Progress in Thermochemical Biomass Conversion
Edited by
A. V. Bridgwater Director of the Bio-Energy Research Group in Chemical Engineering and Applied Chemistry Aston University, Aston Triangle Birmingham, UK
b
Blackwell Science
0 2001 by
Blackwell Science Ltd Editorial Offices: Osney Mead, Oxford OX2 OEL 25 John Street, London WClN 2BS 23 Ainslie Place, Edinburgh EH3 6AJ 350 Main Street, Malden MA 02148 5018, USA 54 University Street, Carlton Victoria 3053, Australia 10, rue Casimir Delavigne 75006 Paris. France Other Editorial Offices: Blackwell Wissenschafts-Verlag GmbH Kurfiirstendamm57 10707 Berlin, Germany
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Contents Volume 1 Preface Acknowledgements
xvii xviii
Gasification Progress in biomass gasification: An overview Maniatis K Steam gasification of wood char and the effect of hydrogen inhibition on the chemical kinetics Barrio M, Gsbel B, Risnes HI Henriksen U,Hustad JE Ssrensen LH
1
32
C02 gasification of birch char and the effect of CO inhibition on the calculation of chemical kinetics Barrio MI Hustad JE
47
C02 reactivity of chars from wheat, spruce and coal Risnes H, Ssrensen LH, Hustad JE
61
Gasification reactivity of charcoal with C02 at elevated conversion levels Struis RPWJ, von Scala C,Stucki S, Prins R Dynamic modelling of char gasification in a fixed-bed &be1 B, Henriksen U,p a l e B, Houbak N Biomass treatment in supercritical water. The way from total oxidation to the gasification Boukis N, Abeln J, Kluth MI Kruse A , Schmieder H, Dinjus E Characterisation method of biomass ash for gasification Moilanen A , Smensen LH, Gustafsson TE, LaatikainenLuntama J, Kurkela E
iii
73
92
109
122
Design of a biomass gasification gas sampling system S&nchezJM, Ruiz E, Cillero E, Otero J , Cabanillas A
137
Characterization of products from biomass tar conversion Morf Ph, Hasler Ph, Hugener M, Nussbaumer Th
150
Guideline for sampling and analysis of tars and particles in biomass producer gases Nee$ JPA, Knoef HAM, Bufinga GJ, Zielke U,Sjostrom K,
162
Brage C, Hasler P, Simell PA, Suomalainen M,Dorrington MA, Greil C Modelling the characteristics of the endothermic reaction potential of tar for flue gas clean-up in advanced thennochemical conversion processes
176
Taralas G Fundamental fluid-dynamic investigations in a scaled cold model for biomass steam-gasification
188
Kehlenbeck R, YatesJG, Di Felice R Stoichiometric water consumption of steam gasification by the FICFB-gasification process
199
Hofbauer H, Rauch R A pilot scale circulating fluidized bed plant for orujillo gasification Garcia-Iban’ez P, Cabanillas A, Garcia- ybarra PL
209
A two-stage pyrolysis/gasification process for herbaceous waste biomass from agriculture
22 1
Henrich E, Dinjus E, Rumpel S, Stahl R Gasification of “liquidized” biomass in supercritical water using partial oxidation Matsumura K Kato A, Sasaki H, Yoshida T
237
Pyrolysis and gasification of black liquors from alkaline pulping of straw in a fixed bed reactor Gea G, Pue‘rtolas R, Murillo MB, Arauzo J
252
iv
Effect of fuel size and process temperature on fuel gas quality from CFB gasification of biomass van der Drifi A, van Doom J
265
Biomass ash - bed material interaction leading to agglomeration in fluidised bed combustion and gasification Visser HJM, Hofmans H, Huijnen H, Kastelein R, Kiel JHA
272
CASST: A new and advanced process for biomass gasification den Uil H
287
Development of a novel, reverse-flow, slagging gasifier van de Beld B, Assink D, Brammer J , Bridgwater AV
298
Study of biomass gasifier-engine systems with integrated drying for combined heat and power Brammer JG, Bridgwater A V
307
Fuel-bound nitrogen conversion: results from gasification of biomass in two different small scale fluidized beds Berg M, Vriesmn P,Heginuz E, Sjostrom K, Espenas, B-G
322
Modelling a CFB biomass gasifier. Part I: Model formulation Corella J , Toledo JM
333
Influence of the reaction atmosphere on gas production and composition in the catalytic conversion of biomass Garcia L, Salvador ML, Arauzo J, Bilbao R
346
The effect of alkali metal on the catalytic gasification of rice straw over nickel catalysts supported on kieselguhr Kim S-B, Lee S-W, Nam S-S, Lee K-W, Choi C-S
358
Hot gas filtration via a novel mobile granular filter Abatzoglou N, Gagnon M , Chomet E
365
Design of a moving bed granular filter for biomass gasification Brown RC, Smeenk J , Wistrom C
379
Redox process for the production of clean hydrogen from biomass Biolluz S, Sturzenegger M , Stucki S
V
388
Hydrogen production from biomass by low temperature catalytic gasification
396
Minowa T, Fang Z Production of substitute natural gas by biomass hydrogasification
405
Mozaflarian M,Zwart RUP A study of carbon dioxide mitigation effect by biomass energy plantation for electricity and methanol
420
YokoyamaS-Y, Tahara K A small-scale stratified downdraft gasifier coupled to a gas engine for combined heat and power production
426
Barrio M,Fossum M , Hustad JE Small scale biomass gasification: Development of a gas cleaning system for power generation
44 1
WalkerM,Jackson G, Peacocke GVC Gas mixing in a pilot scale (500 kWth)air blown circulating fluidised bed biomass gasifier Kersten SRA, Moonen RHW, Prins W, van Swaaij WPM
452
A demonstration project for biomass gasification and power generation in China Wu C, YinX , Zheng S, Huang HI Chen Y
465
Pressurised gasification of biomass and fossil fuels in fluidised bed gasifiers, hot gas cleanup using ceramic filters and pressurised product gas combustion de Jong W, Unal 0, Hoppesteyn P, Andries J, Hein KRG Tri-generation from biomass and residues; options for the co-production of Fischer-Tropsch liquids, electricity, and heat Daey Ouwens C, den Uil H, Boerrigter H Concept for a decentralised combined heat and power generation unit for biomass gasification Romey I, Adomi M, Wartmann J , Herdin G, Beran R, Sjostrom K, Roskn Ch
vi
473
488
499
Biomass power generation: Sugar cane bagasse and trash WaldheimL, Morns M , Leal MRLV Ammonia formation and NOxemissions with various biomass and waste fuels at the Varnamo 18 MWth IGCC plant Goldschmidt B, Padban N, Cannon M , Kelsall G, Neergaard M , Sthhl K, Odenbrand I Tar formation in the 18 MWth biomass IGCC plant in Varnamo and in a 90 kWth pressurised fluidised bed gasifier at Lund University Padban N, Hansson S, Neergaurd M , Sthhl K, Odenbrand I Final report: Varnamo demonstration programme Sthhl K, Neergaurd M , Nieminen J
509
524
536
549
Combustion Examining the thermal behaviour of biomass ash by various analytical techniques Arvelakis S, Gehrmunn H, BeckmunnM , Koukios EG
564
Measuring and modelling the gas residence time distribution in biomass furnaces Biollaz S,Nussbaumer Th, Onder CH
573
A general model for the investigation of packed bed combustion with respect to wood Bruch C,Peters B, Nussbaumer T
585
Reactor network modeling of a biomass dedicated swirling combustor and a fluidized bed gasifier El Asn R, Konnov AA, De Ruyck J
599
New test method to determine efficiency and emissions of slow heat release appliances burning solid fuel Gaegauf CK, Macquat Y
614
Prediction of combustion characteristics for woody biomass fuels - heat output Li J, Giford K, Senelwa K, Hooper RJ, Clemens A, Gong D
630
vii
Parametric modeling study of volatile nitrogen conversion to NO and NzO during biomass combustion LoJler G, Winter F, Hopauer H
641
Profile measurements and modelling studies for optimisation of combustion processes in wood firing systems Unterberger S, Gaegauf CK, Berger R, Hein KRG
656
Visualization and analysis of SEM-EDS data of quartz-bed agglomerates
67 1
Virtanen ME, Tiainen MS, Pudas M I Laitinen RS Model and simulation of heat exchangers and drying silo in a new type of a boiler plant Ytjolu J
678
Testing & modelling the wood-gas turbo stove
693
Reed TB, AnselmE, Kircher K Effect of GR GRANULE used as bed material to reduce agglomeration in BFB combustion of biomass with high alkali metal content Daavitsainen JH, Laitinen RS, Nuutinen LH, Ollila HJ,
705
Tiainen MS, Virtanen ME Elemental gas-particle partitioning in fluidized bed combustion and gasification of a biomass fuel WaffD,Jenkins BM, Turn SQ
713
Evaluation of a novel granular bed filtration system for high temperature applications Risnes HI Smju OK
730
Combustion processes in a biomass fuel bed - experimental results Ronnback MI Axell M I Gustavsson L, Thunman HI Leckner B
743
Combustion performance of New Zealand grown biofuels
758
Senelwa K, Gifford J, Li J, Hooper RJ, Clemens A, Gong D
...
Vlll
Operating parameters for the circulating fluidised bed (CFB) combustion of biomass Smolders K, Honsbein D, Baeyens J Agglomeration and the content of amorphous material in FB combustion. A full-scale boiler test
766
779
Daavitsainen JHA, Nuutinen LH, Tiainen MS,hitinen RS Co-combustion of different waste wood species with lignite in
an industrial steam boiler with a moving stoker firing system Grammelis P, VourliotisP, Kakaras E Biomass and waste-toenergy conversion in the Netherlands by means of (in)direct co-combustion: status, projects and future applications in the Dutch utility sector Konings AJA, Meijer R, Rozendual CM, Ruijgrok WJA,
789
799
de VriesR Whole tree energy power plant
8 12
Ragland KW, Ostlie W ,Berg DA Influence of ash composition on slagging and defluidisation in a biomass fired commercial CFB boiler Tranvik AE, Sanati M , Zethraeus B, Lyberg M
824
Utilisation of bagasse residues in power production
83 1
Beckman D, Solantausta Y Use of thermo-economic analysis based on exergy concepts to evaluate the cost of electricity from sugar cane bagasse in the Brazilian sugar cane sector
843
Coelho ST, Moreira JR, Zylbersztajn D Competitiveness assessment of applications of thermochemical biomass conversion technologies
85 1
Lauer M,Pogoreutz M A comparison of using wood pellets and fast pyrolysis liquid industrially for heat production within Stockholm Ostman A, Lindmun EK, Solantausta Y, Beckman D
ix
867
Development of catalytic wood fired boiler: integrated, deactivation and regeneration of net-based catalysts Berg M, Hargitai T, Brandin J, Berge N
875
Combustion of chlorine-containing biomass: V ~ O S - W O ~ - T ~ O ~ monoliths for C 1-VOCs abatement in the flue exit gas. Part I 887 Corella J, Toledo JM,Gutikrrez M Biomass burner designed to reduce nanoparticle emissions Gaegauf CK, Wieser U,Unterberger S, Hein KRG
896
Emission of UHC and CO from a biomass furnace Griselin N,Bai XS
908
NOx reduction of biomass combustion by optimized combustion chamber design and combustion control Padinger R Polycyclic aromatic hydrocarbons associated to particle size emitted from biomass fluidised bed combustion Saez F, Cabar7as A, Gonzalez A, Escalada R, Martinez JM, Rodriguez-Maroto JJ, Dorronsoro JL, Gdmez F, Saenz D Fuel staging for NOx reduction in automatic wood furnaces Salzmunn R, Nussbaumer Th
918
929
94 1
Estimate of the net C02 reduction by replacing coal and oil with biomass in Japan Dote Y, Ogi T, Yokoyama S
956
The importance of bioenergy and its utilization technologies evaluated by a global energy and land use model Yammoto H, Fujino J, Yamaji K
964
Volume 2
Pyrolysis An overview of fast pyrolysis Bridgwater AV, Czernik S, Piskorz J
977
X
Test bed to turnkey: the introduction of new thermal renewable energy technologies Burdon I
998
Woody and herbaceous biomass feeds - how can we study their composition and their pyrolysis products? Krieger-Brockett B, Rodriguez I
101 1
Biomass selection criteria for pyrolytic conversion processes Ganesh A, Raveendran K
1025
Use of a concentrated radiation for the determination of cellulose thermal decomposition mechanisms 1034 Boutin 0, Lkdk J Modelling and measurements of drying and pyrolysis of large wood particles Larfeldt J, Leckner B, Melaaen MChr Thermal analysis and kinetic modelling of wheat straw pyrolysis Stenseng M , Jensen A, Dam-Johansen K
1046 1061
The potential of multivariate regression in determining formal kinetics of biomass pyrolysis Volker S, Rieckmann Th
1076
A modeling study on cellulose particle pyrolysis under fluidized-bed conditions Yu C, Zhang W, Cen K
1091
Modeling potassium release in biomass pyrolysis Yu C,Zhang W Comparative study on char properties and pyrolysis kinetics of different lignocellulosic wastes Bonelli PR, Della Rocca PA, Cerrella GE, Cukiemn AL The pyrolysis kinetics of a single wood particle Davidsson KO, Pettersson JBC, Bellais M, Lilieduhl T, &'jostromK
xi
1107
1116 1129
Dynamics and products of wood pyrolysis Di Blasi C,Branca C, Santoro A, Hernandez EG, Bennudez RAP The mathematical modeling of biomass pyrolysis in a fixed bed with experimental verification Chen G, Andnes J, h u n g YC
1143
1158
Origin and nature of paramagnetic moieties in pyrolysis oils Dizhbite T, Dobele G, Mironova N, Telysheva G, Meier D, Faix 0
1171
New prospects for biocarbons
1179
Antal MJ Jr, Dai X , Shimizu B, Tam MS, G r d i M Issues in value-added products from biomass
1186
Elliott DC Combined chemicals and energy production from biomass pyrolysis Himrnelblau A, Beck RW Sibunit supported catalysts for hydrogenolysis of a C-0 bond in ‘bio-crude-oil’ components
1197
1207
AksenovDG, Startsev AN, Kuznetsov BN Multi-parameter assessment of sunflower husk-sawdust layer hydraulic resistance
1213
GubynskyyM, Shishko Y, UsenkoA, VvedenskuyaT Organic composition of liquidized model kitchen garbage Inoue S, Minowa T, Sawayama S, Ogi T
1219
The volatility of tars from pyrolysis of biomass materials Oja V, Hajaligol MR
1226
Release of chlorine from biomass and model compounds at pyrolysis and gasification conditions
Stromberg B, Zntl F
xii
1234
The char residues from pyroysis of biomass - some physical properties of importance Suuberg EM, Aama I, Milosavljevic I
1246
Biomass fast pyrolysis in an air-blown circulating fluidized bed reactor Boukis I, Gyftpoulou ME, Papamichael I
1259
Rotating cone bio-oil production and applications Wagenaur BM, VenderboschRH, Carrasco J , Strenziok R, van der Aa BJ
1268
CFD for the modelling of entrainment in fluidised bed fast pyrolysis of biomass Gerhuuser H, Generalis SC, Hague RA, Bridgwater AV
1281
Modelling, scale-up and demonstration of a vacuum pyrolysis reactor Yang J, Blanchette D, de Caum'a B, Roy C
1296
Thermal efficiency of the HTU@ process for biomass liquefaction Goudriaan F, van de Beld B, Boerejljn FR, Bos GM, Naber JE, van der Wal S, ZeevalkinkJA
1312
Thermochemical treatment of radiata pine using hot compressed water Ogi T, Inoue S, Yazaki Y
1326
Chemical conversion of biomass resources to useful chemicals and fuels by supercritical water treatment Saka S, Konishi R
1338
Co-pyrolysis under vacuum of bagasse and petroleum residue Chaala A, Garcia M , Roy C
1349
Preliminary results on wood waste pyrolysis Dudouit C,Schenkel Y
1364
Fast pyrolysis of industrial biomass waste Gerdes Ch, Meier D, Kaminsky W
1374
...
Xlll
Co-pyrolysis of wood biomass and plastic wastes of different origin under the pressure of argon and hydrogen Kuznetsov BN, Sharypov VI, Beregovtsova NG, Mann N, Weber JV Fate of arsenic after fast pyrolysis of chromium-copper-arsenate (CCA) treated wood Hata T, Meier D, Kajimto T, Kikuchi H, Immura Y
1388
1396
Fast pyrolysis of impregnated waste wood - the fate of hazardous components 1405 Meier D, Ollesch T, Faix 0 Low-temperature pyrolysis as a possible technique for the disposal of CCA treated wood waste: metal behaviour 1417 Helsen L, Van den Bulck E Pyrolysis of biomass as pre-treatment for use as reburn fuel in coal-fired boilers Storm C, Unterberger S, Hein KRG Combustion of bio-oil in a gas turbine Strenziok R, Hansen U,Kiinstner H Stirling engine with FLOX" burner fuelled with fast pyrolysis liquid Bandi A, Baumgart F Pyrolysis oil combustion tests in an industrial boiler Oasmaa A, Kyto M,Sipila K
1433 1452
1459 1468
Transport, handling and storage of biomass derived fast pyrolysis liquid Peacocke GVC, Bridgwater A V
1482
Levoglucosenone - a product of catalytic fast pyrolysis of cellulose Dobeie G, Rossinskuja G,Telysheva G,Meier D, Radtke S, Faix 0
1500
xiv
Microporous sorbents produced by pyrolysis and gasification of hydrolytic lignin Plaksin GV, Baklanova ON, Duplyakin VK,Drozdov VA
1509
The formation of petrodiesel by the pyrolysis of fatty acid methyl esters over activiated alumina 1517 Boocock DGB, Konar SK, Glaser G Bio-crude-oil/Diesel oil emulsification: main achievements of the emulsification process and preliminary results of tests on diesel engine Baglioni P, Chiararnonti D, Bonini M , Soldaini I, Tondi G
1525
Production of diesel fuel additives from the rosin acid fraction of crude tall oil Coll R, Udas S, Jacoby WA
1540
Preliminary study of fungicide and sorption effects of fast pyrolysis liquids used as wood preservative Meier D, Andersons B, Irbe I, Chirkova J, Faix 0
1550
Fractional vacuum pyrolysis of biomass for high yields of phenolic compounds Pakdel H, Murwanashyaka JN, Roy C
1564
Production of hydrogen from biomass-derived liquids Czemik S, French R, Feik C, Chomet E
1577
Co-firing of bio-oil with simultaneous SOxand NOx reduction VenderboschRH, Wagenaar BM, Gansekoele E, Sotirchos S, Moss HDT
1586
Improving charcoal kiln performance - do fundamental studies have a role? Connor MA
1603
Effect of four physical characteristics of wood on mass and energy flows from slow pyrolysis in retorts Schenkel Y
1618
xv
Influence of temperature, residence time and heating rate on pyrolytic carbon deposition in beech wood chars Schenkel Y Catalysed carbonisation of fine woodworking industry residues Zandersons J, Zhurinsh A, TardenakaA, Spince B
1633
1642
Thermal desorption technology: low temperature carbonisation of the biomass for manufacturing of activated carbon Somkus GE
1651
Workshops reports
1661
Author index
(see also at the end of Volume 1) 1683
Subject index
(see also at the end of Volume 1) 1687
xvi
Preface There has been considerable progress in the science and technology of thermo-chemical biomass conversion since the previous conference on Developments in Thermochemical Biomass Conversion in Banff, Canada, in 1996. This fifth conference again covers all aspects of thermal biomass conversion systems from bdamental research through applied research and development to commercial applications to reflect the progress made in the last four years. The programme was divided into three major thermal conversion technologies: gasification, combustion and pyrolysis, and this division is reflected in the structure of these proceedings. Each main area of the conference was preceded by a state-of-the-art review to provide a focus for the ensuing presentations and an authoritative reference. One of the major features of this conference series is the high quality of the presentations and papers, which is achieved by subjecting all contributions to a full peer review process. An important aim was to provide the widest opportunities to exchange
ideas, discuss problems with fellow researchers, and to hear about the latest research and development. This was achieved in two ways: all posters were formally presented in short presentations to provide all authors and all delegates with the opportunity to meet each other face to face and benefit from the ensuing intimate interaction of a small and interested group; and the workshop programme further encouraged this interaction in those areas of interest selected by participants. The resultant report at the end of the book provides a summary of the outcomes from the workshops. The research community will continue to provide the lead in developing new science and technology and in stimulating the development of new ideas. The benefits from the interactions at this conference will eventually translate into new andor better products and processes, which will not only provide the justification for continued research and development, but will also more rapidly translate into commercial processes and products and help to deliver the promised benefits of the bio-energy sector. Tony Bridgwater March 2001
xvii
Acknowledgements The following organisations provided considerable financial support to the conference, which was much appreciated by the organiser and delegates: 0 IEA Bioenergy through the Executive Committee, and the Combustion, Pyrolysis and Techno-Economic Assessment Tasks; 0 Austrian Federal Ministry of Transport, Innovation and Technology; 0 Austrian Federal Ministry of Agriculture, Forestry, Environment and Water Management; 0 Department of Trade and Industry, UK; Natural Resources Canada; 0 VTT - Technical Research Centre of Finland The scientific committee listed below provided whole-hearted support throughout the preparation and running of the conference. Their encouragement and efforts in publicising the meeting, refereeing papers, providing constructive feedback on the programme, organising workshops and chairing sessions was invaluable: Larry Baxter, USA Michaei Antal, USA Ton Beenackers, Netherlands Dave Boocock. Canada Esteban Chornet, Canada Robert Brown, USA Jose Corella, Spain Mike Connor, Australia Doug Elliott, USA Columba Di Blasi, Italy Hermann Hofbauer, Austria Johan Hustad, Norway Walter Kaminsky, Germany Barbara Krieger-Brockett, USA Kyriakos Maniatis, Belgium Robert Manurung, Indonesia Dietrich Meier, Germany Rainer Marutzky, Germany Thomas Nussbaumer, Switzerland Tom Milne, USA Jan Piskorz, Canada Ralph Overend, USA Ann Segerborg-Fick, EC, Belgium Christian Roy, Canada Krister Sjostrom, Finland Kai Sipila, Finland Josef Spitzer, Austria Yrjo Solantausta, Finland Chuangzhi Wu, China Erik Suuberg, USA Shin-ya Yokoyama, Japan Particular thanks are due to Nina Ahrendt and Claire Humphreys who provided all the Conference Administration throughout the preparation and running of the conference. Their contribution was invaluable and much appreciated.
xviii
Progress in Biomass Gasification: An Overview K. Maniatis Directorate General for Energy & Transport, European Commission, Rue de la Loi 200, 1049 Brussels, Belgium
ABSTRACT Gasification is an energy process producing a gas that can substitute fossil fuels in high efficiency power generation, heat andor CHP applications, and can be used for the production of liquid fuels and chemicals via synthesis gas. Gasification technology consists of several unit operations, the most critical of which is gas cleaning and conditioning for utilisation in power production engines. Numerous types of gasifiers have been developed and tested and many industrial applications can use the technology. Significant progress has been achieved over the last five years and some applications are on the threshold of becoming commercial. However, for most of the applications the efficient and economic removal of tar still presents the main technical barrier to be overcome. This overview reports on the progress achieved over the past five years in thermochemical gasification of biomass and waste recovered fuels. The status of all major projects is reviewed while new trends are briefly presented. The paper concludes with recommendations for future R&D needs and demonstration requirements while attempting to present a strategy for the commercialisation of gasification technologies.
INTRODUCTION Biomass is considered the renewable energy source with the highest potential to contribute to the energy needs of modem society for both the developed and developing economies world-wide (1,2). Energy from biomass based on short rotation forestry and other energy crops can contribute significantly towards the objectives of the Kyoto Agreement in reducing the green house gases emissions and to the problems related to climate change (3). Biomass fuels and residues can be converted to energy via thermochemical and biological processes. Biomass gasification has attracted the highest interest amongst the thermochemical conversion technologies as it offers higher efficiencies in relation to combustion while flash pyrolysis is still in the development stage. However, although gasification technologies have recently been successfully demonstrated at large scale and several demonstration projects are under implementation ( 4 3 , they are still relative expensive in comparison to fossil based energy and, therefore, face economic and other non-technical barriers when trying to penetrate the energy markets (6,7,8). Their penetration into the energy markets can 1
only be achieved at present via economic. development through biomass systems integration. Thus the ,innovation in practically all demonstration projects under implementation lies not only on the technical aspects of the various processes but also in the integration of the gasification technologies in existing or newly developed systems where it can be demonstrated that the overall system offers better prospects for economic development (9). The overview starts with the present status of the various gasification technologies and after a brief introduction to their market prospects the most important projects are briefly presented and discussed based on their market segments. Developments in the various fields are also discussed as well as the R&D needs for an accelerated penetration of gasification technologies in the energy market.
STA TUS OF GASIFICATION TECHNOLOGIES An extensive review of gasifier manufacturers in Europe, USA and Canada (10) identified 5 0 manufacturers offering ‘commercial’ gasification plants Erom which: (1)
75% of the designs were downdraft type,
(2) (3)
20% of the designs were fluidized bed systems, 2.5% of the designs were updraft type, and, 2.5% were of various other designs.
(4)
However, there was very little information on cost aspects, emissions, efficiencies, turn-down ratios and actual operating hours experience. Above all, no single manufacturer was ready to give full guarantee for technical performance of their gasification technology. This indicates that the actual operating experience is limited and there is little confidence on the technology, which is due to the general poor performance of the various prototypes. Figure I presents a tentative status for gasification technologies in view of their market attractiveness for power generation and the present strength of the various gasification technologies. TECHNOLOGY STRENGTH Strong
Average
Weak
Atm.CFB Atm.BFB
High
0Press. CFB
I Press. BFB MARKET Medium ATTRACTIVENESS
Updraft
0 Downdraft
Low
0Cyclonic
0
Entrained Bed
Figure 1 Technology development and strategic planning for power Atmospheric Circulating Fluidized Bed Gasifiers (ACFBG) have proven very reliable with a variety of feedstocks and are relative easy to scale up from few MWth
2
up to 100 MWth. Even for capacities above 100 MWth, there is confidence that the industry would be able to provide reliable operating gasifiers. It appears to be the preferred system for large scale applications and it is used by most of the industrial companies such as TPS (1 l), FOSTER WHEELER (12), BATTELLE (13), LURGI (14) and AUSTRIAN ENERGY (15). Therefore ACFBG have h g h market attractiveness and are technically well proven. Atmospheric Bubbling Fluidized Bed Gasifiers (ABFBG) have proven reliable with a variety of feedstocks at pilot scale and commercial applications in the small to medium scale; up to about 25 MWth. They are limited in their capacity size range as they have not been scaled up significantly and the gasifier diameter is significantly larger than that of ACFBG for the same feedstock capacity. On the other hand ABFBG are more economic for small to medium range capacities, thus their market attractiveness is relative high as well as their technology strength. Companies promoting ABFBG are CARBONA (16 ) and DINAMEC (17). Pressurised fluidized bed systems either circulating (PCFBG) or bubbling (PBFBG) are considered of medium market attractiveness due to the more complex operation of the installation and to the additional costs related to the construction of all pressurised vessels. On the other hand, pressurized fluidized bed systems have the advantage in integrated combined cycle applications as the need to compress the fuel gas prior its utilisation in the combustion chamber of the gas turbine is avoided. Pressurised systems have been proposed mainly by CARBONA (16 ) and FOSTER WHEELER (18) with the successful application of SYDKRAFT’s Varnamo IGCC plant in Sweden. Atmospheric Downdraft Gasifiers (ADG) are attractive for small scale applications (4.5 MWth) as there is a very big market not only in developed but developing economies too (19). However, the problem of efficient tar removal is still a major problem to be addressed and there is a need for more automated operation especially for small scale industrial applications. Nevertheless, recent progress in catalytic conversion of tar (see below) gives credible options and ADG can therefore be considered of average technical strength. Atmospheric Updraft Gasifiers (AUG) have practically no market attractiveness for power applications due to the high concentration of tar in the fuel gas and the subsequent problems in gas cleaning. Also the technology is considered weak for the same reasons. There is no company proposing AUG for power at present. Atmospheric Cyelonic Gasifiers (ACG) have only recently been tested for biomass feedstocks and although they have medium market attractiveness due to their simplicity, they are still unproven. No well known company is promoting ACG. Finally, Atmospheric Entrained Bed Gasifiers (AEBG) are still at the very early stage of development and since they require feedstock of very small particle size, their market attractiveness is very low. No company is presently developing pressurised systems for downdraft, updraft, cyclonic or entrained bed gasifiers for biomass feedstocks and it is difficult to imagine that such a technology could ever be developed into a commercial product due to the inherent problems of scale, tar removal and cost. In conclusion, for large scale applications the preferred and most reliable system is the circulating fluidized bed gasifier whle for the small scale applications the downdraft gasifiers are the most extensively studied. Bubbling fluidized bed gasifiers can be competitive in medium scale applications. Large scale fluidized bed systems have become commercial due to the successful co-firing projects (see below) while moving bed gasifiers are still trying to achieve this.
3
For heat applications there is no need to eliminate the tar from the fuel gas and thus any reliable gasifier system can be used successfully. However, although heat applications are relative easy, there are very few examples in the market. The most successful has been the BIONEER updraft gasifier (20), which has been used successfully in ten commercial applications in Finland. This gasification technology was originally commercialised for lime-kiln'applications with peat as main fuel and was later applied to co-utilisation of locally available residues and wastes in existing boilers.
FEEDSTOCK TECHNOLOGY RELIABILITY One of the most important barriers to an accelerated penetration of all biomass conversion technologies is that of adequate resource supply. Figure 2 depicts the technology reliability of using the most important feedstocks in gasification applications.
0 Low MARKET POTENTIAL
RDF Grasses
0 0
Woody Biomass
0
Straw
S RF
High
0
Sludge
High
Low
OVERALL TECHNOLOGY RELIABILITY
Figure 2 Status of feedstock technology reliability and market potential Clean biomass feedstocks are becoming scarce and there is hardly any reliable supply. In some countries like Germany, all industrial wood waste and other wood residues are consumed completely and there is no other clean biomass available to increase the contribution of bioenergy. Thus, the industry has been obliged to look into relative difficult fuels and fuels with little practical industrial experience in order to create new market opportunities. Waste recovered fuels present the advantage that they often have a negative cost associated with their disposal, which can significantly decrease the operating costs of a plant. In addition, since the last decade there has been a significant interest in energy crops and especially short rotation forestry (SRF) as a means to increase the production of biomass fuels while simultaneously creating new jobs for the farming community. SRF operations can also contribute significantly towards sustainability and meeting the Kyoto's obligations. Woody biomass has the hghest reliability in feeding into a gasifier and most problems related to bed sintering in fluidized bed gasifiers or slag formation on heat exchange surfaces are relative well understood and the industry has sufficient confidence to use effectively most types of woody biomass (21). The industry has also attained a high degree of reliability for the pretreatment operations such as 4
drying, size reduction and storage. However, the market potential of woody biomass is limited as most of the locally available feedstocks are already consumed in various industrial or district heat applications. Short Rotation Forestry has relative good potential to be used in non arable land and provides a sustainable approach to energy (22), however, since the land has to be blocked for about 15-20 years, farmers in the EU are reluctant to implement SRF schemes. The only exceptions are Sweden, where there is a long tradition for SRF mostly for pulp and paper and the UK where recently successful schemes have been introduced to the farming community (23). The USA (24) has also an ambitious programme for the development of SRF while Canada also has carried out significant work and is examining various SRF implementation schemes. Brazil has successhlly established eucalyptus plantations (23). On the other hand, very few tests have been carried out with SRF feedstocks and the industry is a somewhat uncertain about the properties of SRF fuels. A sensitive area is that of heavy metals some of which are easily up taken by the plants (e.g. cadmium). Grasses (25) have attracted interest recently since they can be cultivated on various places, even on the sides of highways, however, their market potential is still uncertain as there are no dedicated plantations yet and there is relative little experience with such feedstocks. Technically grasses present problems in all pretreatment operations such as size reduction, storage, drying and even their relative fast biodegradability which can result in significant weight loss unless dried and properly stored. Their low bulk density results in solids flow problems and can create local hot spots in the gasifier. Straw has a relative low market potential for gasification applications since successful combustion technologies have been developed. There is little experience with straw gasification and severe problems of ash sintering and bed agglomeration are known to exist in fluidized bed gasifiers. Due to the low bulk density it is not possible to use straw in moving bed gasifiers unless the straw has been palletised; an expensive operation. However successful operation of the Varnamo plant was achieved with 100% straw feeding ( 18). Refuse Derived Fuel has significant potential for gasification applications since gasification does not have such a negative public image as incineration and there is sufficient experience by TPS (26) and FOSTER WHEELER (12). However, the feeding systems for fluff RDF need to be developed fiuther to ensure reliable operation and more experimental results at large scale applications are needed to prove efficient operation. Finally sludge can also be utilised in gasification applications and although there exists little experience, it is expected that the application with sludge may increase in the future. Technical reliability still has to be demonstrated. Recently two feedstock databases have been established where a significant amount of information is provided for a variety of biomass feedstocks (27, 28). In the databases, the basic physico-chemical properties of biomass fuels can be found whch will provide basic information to gasifier developers as well as gasifier users on the quality and suitability of the various fuels for the gasification technologies they either develop or use.
BIOMASS & RECOVERED FUELS STANDARDISATION The guaranteed supply of biofuels is an important element for the promotion of bioenergy in general and gasification technologies in particular. This requires the
5
creation of a biofuels market, which necessitates the development of standards to govern any transaction between producers and users of biofuels. The availability of standards for biofuels would provide guidance to the farmers, foresters and producers of recovered fuels of the types and quality of the fuels the market requires as well as a guarantee to the users of the quality of the fuels they procure from the market. The European Commission took the initiative to instigate an action in 1998 with the industry and several EU National Organisations on this issue, which resulted in a mandate to the European Centre for Normalisation (CEN) for the elaboration of standards for Solid Biofuels (CEN/TC 335). For Solid Recovered Fuels, originating from waste streams, the European Commission gave only a programming mandate to CEN (CEN/BT/Task Force 118), which may eventually result into a full mandate subject to the contents of the various classes of fuels. Both mandates are to be supported by quality assurance systems to ensure the quality of the fuels entering the future biofuels market. This action was evolved to an IEA Bioenergy Task, “Task 28 Standardisation of Solid Biomass Fuels” with the inclusion of US Department of Energy in the standardisation activities (29). In particular for gasification technologies, which require well calibrated and defined feedstocks in terms moisture, size, ash and inerts, the standards will facilitate the production of dedicated fuels for gasifiers.
TAR REMOVAL The efficient removal of tar still remains the main techmcal barrier for the successful commercialisation of biomass gasification technologies and unless this barrier will be properly addressed biomass gasification applications for power, with the exception of pressurised IGCC, will never materialise. There are several groups which have been working extensively on tar, however the most prominent teams are those (in alphabetical order) of the University of Madrid (30, 31), The Royal Institute of Technology in Stockholm (32,33) and VTT (34,35).
TAR REMOVAL TECHNOLOGIES The main attempts to eliminate tar concentrate on three approaches: scrubbing, catalytic reforming followed by scrubbing and hot gas clean up. In the later case the producer gas is kept above 400 “Cin order to avoid tar condensation and the hot fuel gas is burned in the combustion chamber of a gas turbine. However, this approach applies only to pressurised gasification IGCC systems and has been successfully demonstrated at the Varnamo plant (18) while recently the quality of the tar produced by the FOSTER WHEELER gasifier has been reported (36). This approach has been proven successful as there were no problems due to tar (either in the filters or in the gas turbine) during the operation of the Vamamo plant for more than 3600 h on IGCC operation. The simple scrubbing approach has failed repeatedly to prove long term operational reliability and in addition it creates a serious environmental problem because of the large quantities of condensate produced. Between wet and catalytic cleaning methods, the latter is preferred because it actually destroys/modifies the tars instead of transferring them to a liquid phase, which needs further and expensive waste water treatment. Among the possible catalytic cleaning methods, two are being adopted by most of the institutions and companies working on biomass gasification:
6
they are either based on the use of calcined dolomites (or related materials) or of steam reforming (nickel-based) catalysts located downstream the gasifier. Calcined dolomites have proved their usefulness for tar removal but they have two operational problems. First they have low mechanical strength, which result into significant erosion and thus catalyst consumption, with the associated increase of the particulates content in the fuel gas. Second because of their low catalytic activity, it is very difficult to reduce the tar contents in the exit fuel gas below 0.5-1.O giNm3. Catalytic tar elimination over nickel-based catalysts mainly proceeds by steam and dry (COz) reforming reactions, although there can be simultaneous thermal reactions of cracking and, perhaps, of hydrocracking. Therefore, the steam and C 0 2 contents in the flue gas have an important role in the overall tar elimination. Tar conversion (elimination) depends on the properties of the catalyst used, on the spacetime (or space-velocity), bed temperature, H2O to carbon to be reformed ratio and on the operation variables of the upstream gasifier such as equivalence ratio and temperatures in the bed and in the freeboard. The combination of catalytic reforming followed by scrubbing offers the only remaining hope for the gasification industry to effectively address the tar problem. Three main configurations have been proposed and are under development: 1. reverse flow catalytic bed with dolomite, 2. second, fluidized bed with dolomite, and, 3. catalytic bed with monolith based catalyst.
Reverseflow catalytic bed with dolomite The reverse flow catalytic bed has been proposed by BTG (37) and Wellman (38). The principle is based on transferring heat by the partial combustion of the he1 gas to maintain the temperature of the catalyst bed to about 900 "C from the gas to the catalyst bed and vice-versa. The hot reaction front moves along the catalyst bed and when the reaction front reaches the end of the catalyst bed the flow of the gas changes (Fig. 3).
Figure 3 Schematic of the catalytic fixed bed reverse flow tar removal system This technology has been demonstrated at pilot scale by both organisations, however, it still has to be operated in demonstration scale applications to prove its reliability and economics. BTG will use the technology in a demonstration project (39), which is schedule to start operation by 2002. 7
Secondary fluidized bed with dolomite This approach has been proposed by TPS (1 1,40), (Fig. 4) and has been demonstrated extensively at pilot scale with a variety of feedstocks. Full scale demonstration is expected to be achieved soon when the ARBRE plant (41) is commissioned in Spring 2001. An important element in the ARBRE demonstration project would be the operational cost for the calcined dolomite catalyst as well as the associated waste disposal costs. The successful demonstration of the ARBRE project would provide reliability for large scale IGCC and will also be the ‘springboard’ for other large scale projects in preparation such as the second generation of ARBRE (35-40 MWe) in the UK, (41) and the Brazilian project (30-35 MWe), (1 1). Dolomite Gasifim
Cradter
Gzg Cooling
Gar Purnmtion
Fuel GES
Figure 4 Schematic of second fluidized bed and subsequent gas cleaning Catalytic bed with monolith based catalyst The monolith catalysts are the least tested in pilot scale, however they have the advantage that they offer good mechanical strength and have high catalytic activity. On the other hand their cost is considerably higher and they are more prone to poisoning and deactivation than dolomite and related catalysts. Because of their cost, the most important operational variable is the life of the catalyst. There are three causes that can deactivate the catalyst, coke (formed from tar), sulphur poisoning, and particulates. Deactivation by coke is not important if the tar content in fuel gas is below 2 g tar/Nm3 and this can be achieved with a good operation of the upstream gasifier (with in-bed dolomite and high temperatures in the gasifier freeboard, for instance). Deactivation by sulphur remains to be verified in very-long term tests although the sulphur content in biomass feedstocks is not high. Finally deactivation by particulates can be avoided if monoliths are used instead of rings and the particulates content in the fuel gas is carefully controlled. The main catalyst manufacturer and provider is BASF AG of Germany, while the two main research teams whch have extensively studied Nickel-based monoliths are the University of Madrid and VTT. There are two recent projects, which plan to use t h s approach in eliminating tar (42, 43). Both are based on bubbling fluidized bed gasifiers (Fig. 5) and are
8
addressing the small to medium scale for power applications (1-5 MWe). In both cases the fuel gas is burned in a gas engine to generate electricity while in one of the projects (42) the waste heat is recovered in a district heat application. The success of these two projects will be critical for this market segment as other tar removal options are uneconomic or unreliable. Gas
FluidBcd G a*r
CycLr
1
i Air
Clean water to vtoragdsawn
Figure 5 Schematic of catalytic reactor arrangement with monolith Ni-based catalyst
TAR PROTOCOLS A number of different sampling and analysis methods have been developed by manufacturers and various institutes workmg in this field to determine the level of particulates and tar in the gas exiting the gas cleaning system of a gasification installation. This diversity of methods makes the comparison of operating data from different sources very difficult and represent a significant barrier to the further development and commercialisation of the technology. The members of the IEA Bioenergy Gasification Task, the European Commission and the US DOE have been aware of this barrier for some time. In order to address it they called a joint meeting in the Spring of 1998, where it was decided to draft two sampling and analysis protocols - one for small scale, fixed bed, engine based systems and the other for larger utility scale plants (44). These protocols were to describe the best available procedure for each scale of operation and should reflect as far as possible the collected experience and expertise of the international gasification community. It was intended that the procedures should be used as reference methods and would eventually be submitted as the bases of European and US standard methods. More specifically the objectives were to describe a procedure and set of apparatus that would allow a skilled technician to: 0
Determine the concentration of particulate material in the fuel gas stream gravimetrically; Determine the concentration of higher hydrocarbons or tars gravimetrically.
9
The first results (45, 46) were reported in a dedicated workshop during the 10" European Conference on Biomass for Energy & Industry, Wurzburg, 1998 and were published in a special issue of Biomass & Bioenergy, volume 38 (44). This work prompted a significant amount of high quality dedicated publications from several groups working on this problem (47-50). It also concentrated the European scientific resources in a collaborative effort to address the tar & particulate measurement during parallel tar measurements sponsored by the Danish Energy Agency, NOVEM, VTT & the Swiss Federal Office of Energy (51). In addition, an international concerted action was sponsored by the European Commission to continue the work initiated by the Gasification Task of IEA Bioenergy and complete the tar protocols (52). The first conclusions of the work was that the protocols should be integrated into one for sampling and analysis of tar from all biomass gasifiers under all relevant conditions (0-900 "C; 0.9-60 bar) and concentrations in the range of lmg/Nm3 to 100 g/Nm3. Compared to the Wurzburg Protocols, another solvent will be used because dichloromethane is not considered to be suitable for reasons of safety and health. APPLICATIONS
The various gasification applications for power and or heat are shown in Fig. 6 in terms of their market potential and overall technology reliability. Each of these applications will be discussed in the subsequent sections and the most advanced plants in each application will be presented in terms of their status and future prospects. It is of course beyond the scope of this overview to present all known activities, however, the most significant of these will be discussed as a means of presenting their achievements for the benefit of the other projects, which are still in the development stage. All demonstration projects had to overcome numerous technical and non-technical barriers as this is an emerging technology, however, many of these problems are common to all projects in the same application field and thus the projects still in the development face could learn from the experiences of the others.
Low MARKET POTENTIAL
High
High
Low
OVERALL TECHNOLOGY RELIABILITY
Figure 6 Status of applications for market potential and technology reliability
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FIRING IN BOILERS OR HEAT APPLICATIONS
Firing the raw gas in boilers or heat applications such as kilns after removal of dust and particulates is the simplest application since the gas is kept hot and the tar problem is avoided. However, surprising enough, there are very few known successful applications, which have been operating in a commercial environment. This market is one where all types of gasifiers can compete and more concerted efforts have to be undertaken by the gasification industry to increase the number of successful cases. Between 1985 and 1986, when fuel oil prices were high, eight commercial BIONEER plants, with capacities ranging from 4 to 5 MWth, were commissioned: five in Finland and three in Sweden. Four plants are operated with wood or wood and peat mixtures while the rest are operated with peat only. Most of the gasifiers are still in operation at small district heating plants to provide circulating hot water. Ahlstrom Corporation bought the BIONEER company originally owned by YIT Corporation. After Foster Wheeler acquired Ahlstrom, in 1996 a 6.4 MWth plant was installed at Ilomantsi, in eastern Finland. The estimated investment cost for district heating applications is about 350 kECU/MWth, operating cost is about 17 ECUMWh, and heat generation cost is about 20 ECUkWh (53). The first commercial Ahlstrom Pyroflow CFB gasifier was commissioned in 1983 at the present Wisa Forest Pulp and Paper Mill in Pietarsaari, Finland. The fuel for the 35 MWth (about 150 TPD of biomass) gasifier is primarily bark and saw dust. The biomass is fed from the side into the circulating sand of an air-blown CFB gasifier maintained at about 900 C. The hot fuel gas at 700 C, is fed directly to a lime kiln. The fuel gas replaces about 85% of the fuel oil for the lime kiln. Between 1985 and 1986, three more gasifiers, two in Sweden (25 MWth at Norrsundet Bruks, and 27 MWth at ASSI, Karlsborg) and one in Portugal ( 15 MWth at Portucel, Rodao Mill), were built and commissioned for firing lime kilns (53). Corenso United Oy Ltd. is commissioning a new gasification plant for energy production and aluminium recovery at its core-board mill in Varkaus. The plant will enable the complete exploitation of used packages containing wood fibre, plastic, and aluminium. It will be the first plant in the world able to recycle the aluminium in used liquid packaging to create a raw material for foil for its original purpose, while simultaneously exploiting the plastic contained in the packages to produce energy. The fibre material in multi-layer packages (Tetra-Pack) can be recycled in core-board while the aluminium being recycled as raw material for foil. The remaining plastic will be gasified to generate 40 MWth, with an estimated annual total energy production of about 165 GWh. The estimated cost of the new plant being built is around EUR 17 million, with completion scheduled for the autumn of 2000. The investment includes the gasifier, an aluminium recovery unit and a new boiler designed specially for gasification gas (53). The first commercial TPS CFB gasification process was built for Refuse Derived Fuel (RDF) gasification at Greve in Chianti and started commercial operation in 1993. RDF pellets, up to 150 mm long, are fed into the lower section of two 15 MWth capacity CFB gasifiers, at a rate of about 3 t/h. The air blown TPS gasifier operates at a temperature of about 875 "C.The fuel gas has a heating value of 8 MJ/Nm3 (54). The raw gas from one of the gasifiers passes through two stages of solids separation before being fed to a furnacehoiler to generate steam for producing 2.3 MWe in a condensing steam turbine. The overall power generation efficiency is about 19 to 20%. The gas produced in second gasifier is supplied to the neighbouring cement
11
factory for a direct combustion in the cement luln. The gas leaves the cyclone at a temperature of about 850°C and is sent to a oil-filled heat exchanger to be cooled to a temperature of about 450"C, before it is sent to the cement factory, (Fig. 7) (55). At present the owner of the plant is installing a modem second boiler and fuel gas cleaning equipment in order to generate electricity from the second gasifier as well to a total capacity of 6.7 MWe. The plant has been operated intermittently due to difficulty in obtaining continuous supply of RDF pellets.
.
I
Cement Factory
Figure 7 The Greve in Cluanti process flowsheet The BattelleRERCO project in the US was built at the McNeil power plant in Burlington, Vermont. The 200 ton per day project employs the low pressure Battelle gasification process that consists of two reactors: (a) a gasification reactor in which the biomass is converted into a MCV gas and residual char at a temperature of 70085OoC, and (b) a combustion reactor that burns the residual char to provide heat for gasification. Heat transfer between reactors is accomplished by circulating sand between the gasifier and combustor (Fig.8). Since the gasification reactions are supported by indirect heating, the primary fuel gas is a medium calorific value fuel gas. The estimated HHV of this fuel gas is 17.75 MJ/Nm3. Full plant operation was achieved in mid 2000 using wood chips. It is envisaged that in subsequent phases, the fuel gas will be cooled for heat recovery, scrubbed, and recompressed prior to energy conversion and recovery in a 15MWe gas turbine system (13, 56).
IGCC PROJECTS Several project have been initiated for IGCC applications over the last decade, however, only two have been implemented, the SYDKRAFT plant at Varnamo based on FOSTER WHEELER technology (18) and the ARBRE plant based on TPS technology (41). The Vermont project based on BattelleEERCO may be upgraded to an IGCC plant in the medium to long term, however, there are no concrete plans at
12
present. The Energy Farm project in Pisa with LURGI technology (57) and the Brazilian project with TPS technology (58) still face implementation problems and their future is uncertain. T h s indicates that such large scale projects still face barriers which are mainly related to high installation cost and high technical risks due to the emerging technology status of gasification. However, the successful operation of the ARBRE project, the first commercial IGCC, will provide reliability for the technology and a basis for scaling up with confidence so that the second generation ARBRE could be built with reduced costs indicated by learning effects. In addition there is a small scale IGCC plant based on an hot gas indirect fired gas turbine which is described below in the section concerning hot gas operations of gas turbines.
Figure 8 Schematic of the Battelle gasification technology
The Varnamo Plant The plant in Varnamo (Fig. 9) produces about 6 MWe electricity to the grid as well as 9MWth to the district heating system of the city of Vamamo, from a total fuel input equivalent to 18 MW (18, 59). The accumulated operating experience amounts to about 8500 hours of gasification with more than 3600 hours of gas turbine operation on gas. A successful test programme was completed in 2000 addressing fuel flexibility and NOx emission problems. Fuels including wood, bark, forest residues, willow grown on energy crops, straw and RDF have been used without any major operating problem. However, some problems occurred in the hot gas filtration system, where some ceramic filter candles broke. The reason for the cracking was found to be mechanical fatigue due to micro craclung in the filter elements and since 1999 sintered metal filters are installed. No problem was experienced with tlus type of filter.
13
Figure 9 The process flow diagram of the Varnamo plant
In addition to the demonstration programme, development work aiming to substantially reduce the NOx emissions from gasification plants originating from fbelbound nitrogen compounds was carried out under the coordination of VTT Energy. The method is based on controlled and selective oxidation of fured nitrogen species, primarily ammonia and hydrogen cyanide, of the gasification product gas to N2.The central part of this research is the development of a new SCO (Selective Catalytic Oxidation) technology. Tests in reducing the bed material feed were also made in order to check the possibilities of reducing the operating costs for the IGCCtechnology. It has also been possible to make a drastic reduction of the bed material cost as the tests have proved that a significant amount of the bottom ash from the gasifier can be fed back into the process as bed material, thus reducing the required amount of "fresh" bed material to a minimum. The plant dynamics have been tested with load acceptance tests, change-over of gas turbine he1 during operation from diesel oil to bio-gas and vice versa. These tests have verified the calculated dynamic properties of the plant and fuel switch-over is now made as a filly automatic procedure, remote surveyed from the control room. The gas quality has been at the calculated levels during all sorts of operating conditions and the gas has a lower heating value in the range 5.3 - 6.3 MJ/Nm3. Other gas components, such as tars, were not continuously monitored, but extensive measurement have been made during most of the tests, however tars have never caused any problems in the plant. An example of the influence of fuel composition on the amount of tars in the gas is given in Table I.
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Table 1 Effect of feedstock on tar
Fuel Bark 60 % + Forest residue 40 % Cellulose chips
Benzene (mg/nm3) 5000 - 6300 7000 - 9000
Light tars (mg/nm3) 1500 - 2200 2500 - 3700
The ARBRE Plant Construction of the ARBRE plant, situated just South of Selby, North Yorkshre, UK started in July of 1998. Construction is now been completed and all equipment has been installed. The plant will provide a net electrical output of 8 MWe with an efficiency of 30.6%. The sale of electricity to the grid is supported by the NNFO programme in the UK. The process flowsheet is given in Fig. 10. The supply of the coppice will be undertaken by Yorkshire Environmental Ltd., which will try to maximise the utilisation of biomass fuels supplied by short rotation forestry. Slurry of domestic treated sewage sludge supplied by Yorkshlre Water Services Ltd. will be applied to the coppice plantations as organic fertiliser to increase crop yield. The sludge will be low in heavy metal, pathogens and odour. Waste ash from both the gasifier and the catalytic cracker will be recycled to the coppice plantations as a soil conditioner and source of base cations and micronutrients to improve soil fertility. Monih
Figure 10. Process flowsheet of the ARBRE plant.
Coppice fuel will be used throughout the year, necessitating storage facilities. The chips will be in two covered warehouses adjacent to the site, and in clamps on local disused airfields. The chopped wood will be dried to 10-20 % moisture content by the low grade heat taken in the form of warm dry air from the air cooled
15
condensers downstream the waste heat boiler. This warm dry air will be led underneath a four day chipped wood store and will be blown upwards through the floor of the grain floor wood chip dryers. It is expected that drymg will take up to two days to be completed. A lock hopper system will feed the chips into the TPS circulating fluidized bed gasifier operating at between 850 - 900 OC. Sealing gas is introduced to the fuel feeding system at a flow rate sufficient to prevent backward gas leakage from the gasifier. In the TPS process the tars are cracked catalytically to simpler compounds in a second circulating fluidized bed reactor, which is built in the same way as the gasifier, is and has similar dimensions. The gas is introduced together with some air at the bottom of the catalytic reactor and is contacted with the bed material, which consists of dolomite at temperatures of about 900 OC. The gas is then cooled to 180°C in a cooler, which is used to generate high pressure saturated steam. The gas leaving the cooler enters a conventional high efficiency bag-house utilising needle fibre bags. The gas is then supplied to a combination cooler and scrubber where its temperature is lowered to 25 "C and any remaining alkalis, naphtha and ammonia are removed. Liquid effluent from the gas scrubber is treated in a wastewater treatment plant. Facilities will be provided for pelletisation of dolomite containing ash for use as slow release fertiliser to be applied in the energy plantations. The gas after compression to about 20 bar is fired in an ABB Alstom Typhoon gas turbine (the same as in the Varnamo plant). The exhaust gas from the gas turbine exits at approximately 475 "C. The steam produced is used to generate 5.5 MWe in a steam turbine. Waste heat is used for drying the biomass fuel.
CO-FIRING WITH COAL Co-firing application are perhaps the most interesting at present for an accelerated market penetration potential as the overall costs are relative low due to the existence of the power cycle in the coal fired power plant. In addition, co-firing has the advantage over co-combustion, where the biomass fuels are mixed with coal before or during the combustion process, that the biomass residual ash is not mixed with the coal ash, which has an existing market as a construction material. Also the technical risks are low as the gas is utilised hot and therefore there is no tar problem. In reburning applications, (when the fuel gas in introduced almost at the top of the coal boiler) it has been shown that the environmental performance of the power station is significantly improved in addition to the replacement of fossil fuels by renewable biomass fuels (60).
The Lahti Plant The utility Lahden Lampovoima Oy has built a FOSTER WHEELER CFB gasifier for its Kymijarvi Power Station in Lahti, Finland. The gasifier uses industrial waste wood, chips, fuel peat and recycled fuel (REF) as fuel. The gasifier-produced gas is burned in the Kymijarvi Power Station's steam boiler with a high flame temperature, guaranteeing the purity of the flue gases. The objective of the gasifier is to replace 50 MW of the Power station's steam boiler's 350 MW fuel effect by biofuels (Fig. 11). One third of the gasifier fuels is recycled fuel (REF), which is in-origin classified refuse from households and industry and the other two- thirds is composed of different kinds of biomasses. One possible fuel in the future is shredded tires, which has been tested during the past operating seasons. By using biofuels it is possible to
16
reduce the power station's emissions and diminish the environmental hazards. The decrease of carbon dioxide has been calculated to be 60 - 80 000 t/a. The product gas for combustion is led directly from the gasifier through the air preheater to two burners that are located below the coal burners in the boiler. The gas is burned in the main boiler and it replaces part of the coal. When the fuel is wet, the heating value of the gas is very low. Typically, when the fuel moisture is about 50 % the heat value of the gas is only about 2.2 MJ/m3n. The combustion air for the burners is extracted from the main boiler air supply. Air is divided into primary and secondary air to ensure optimum combustion conditions. With regard to the gasification plant itself, the problems faced were related mostly to the use of shredded tires as a fuel in the gasifier. In addition to that it has been observed that the light fuel fractions can occasionally cause some fluctuations in the fuel feed rate to the gasifier. On several occasions the wire content of the tires (there is no wire removal process, such as magnetic separation, after tire shredding) was so high that accumulated wires blocked the ash extraction system and the gasifier had to be shut down. However, in general the operation of the gasification process was good.
Figure f f Schematic of the Lahti Co-Firing plant Concerning the gasification process itself, the results have met expectations. The operating conditions as measured by temperatures, pressures and flow rates have been as designed and the process measurements of the product gas, bottom ash and fly ash compositions have been very close to the calculated values. The operating temperature of the gasifier has been typically 830 - 860°C and the feed flow rates of bed materials as designed. Typically the gasifier effect varied between 35 MW and 55 MW depending on the gasifier fuel moisture content and on the required gasifier load. The stability of the main boiler steam cycle has been excellent. The large openings that were made for the low Btu gas burners have not caused any disturbances in the waterheam circulation system. Furthermore, as regards the operation of the product gas burners, the product gas combustion has been stable even though the moisture content of the solid fuel has been mostly high and the heating value of the gas very low. The stability of the main boiler coal burners has been normal despite the fact that the product gas burners were integrated very close to the 17
lowest level coal burners. The main boiler emissions were perhaps of the greatest interest as regards the measurement program of the monitoring phase. In summary,it can be stated that the changes in the emissions were rather low. As indicated earlier, the main boiler is not equipped with DeNOx or DeSOx plants and today the limit values for the emissions are as follows: NOx 240 mg/MJ (as N02) and SOX 240 mg/MJ. Table 2 summarizes the effect of the co-combustion of the gasifier product gas on the main boiler emissions (60).
The BioCoComb Plant in Zeltweg This demonstration plant has been installed at the Zeltweg power plant operated by DRAUKRAFT (15). The BIOCOMB process is designed for preparation of biofiels for co-combustion by partial gasification and attrition due to mechanical and thermal stress in a circulating fluidized bed reactor (CFB) (Fig. 12). Table 2 The effect of the Lahti gasifier to the main boiler emissions.
Emission NOx
Change caused u a s i f i e r ' T D z r e a s e d by 10 mg/MJ (= 5 to 10 %) 1 Decreased by 20 - 2rmg/MJ
sox
__.__I_
1 Increased by 5 mg/MJ * 1 No change
--HCI
co
-
~
I
^
-
--
1
-.. -
-
-
_
~
-
I
-
_
Particulate
Decreased by 15 mg/m3n Heavy metals -----___ Slight increase in some elements, base level low Dioxins Furans PAH No change Benzenes Phenols -__.
_"
-"
coal
biomass -
itQ h mill
4 ash Figure 12 Schematic of the BioCoComb plant in Zeltweg
18
I
~
-
-
,
The product gas is fired in the furnace of a coal-fired power plant. The portion of biofuels reaches to about 3-5% of the total thermal input corresponding to 10 MWth. The fluidizing medium of the CFB is hot air, which is taken from the air preheater of the power plant. The CFB reactor operates at conditions where the biomass will be partly combusted and partly gasified at temperatures between 750 and 850 "C. The produced char is ground by mechanical attrition and thermal stress to a fine powder. The attrition is maximised by optimal operating conditions of the CFB. The cyclone of the CFB-reactor is designed in such a way that only char particles, which are small enough to burn completely in the coal fitmace pass the cyclone. They are fed into the furnace together with the hot gas, which contains combustible components from the gasification. Larger char particles will stay together with the bed material - in the CFB process until they are small enough or gasified completely. The efficient combustion system of the boiler combined with the very efficient flue gas cleaning system of the plant guarantees a minimised impact on the environment. The substitution of part of the coal by biomass reduces COz emission from fossil fuels, too. There is a high potential for the fuel gas to be used as a reducing gas in'the reburning zone of the combustion chamber and thus reduce or even avoid other additional De-NO, measures (reduction of NH3 consumption). More than 5,000 tons of biomass and supplementary fuels have bee gasified since start up. Main base fuel was spruce bark with moisture content of about 55% but also chopped wood and sawdust. Operating experiences are very positive with gasification as well as the combustion of the gas been according to design. The critical changeover from gasification to combustion and reverse is smooth with a slight and acceptable temperature increase. The power range of the gasifier was varied between 5 and 20 MWth, the maximum load depending on the moisture content of the biomass fuel. The quality of the fuel gas was similar to pre-calculated values while the bum out of carbon is excellent with almost no carbon found in the discharge bed material. The reburning effects in the boiler have an astonishing good performance, where a decrease of 10-15 % of the ammonia water consumption was gained with only 3% of the total thermal input coming from biofuels (62). The AMER Project This wood gasification plant is located at the relatively new coal fired heat unit of EPZ, Amer 9 at Geertruidenberg in The Netherlands. The aim of the project is to achieve savings of 70,000 tons of coal (33,700 TOE) based on 150,000 tons of wood waste (construction & demolition waste) corresponding to 170,000 C 0 2 t/y reduction in emissions. The equivalent electricity capacity is 29 MWe. The gasification plant has been supplied by LURGI. Chipped demolition wood is transported to the plant by ship and by truck is stored in silos after passing a rotating disk separator (for separating off big parts) and a magnetic separator. From the storage silos the chips are continuously transported to two day silos, feeding two screw conveyor feeding system of the gasifier. The gasifier is of the atmospheric circulating fluidized bed type, operating at temperatures of 800 - 950 "C with the addition of bed material and possible limestone or dolomite. After passing the cyclone, the raw product gas is cooled down in a gas cooler to a temperature of about 200 "C; in that cooler intermediate temperature steam is produced and slightly superheated. At the temperature of 200 - 250 "C, the gas is dedusted in a bag house filter. The dust-free gas is then washed with water in a scrubbing section, in order to remove mainly the ammonia. After this scrubber, the 19
gas is reheated to about 100 "C and at that temperature it is fed to special burners in the existing coal fired boiler of Amer unit 9. The wastewater from the scrubbing section is stripped to remove the ammonia. The ammonia is recycled to the gasifier. From the waste water system, a bleed stream is injected into the coal fired boiler. The plant has been commissioned, however, it has not entered into full scale operation since problems of slug formation have been reported on the gas cooling unit (63). Furthermore, the owner plans to install a second feeding system dedicated for difficult to handle biofuels such as RDF and grasses. The objective is to increase the capabilities of the plant to operate under multi-fuel conditions in order to ensure competitive feedstock cost and flexibility.
.-
Ceah 31h
m-Pnar*."
..........
,
(...... . A-..........i 3 -
...y
.............. : -+be
1
*.*I
hrrn
................... ....
........J
L
'
.H-
"
................
I
.........._ ........-
I
Figure 13 Process flow diagram of the AMER gasification plant CO-FIRING WITH NATURAL GAS As with coal, fuel gas produced by biomass gasification can be co-fired with natural gas either directly in turbines, boilers or duct burners or as reburning fuel. Very little work has been published on this issue, however, this could significantly enlarge the market options for biomass gasification. Calculations show a substantial increase in flame temperature, laminar burning velocity and lower heating value by adding 25% methane to gasifier fuel gas (64).
INDIRECT FIRING OF GAS TURBINES One innovative way of eliminating the tar problem for relative small scale applications is the indirect cycle or hot air gas turbine process. In such a configuration, the fuel gas produced by the gasifier is combusted directly in a heat exchanger where clean air supplied by the compressor of the turbine is heated up to the range of 850" to 950' C. The hot air is then fed to the gas turbine. The main problem with this type of process is the size and operational problems including fouling and eventual corrosion of the heat exchanger.
20
The Free University of Brussels project, BINAGAS The Binagas project is located at the campus of the VUB in Brussels. The gasification train consists of the feeding system, an atmospheric fluidized bed gasifier operating in the temperature range of 725 to 850 "C and a cyclone for the removal of particulates. The gas produced is fed to the combustion chamber of the heat exchanger through insulated lines and a high temperature valve, which isolates the gasification train (when the fuel gas is flared) and the metallic air heatedgas turbine system (Fig. 14). The operating procedure has been designed in order to avoid any condensation of tars with well defined starting up and shut down sequences (65). The compressor of the Volvo gas turbine supplies the air through the heat exchanger where it is heated up to 850 "C, whch is the limiting temperature dictated by the construction materials. Some natural gas topping combustion is included to raise the temperature to about 1,000 "C. Water injection in the air heater is included to enhance the power output and allow flexible power to heat ratios. The demonstration scale is 500 kWe, for production of power and heat for the University campus district heating. Targeted maximum performances are 24% electrical and 70% total. When injecting water at power mode, peak electrical efficiencies of 30% and power output of 700 kWe are expected. Iw tenperature
---
-
air airandLjquiwter
high tenpesature
WI I
Figure 14 Process flowsheet of the VLTB plant. The complete plant has been operational for several hundred hours, but control problems were experienced with the gas turbine which was a prototype. Also the particulate accumulation in the heat exchanger has been higher than expected but no major problems have been encountered with this part of the plant. The gasification island operates efficiently and has met the design specifications, even though it is basically fed with sawdust instead of pellets. At present the modification to the gas turbine have been completed and the test program will restart. The W B plans to use the installation as a test facility until all the problems have been overcome.
The Freiberg Project Pipeline Systems GmbH (PPS) has constructed an IGCC indirect fired wood
21
gasification CHP plant at Siebenlehn in the Freiberg District, Saxony, Germany. PPS is the general constructor and works manager of the plant. The consumer of the process heat is an industrial user located 50 m from the CHP plant itself. The gas turbine nominal capacity is 1.3 MWe while the steam turbine’s nominal capacity is 0.7 & 1.0 MWe, respectively for extraction or condensation operation mode. The primary energy carrier is a mixture of forest cut wood, industrial wood residue and forests residues. In the co-current downdraft moving bed gasifier the fuel is fed from the top and the gasification medium is air supplied through a central blow-in pipe. The fuel gas is removed through an induced-draught fan. A small part of the fuel is burned in the oxidation zone in order to achieve the extremely high reaction temperatures, which characterise this gasifier. A rotating grate of conical shape and which is heightadjustable is built into the lower part to regulate the flow of the fuel. The firebox of the gasifier is made of a special high-temperature-resistant ceramic material. Through the special introduction of gasification air and the utilisation of high-temperatureresistant ceramic material, temperatures up to 2000 “C are generated in the oxidation zone of the gasifier. At these temperatures long-chained hydrocarbon compounds are cracked and transformed into combustible gases and the mineral and metallic components of the fuel are fused into an inert, glasslike slag (66), (Fig. 15). b
Power station
Gasijkation section
Wood storage
Figure 15 Schematic of the TU Bergakademie Freiberg plant
The fuel gas with a higher heating value of 5 MJ/Nm’, has a high concentration of H2 (up to 20 Vol%). Formation of tar and contamination of the he1 gas is avoided to the greatest extent possible through the high temperatures in the reaction zone. This reduces the cost of any subsequent gas cleaning. The slag, which is formed corresponds to the storage regulations of mineral andor inert material disposal sites. The gas cleaning consists of coarse dust removal by a cyclone and of a fine filter. The dust removed is made up of up to 99% carbon. Fine dust which passes through the gas cleaning devices is burned in the subsequent combustion chamber together with the fuel gas. The combustion of the wood-derived gas takes place at atmospheric pressure in an external combustion chamber. In a high-temperature heat exchanger, the combustion gas which arises raises the temperature of compressed air, which is to be the working medium of the gas turbine, up to the gas turbine entry temperature. The gas turbine aggregate consists mainly of an air compressor, the expansion unit
22
(turbine), reducing gear, a generator and the auxiliary systems required for operation. The flue gas whch emerges from the heat exchanger and a partial flow of the gas turbine exhaust are channelled to the heat recovery steam generator. The utilisation of t h s exhaust beat for steam generation corresponds to the state of technology of the conventional gas and steam turbine process. In the heat recovery steam generator, equipped with economiser, evaporator and superheater, the fresh steam is generated for the steam turbine. A flue gas heat exchanger for generation of hot water is attached in series to the heat recovery steam generator for extensive utilisation of the exhaust gas energy and for minimisation of the flue gas temperature at the stack. Clean hot air suctioned by the compressor of the gas turbine and then compressed to the required turbine entry pressure serves as the working medium of the gas turbine. A high temperature heat exchanger is placed between compressor and gas turbine, in whch the air is heated to the h g h gas turbine entry temperature. This heat is supplied by the combustion of the fuel gas in a combustion chamber. The exhaust of the gas turbines provides the combustion air for this, whereby an increase of the combustion chamber temperature is also achieved, due to the high energy potential. The gas turbine exhaust not needed for combustion is fed through a bypass to the heat recovery boiler for waste-heat recuperation, as is the cooled down combustion gas in the high temperature heat exchanger. The heat recovery steam generator produces fresh steam for the extraction-condensation steam turbine, the extraction steam of which serves to cover the process heat requirements of the industrial user through its transformation into hot water. The plant has been commissioned and the first results on the performance are expected soon. Design data give an efficiency to power of 23 and 27 % respectively for extraction and condensation mode of operation whle the overall plant efficiency is 73 and 39 % respectively.
ENGINES Work for running engines with producer gas has been continuing for decades now but with few breakthroughs if any. In the field of reciprocating engines there have been two engine manufacturers who are the key players - Jenbacher and Caterpillar (67). The main problem relates to efficient removal of tar, however, the engine manufacturers have not been able to design and construct more robust engines, whch can tolerate some tar in the gas. In developing economies like Chna, where there is h g h demand for energy and cheap labour, there are examples where engines are carefully and continuously maintained and are operated with a relative dirty gas (68). This approach results however in significant quantities of condensate which accumulates and causes an environmental hazard. This is not acceptable for developed economies and overall there is little development to report on engines. In recent configurations, the product gas is kept above its dew point (>75 "C) throughout the installation downstream of the cooling section in order to avoid condensation of any water and any remaining hannful tar components. T h s eliminates the need for an extensive water treatment plant. The engines proposed are lean-bum, high-speed, high-efficiency, turbo-charged gas engines. The turbo-charger is mounted before the engine in order to increase the gas pressure, and therefore the gas energy density, compensating for the loss in efficiency due to the lower density fuel gas caused by the relative high temperature of the fuel gas (> 75 "C). The electrical efficiency of such a configuration can reach 40%. Work on gas turbines has proven successful with the operation of the TYPHOON
23
gas turbine at the Varnamo plant and the expected operation at the ARBRE plant in the first half of 2001 (59, 41). In addition, NUOVO PIGNONE has carried out a successful series of combustion chamber tests for a 10 MWe gas turbine for the Energy Farm project in Italy (57); however, the gas turbine has not been built yet and it is only planned to be commissioned in 2003 if this project proceeds to completion. This is an area which presents a dilemma for the gasification community since the gas turbine manufacturers do not wish to develop more robust engines that could operate with some contamination in the gas as the market is still considered very small. Thus the main task falls on the gasification industry to deliver a ‘clean’ gas to the combustion chamber of the gas turbine. On the other hand this has prompted the gasification industry to develop gas cleaning technologies which, if proven reliable, will form the basis for synthesis gas from biomass. This will open new opportunities and markets for bioenergy and especially gasification technologies as discussed below.
METHANOL, HYDROGEN & FISCHER-TROPSCH These chemicals as well as energy vectors can be used in several applications or can be W h e r upgraded into other useful products. They can all be produced via synthesis gas (CO + H2) that has been the subject of extensive investigations and commercial industrial processes based on fossil based synthesis gas exist. The advantage of these vectors is that they can be either used in fuel cells for electricity or transport applications, or alternatively, they can be processed to liquid transport fuel additives such as dimethylether (DME) and dimethoxymethane (DMM). Advanced biomass gasification processes can also produce synthesis gas as has been demonstrated in the 1980s by the “Methanol from Wood” programme of the European Commission (69). During that work four pilot plants were operated at design capacities ranging from 4.8 to 12 dry t/d (70). These plants were developed by Framatome, Lurgi, John B r o f l e l l m a n & Italenergie while TPS developed the MINO process. All technologies were based on fluidized bed with various mixtures of 02,steam and air while the John B r o f l e l l m a n process used an innovative approach of chemically active solids as an O2 carrier. Similarly various groups have looked into the production of H2 from gasification by a slurry of hydrated metal alloys (70). The National Renewable Energy Laboratory in the USA has also investigated extensively the conversion of synthesis gas to energy hels with positive results (71). In principle, if a clean synthesis gas can be produced from biomass, there should not be any serious technical barrier for it’s subsequent conversion to methanol or Fischer-Tropsch liquid products as these processes have been demonstrated to some extent by the novel methanol to gasoline process in New Zealand or by SASOL in South Africa (72). Various Dutch research centres have proposed a once through configuration for methanol combined with an IGCC fuelled by the lignin by-product of an ethanol facility in order to improve the overall economics and process efficiency while simultaneously producing ethanol (73). However, such schemes are only potentially viable at large scale and require a concerted approach by several stakeholders. Others have looked into the methanol route specifically (74), but these studies are still in the laboratory scale and a significant amount of work is required before industrial scale processes could be developed. On the other hand the results are promising. A recent extensive techno-economic analysis of the various transport fuel chains from biomass (75) concluded that the most promising chains were ethanol substitution
24
of gasoline followed by Fischer-Tropsch produced diesel. On the other hand a similar study in Sweden (76) resulted in DME as the most promising liquid biofuel. What is of importance is that the main car manufacturers are seriously considering these liquid biofuels and try to develop dedicated engines. A ROAD MAP FOR GASIFICATION Gasification technologies offer huge potential as they can produce energy and chemical vectors from a variety of lignocellulosic materials. These vectors can be used in numerous applications directly or after further processing and upgrading and can be either in the gaseous state or liquid state depending on the processes and applications. This section attempts to provide a road map and benchmarking of the gasification technologies in relation to existing and possible future markets and as a whole it can form the basis of a strategy for the commercialisation and market penetration for gasification applications. Figure 16 shows the relationship between technology and markets. Each of the four quadrants relates new and existing technology to new and existing markets.
EXISTING MARKETS
NEW MARKETS
EXISTING TECHNOLOGY
NEW TECHNOLOGY
Market Penetration
Product Development
Co-Firing Firing Waste + BM
M eth ano 1 H2 Fischer-Tropsch
Market Development
Product Diversification
IGCC - 30-75 M W e Gasification 1-5 M W e Hot Air Gas Turbines
Chemicals from Biomass Monomer recovery Materials recovery
Figure f 6 Developing growth strategies and market opportunities At present the most reliable applications for biomass gasification are co-firing and direct firing of the fuel gas in a boiler for heat or steam cycle. These applications present the least technical risks as the problem of tar is avoided and therefore the main task of the industry is to increase their market penetration. It is important that more plants will be built and operated in existing markets in order to increase the degree of confidence for the users and especially the utilities as well as to improve the industrial capabilities with various problematic but cheap fuels. The bottom left quadrant in Figure 16 shows the applications for which new markets must be developed based on the expected success of a few key projects, which are now in the demonstration phase. Commercial IGCC, medium scale gasification and hot air gas turbines are expected to become commercially available in the short to medium term of about 4-6 years. If this happens, then the gasification technology will be able to address all basic requirements for energy applications
25
either as biomass dedicated plants or in combination with fossil fuels. New technology development is needed for chemicals and liquid biohels, whch can have direct utilisation in existing structures of modem society and especially as transport hels. Although bio-ethanol and biodiesel from grain and seed respectively are supposed to be more competitive, various studies indicate that Fischer-Tropsch diesel and DME can be competitive in the medium to long term of about 5-10 years. The progress achieved at Vamamo gives confidence that the gasification industry would be in position to deliver a clean gas for further processing to synthesis gas while gasifier operation with mixtures of oxygen and steam does not appear to create any serious technical barrier. Finally hydrogen will always remain the cleanest fuel of all and any process that can produce hydrogen form biofuels under economically competitive conditions will be an immediate market success. The last quadrant in Figure 16 shows the areas for eventual product diversification from energy and/or liquid biofuel vectors. Although these do not appear to offer great market potential they can address niche markets for the production of bio-chemicals andlor the recovery of the valuable products such as the monomer from waste polymers and aluminium from drink packaging. Considering that financial resources are limited and there is very strong competitions with other bioenergy technologies such as fast pyrolysis and combustion, in addition to other renewable energy sources which may be more appealing to the general public and even utilities and authorities such as wind power, it is strongly advised that the finite financial resources available should be carefully targeted to address actual technical and market barriers. It is therefore recommended that the target areas would be those included in the bottom left and top right quadrants of Figure 16. It must be repeated once more that the main barrier remains the delivery of a clean gas.
CONCLUSIONS Biomass gasification technologies have reached the point where the first simple applications with minimal technical risks are becoming commercial. In addition the first biomass based IGCC plants are being demonstrated and are expected to reach commercial status within about 5 years. Future market opportunities exist for liquid biohels production via synthesis gas, however, although these systems have been examined periodically since the 1980s, a significant amount of work has still to be done before such plants could be considered by the financial community. The development of the technology has moved beyond the element of the “gasifier” to the critical area of the supply of a “clean gas”, free of particulates and tar. If this will be achieved, then the power market will be slowly but steadily penetrated on condition that sufficient feedstocks can be secured. After this stage and on the basis of the various policies concerning liquid biofkels by the European Union, USA, Canada, China and India synthesis gas could become an important market. At this stage the technology will move from the supply of “clean gas” to the production of “synthesis gas”. This would necessitate a dedicated policy for the production of very large quantities of biomass fuels to satisfy the demand for power and liquid biofuels.
26
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46.
gasification) elimination. In Proceedings 1" World Biomass Conference, In Proceedings 1'' World Conference h Exhibition on Biomass for Energy & Industry, (Ed. by S . Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. Corella, J., et al., (2001) Testing commercial full-size steam reforming catalysts for tar elimination in biomass gasification at pilot scale. In this proceedings. Paul.P.J., et al., (1997) Tar characterisation in new generation agro-residue gasifier-cyclone and downdraft open top twin air entry systems. In Biomass Gasification and Pyrolysis, State of the Art and Future Prospects, (Ed. by M. Klatschrmtt and A.V. Bridgwater), Cpl Press Newbury. Zanzi, R., et al., (2001) Rapid pyrolysis of biomass at high temperature as the initial stage in gasification. In Proceedings 1" World Biomass Conference, In Proceedings I" World Conference h Exhibition on Biomass for Energy di Industry, (Ed. by S. Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. Simell, P., & Kurkela, E., (1997) Tar removal from gasification gas. . In Biomass Gasijication and Pyrolysis, State of the Art and Future Prospects, (Ed. by M. Klatschmitt and A.V. Bridgwater), Cpl Press Newbury. Kurkela, E., (1996) Formation and removal of biomass-derived contaminants in fluidized bed gasification processes.. VTT Energy publications 287, Espoo. Padban, N., et al., (2001) Tar formation in a 18 MWth biomass IGCC plant in Varnamo and in a 90 kWth pressurised fluidized bed gasifier at Lund University. This Proceedings. Beenackers, A.A.C.M.& Maniatis K. (1997) Gasification Technologies for Heat and Power from Biomass. In Biomass Gasification and Pyrolysis, State of the Art and Future Prospects, (Ed. by M. Kaltschmitt and A.V. Bridgwater), Cpl Press Newbury. McLellan, R., (1997) Welman Biomass gasification technology. In Proceedings 4Ih International Wood Fuel Conference, DTI, London. Knoef, H., Private communication from BTG, and EC ENERGIE contract NNES- 1999-528. Kaltschmitt, M., et al., (1998) Chapter 4: State of the art of biomass gasification. In Biomass Gasification in Europe, European Commission, DG XII, EUR 18224 EN, Luxemburg. Pitcher, K., & Weekes, A., (2001) Arable Biomass Renewable Energy (ARBRE)-The development of a biomass gasification combined cycle plant. In Proceedings 1'' World Conference h Exhibition on Biomass for Energy h Industry, (Ed. by S. Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. Madsen, M., Private communication from FLS Miljoe and ENERGIE contract NNE5-2000-124 Kalogeropoulos, P., Private communication from Envitec and EC ENERGIE contract NNE5-2000-3 12. Maniatis, K., & Beenackers, A.A.C.M., (2000) Tar Protocols. IEA Bioenergy Gasification Task, Editorial. In Biomass h Bioenergy, 18, No 1, 1-4. Abatzoglou, N., et al., (1999) The development of a draft protocol for the sampling and analysis of particulate and organic contaminants in the gas from small biomass gasifiers. In Biomass and Bioenergy, Vol. 18, 5-1 7, Pergamon, Oxford. Simell, P., et al., (1999) Provisional protocol for the sampling and analysis of tar
29
47. 48. 49. 50.
51. 52.
53. 54.
55.
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59. 60. 61. 62.
and particulates in the gas from large scale biomass gasifiers. Version 1998. . In Biomass and Bioenergy, Vol. 18, 5-1 7, Pergamon, Oxford. Hasler, P., et al., (1998) Gas cleaning and waste water treatment for small scale biomass gasifiers. Final report, Swiss Federal Ofice of Energy & Swiss Federal Oficefor Education a Science, VEEZENUM Research, Zurich. Knoef, H.A.M. & Koele, H.J., (1998) Tar Measurement Protocol: Realisation of a standard procedure for tar & particle content determination in producer gas from biomass gasifiers. Report E WAB9832, NOVEM, Utrecht. Milne, T.A. et al., (1998) Biomass Gasifier Tars: Their nature, formation and conversion. Report NREL/TP-S70-2535 7, NREL, Golden Colorado. Stahlberg, P., et al. (1998) Sampling of Contaminants from Product Gases of Biomass Gasifiers. VTTResearch Notes 1903, VTT, Espoo. Zielke, U., et al., Parallel Measurements of Tar and Particulates. Report for the Danish Energy Agency, NOVEM, VTT & Swiss Federal Office of Energy, Danish Technological Institute, Arhus. Neeft, J.P.A., et al., (2001) Tar Protocol. Development of a standard (protocol) for the measurement of organic contaminants (tars) in biomass producer gases. Thisproceedings. Kurkela, E., & Simell, P., (2000) Gasification Survey Country-Finland. In Status of Gasification in countries participating in the IEA Bioenergy gasification activity (Ed. KW. Kwant) Gasification Task, IEA Bioenergy, NOVEM, Utrecht. Rensfelt, E., (1997) Atmospheric CFB gasification-the Greve plant and beyond. , In Biomass Gasification and Pyrolysis, State of the Art and Future Prospects, (Ed. by M. Klatschrmtt and A.V. Bridgwater), Cpl Press Newbury. Scoditti, E., (2000) Gasification Survey Country-Italy. In Status of Gasification in countries participating in the IEA Bioenergy gasification activity (Ed. KW. Kwant) Gasification Task, IEA Bioenergy, NOVEM, Utrecht. Paisley, M.A., et al. (2000) Preliminary operating results from BattelleEERCO gasification demonstration plant in Burlidington, Vermont, USA. In Proceedings I" World Conference & Exhibition on Biomass for Energy & Industry, (Ed. by S . Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. de Lange, H.J. et al., (2000) The realization of a biomass helled IGCC plant in Italy. In Proceedings I" World Conference & Exhibition on Biomass for Energy & Industry, (Ed. by S. Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. Carpentieri, E., & Silva, A., (1998) WBPBIGAME the Brazilian BIG-GT demonstration project actual status and perspectives. In Biomass and Bioenergy, VoI. 15, N"3, (Guest Ed. A.A.C.M. Beenackers & K. Maniatis), Pergamon, Oxford. Stahl, K., et al., (2000) Final report: Vamamo demonstration programme. In this proceedings. CRE Group Ltd, (2000) Technical review on opportunities and markets for coutilisation of biomass and waste with fossil fuels for power generation. Report prepare for the European Commission, Brussels (in print). Kivela, M., (1999) Final report: The Lahti gasification plant. Report by Lahden Lampovoima Oy, Lahti. Mory, A. & Tauschitz, J., (2000) BIOCOCOMB-Gasification of biomass and co-combustion of the gas in a PF boiler in Zeltweg power plant. In Proceedings EU Seminar The Use of coal in mixture with wastes and residues II, (Ed. V.
30
Breme), BEO, Julich. 63. Willeboer, W., (2000) AMERGAS biomass gasifier starting operation. In Proceedings EU Seminar The Use of coal in mixture with wastes and residues II, (Ed. V. Breme), BEO, Julich. 64. Fossum, M., & Beyer, R.V., (1998) Co-Combustion of natural gas and low calorific value gas from biomass. SINTEF Energy Research, Report prepared for IEA Biomass Gasification Activity, Trondheim. 65. De Ruyck, J. et al. (1996) An externally fired evaporative gas turbine cycle for small scale biomass gasification. In Biomass for Energy & the Environment, gh European Bioenergy Conference,(Ed. P. Chartier et al.) Pergamon Oxford. 66. Franke, B., & Bizaj, B., (2001) Wood -Fuelled combined heat & power plant. Private communication & promotional brochure, PPS Pipiline Systems. 67. Barker, S.N., (1998) Gas turbines, reciprocating engines and other conversion devices in biomass to electricity systems. AEA, Report prepared for IEA Biomass Gasification Activity, Harwell. 68. Bridgwater, A.V., et al., (1999) An assessment of the possibilities for transfer of European Biomass Gasification Technology to China. European Commission Report, Luxemburg. 69. Beenackers, A.A.C.M. & Bridgwater, A.V., (1989) Gasification & Pyrolysis of biomass in Europe. In Pyrolysis & Gasification, (Ed. G.L. Ferrero, K. Maniatis, A. Buekens & A.V. Bridgwater ) Elsevier Applied Science, London. 70. Beenackers, A.A.C.M. & van Swaaij, W.P.M. (1986) Advanced Gasification. Reidel, Doordrecht. 71. Mills, G., (1993) Status and future opportunities for conversion of synthesis gas to liquid energy fuels: Final Report. (Ed. R. Overend & R. Bain), NREL, Golden, USA 72. Larson, E. D., (1999 Advanced technologies for biomass conversion to energy. In Proceedings 2"' Olle Lindstrom Symposium on Renewable Energy, BioEnergy, Royal Institute of Technology, Stockholm. 73. Ouwens, C.D., & Boerrigter, H., (2001) New developments in the field of trigeneration from biomass & waste: A survey. In these proceedings. 74. Sakai, M., & Kaneko, M., (2001) The development of methanol synthesis with biomass gasification. In this proceedings. 75. van den Heuvel, E.J.M.T., (2001) Bubbling with Energy: New chances for climate neutral liquid and gaseous energy carriers in the Netherlands. In Proceedings Is' World Conference & Exhibition on Biomass for Energy & Industry, (Ed. by S . Kyritsis, A.A.C.M. Beenackers, P. Helm, A. Grassi & D. Chiaramonti), James & James. 76. Waldheim, L.,et al., (2000) DME Development plant at Varnamo. TPS & SYCON, Stockholm.
31
Steam gasification of wood char and the effect of hydrogen inhibition on the chemical kinetics M. Barrio, B. G@bel+,H. Rimes, U. Henriksen', J.E. Hustad and L.H. Seremen* Norwegian University of Science and Technology, Department of Thermal Energy and Hydro Power, 7491 Trondheim, Norway +Technical University of Denmark, Department of Energy Engineering, Nils Koppel Alli, DTU-Building 403, DK-2800 Kongens Lyngby *ReaTech c/o Centre for Advanced Technology (CAT), Postbox 30, DK-4000 Roskilde
ABSTRACT: Gasification kinetics parameters have been derived for birch and beech char samples (45pm
INTRODUCTION The gasification process requires an oxidising agent that provides oxygen for the formation of CO from solid fuel. The oxidising, or gasifying, agents are air, oxygen, steam and C 0 2 .C 0 2 is produced during the pyrolysis and early oxidation processes and generally not externally added. The most common agent is air because of its availability at zero cost. Air, though cheap, is not a perfect agent because of its nitrogen content. The product gas from air gasification has generally a low heating value of 4-7 MJ/Nm3. Oxygen gasification produces a higher heating value (10-18 MJ/Nm3) but has a drawback due to the high production cost of oxygen. Steam is another alternative. It also generates a medium calorific value gas (10-14 MJ/Nm3) and moreover increases the hydrogen content of the product gas. The presence of steam is important in case of further catalytic upgrading of the product gas'. Steam gasification is however a highly endothermic reaction and requires a temperature above 800 "C to take place2 if no catalyst is present3r4.The heat required for the reaction has to be transferred either by partial char combustion in the same reactor -mixing H20 with ~xygen/air''~or by indirect heating6.'.
32
Because of biomass moisture, and steam from pyrolysis in downdraft gasification, steam will always be present in gasification whether it is used or not as a gasification agent. Hydrogen is one of the products of steam gasification and its effect on the reaction is also relevant. Some kinetic data for steam gasification of biomass have been published2.8.9,10.11.12.13.14.15.16.17.18 , but very few considering the effect of H2 inhibition19,20,21,41 The diversity in evaluation of the results from char reactivity experiments is large. The definition of gasification rate varies among researchers and so does the criteria to select the reactivity values from the experiments. Few authors'6,22 have concerns regarding this. This study presents the kinetic parameters and reactivity profiles for steam gasification of birch and beech char. The inhibition effect of hydrogen is also studied ;sing Langmuir-Hinshelwood kinetics. In addition, the influence of the treatment of the experimental results is analysed by comparing the kinetic parameters differently obtained from the same experiments. The same birch char has been used for CO2/CO ga~ification~~. The kinetic study of char gasification in H20/H2/C02/C0mixtures will be a continuation of the work presented.
THEORETICAL BACKGROUND H20/H2 REACTION MECHANISMS The overall steam gasification reaction can be represented by:
C,
+ H,O=CO+
H,
(1)
However, the reaction is much more complex and involves several steps. Numerous studies have been conducted in order to understand the mechanisms of the steam gasification reaction. The catalytic activity of the ash plays an important role in this is more complex than C 0 2 gasification because not d i s c ~ s s i o n ~ H2O ~ * ' ~gasification ~~~. only H20 is involved but also H2, C02 and CO due to the equilibrium of the water gas shift rea~tion'~'~'. Huttinger and Merdes26give a comprehensive description of the models proposed in the literature for the carbon-steam reaction. Basically, there are two models of the reaction mechanism: the oxygen exchange model and the hydrogen inhibition model. The equations involved are: C,
+ H,O kl,
C ( 0 )+ H ,
klb
C(0)
k3
>co+c,
The oxygen exchange model is based on equations 2 (reversible -klfand klb-)and 3, the traditional hydrogen inhibition model is based on equations 2 (irreversible -only klr), 3
33
and 4 and a different version of the hydrogen inhibition model substitutes equation 4 by equation 5. Each model has a different explanation of the inhibition effect of hydrogen. According to the oxygen exchange model, it is due to the equilibrium of the dissociation reaction (Eq. 2). For the traditional hydrogen inhibition model, the formation of the C(H)2 complex is the reason for inhibition. Finally, the second version of the hydrogen inhibition model involves a dissociative chemisorption of hydrogen on the active blocking them for the oxygen transfer reaction with steam. The reaction rate for the models presented is similar, with the exception of dependency on hydrogen partial pressure:
where
klb f( P H 2 ) =PH2
, oxygen exchange model
(6.1.)
k3
k4f f ( p H 2 )=-
p H 2 ,hydrogen inhibition model (traditional)
(6.2.)
k4b
k5f
~ ( P H z=-)
0.5
PH2
, hydrogen inhibition model (second version) (6.3.)
k5b
According to Huttinger and Merdes26, it is not possible to determine which is the dominating hydrogen inhibiting mechanism by looking at the reaction rate because the equations are identical, with exception of the second version of the hydrogen inhibition model. It is quite common to reduce equation 6 to the following
r=
K1PH20
(7)
1+ K 2 P H 2 0 -tK 3 P H 2
where K2 and K3 represent a ratio between rate constants but are not rate constants themselves. Other authors"*29rather use empirical equations to model the chemical kinetics. In this work, the kinetic parameters have been obtained according to the oxygen exchange model, equations 2 and 3, and also according to n* order kinetics.
INFLUENCE OF FUEL TYPE Several studies have focused on the influence of wood type on C 0 2 g a s i f i ~ a t i o n ~ ~ * ~ l * ~ * . A general conclusion is that the ash content, and steam gasification2914*17*"733~34~3s composition and its catalytic properties explain the differences among the fuels. In particular, Hansen et a1.20refer to the potassium content of the ashes as being especially relevant. Moilanen et al.I4 present their results from steam atmospheric gasification of chars from different origins: wood, black liquor, cellulose fibres, peat and coal. All chars, apart from peat, present an increasing reaction rate with conversion, especially wood.
34
Stoltze et a1.I' find that the gasification of hardwood is 2-3 times slower than straw, probably due to the different char structure and composition. However, since the density of the hardwood char is 5 times higher than the one of straw, in a volume basis the reactivity of wood char is double than of straw. The direct consequence of this fact is that the gasifiers for wood char only require half the volume of a straw gasifier. Finally, it is important to mention that the pyrolysis conditions also have influence on the char reactivity, as several investigations have proved.
TREATMENT OF THE EXPERIMENTAL RESULTS There are two definitions of the reactivity commonly used:
where m, is the char mass at the beginning of the gasification and mf is either negligible, or represents the mass of ash, or - as in this work- the residual mass after gasification. The degree of conversion is obtained as:
Therefore, the relation between the two definitions of reactivity presented above is: r, = r * ( l - X )
(1 1)
It is widely accepted that the reactivity depends on the degree of conversion but there is no agreement about how to define one representative value of reactivity for each experiment. The representative value of reactivity from an experiment is most frequently obtained as the average reactivity between two degrees of conversion: 0-50%36,070%16,0-75%1°,40-60%20, 10-50%37,60-80%30. Bandyopadhyay et aL3* selects the representative value of reactivity as the reactivity at 5% conversion. Using an earlier value might introduce error because of the gas changing, but a later value would not correspond to a known condition of the sample inside the sample cup holder (depth, mainly). Stoltze et a1.16 propose a mass-weighed mean reactivity in order to give less importance to the latest stages of conversion. Finally, other researchers consider the reactivity as a function of the chemical reactivity, dependent of temperature and reactants partial pressure but independent of conversion, and of a structural factor, solely dependent on the degree of c o n ~ e r s i o n ' ~ . ~ ~ . Still, it is possible to find other methods to obtain reacti~ity"~''.
35
EXPERIMENTAL SECTION Kinetics for a Norwegian birch and a Danish beech have been determined. Apart from their origin there are also other differences between the woods. The beech sample is first received as wood chips whose surface has been exposed to the ambient and that partially contains bark. The birch sample comes from a wood log that has been cut into small cubes of lxlxlcm, removing the bark. The proximate and ultimate analysis is shown in Table 1 and the ash analysis in Table 2. Table 1 Proximate and ultimate analysis of birch and beech wood. Proximate analysis Birch wood Beech wood Ultimate analysis Birch wood (wt%, mf) Beech wood (wt%,mf)
Moisture
Fixed carbon
Ash
11.13% 14.16%
Volatile matte; 78.7%, mf 75.2%, mf
20.9%, mf 24.2%, mf
C
H
N
0.37%, mf 0.56%, mf 0 (by diff.) 44.45 44.82
48.7 6.4 0.078 48.1 6.4 0.081 * Pyrolysis conditions: Heating at 24 "Umin until 600 "C,held for 30 min, natural cooling.
Table 2 Ash analysis of birch and beech wood (%). Species Beech Birch
Species Beech Birch
Si 1.2 0.03
A1 0.14 0.01
P
C1
2 3.4
0.29 0.03
Fe
Ca
Mg
K
Na
Ti
1.8 0.17
25 30
7.1 4.8 Ni 0.02 0.01
28 28
2 0.08
0.029 0.007
Pb
Cd
Hg
Cu 0.03 0.06
Zn 0.2 0.06
S 0.75 0.64
0.01 <0.001 <0.001 0.02 <0.001 <0.001
Both woods have been pyrolysed at the Technical University of Denmark, Department of Energy Engineering (DTU, ET), in a macro-TGA, heated at 24 "C/min to 600 "C, held at that temperature for 30 min and then cooled down to room temperature naturally. Both chars were thereafter crushed and sieved to 45-63 pm. The instrument used for the reactivity study is a Pressurised Thermogravimetric Analyser (PTGA) at ReaTech, a modified Du Pont Thermogravimetric Analyser. The sample (-5 mg) is placed on a small platinum tray, hanging on a horizontal balance arm. The sample temperature is measured with the help of two thermocouples, near to, but not in contact, with the sample. This investigation is limited to atmospheric pressure although the instrument is prepared for high pressure operation. Rathmann et al.40 and S~rensen'~ give a detailed description of the PTGA and Hansen et aL2' describe the modifications required for the instrument to tolerate steam. Once the char sample is introduced into the PTGA, it is first dried in N2 during 10 min at 200 "C, then is heated at 24 "C/min to 1000 "C and held at this temperature for 30 min. After this the sample is cooled to the gasification temperature and when conditions are stable, the steam is allowed into the reaction chamber. The sample is hold isothermal until the gasification reaction is complete and then the temperature is raised to 1000 "C to complete the reaction. The sample size is ca. 10 mg and the gas flow 1000 mumin. The objective of increasing the temperature up to 1000 "C previous to gasification is to simulate the history of the particle in the two-stage gasifier at DTU, ET. This is also the reason for the heating rate of 24 "C/min. During the 30 min. period at 1000 "C
36
in the nitrogen atmosphere some fraction of the catalytic species K and Na devolatilise and are carried away from the sample and therefore the char could be less reactive. The experimental matrix for this investigation is shown in Table 3. Table 3 Experimental matrix for H 2 0 gasification experiments
I
n.s I 00 I 000 I 00 I 01 (0,0,@,0) Birch char; (0,0,@,@) Beech char. The numbered symbols indicate the partial pressure of hydrogen (x10 bar).
The design of the installation is described in Fig. 1 .
Manual O d o f fvalve Manual Odoff valve
Electrovalve with purge
d Two-way manual valve Flow controller
Fig. 1 Schematic drawing of the installation set-up.
RESULTS AND DISCUSSION GASIFICATION RATE Fig. 2 shows the mass loss curve for one of the experiments. The initial and final weights for the gasification reaction are also indicated. Fig. 3 shows the reactivity as a function of the degree of conversion, i.e. the reactivity profile, for the same experiment according to equations 8 and 9. In addition, the figure shows the average reactivity (from eq. 8) between 20 and 80% conversion.
37
It is important to notice that the shape of the reactivity profile is very dependent on the reactivity definition. For the following discussion, the reactivity has been obtained according to equation 8.
1000
800
600 t 400
200
0 time
Fig. 2 Temperature and weight signal as a function of time. Experimental data. 4 OE-04
3 5E-04 3.OE-04 2.5E-04 'm
z
2.OE-04 1.5E-04 1 .OE-04 5.OE-05 O.OE+OO
4 0
I
I 0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
Degree of conversion (X)
Fig. 3 Reactivity as a function of conversion. (0:Average reactivity between 20 and 80%conversion).
PURE STEAM EXPERIMENTS Fig. 4 shows the reactivity of the pure steam experiments as a function of temperature and steam partial pressure. The representative reactivity value has been obtained as the reactivity at 50% conversion. The continuous line shows the n' order reaction model for the birch experiments. The figure shows that beech is more reactive than birch at
38
low temperatures (750-800 "C). The kinetic parameters obtained according to n' order kinetics are shown in Table 4 together with results from other references. 1.OE-02 ,
, , ,
,
,
, , ,
, , ,
, , ,
, ,
,
,
,
,
,
,
, , ,
,
,
,
I
,
,
,
,
,
I
,
1 .OE-03
u)
Y i 1 .OE-04
1 .OE-05
0.3
0.1
0.5
1
PWO(bar)
Fig. 4 Reactivity as a function of steam partial pressure and temperature. (Filled symbols: beech, hollow symbols: birch).
Table 4 Kinetic parameters comparison for steam gasification experiments.
Reference This work' This work' Capart et a1.12 Hemati et al.13 Richard et a1.* Li et a1.9 Whittylg Timpe et a1.l'
Char origin Birch Beech Woodchar Woodchar Fir wood Black liquor Black liquor Poplar Cattails Moilanen et al.'4*15 Wood Black liquor Stoltze et a1.16-17 Straw (Large TGA) Wood chips Rensfelt et a1.2 Poplar wood Straw Groeneveld l 8 Wood char
N" order kinetics n E(kJ/mol) k, 0.57+ 0.03 2372 0.4 2.62~10*+5.106s~'bar~n 21 1+ 6.1 1.71.107+l.107s%ar-" 0.5 1k0.05 1.oo 138 1.79.103s-latrn-' 198 1.23.107s"atm'" 0.75 104.5+8 210210 230 0.56 27 1 262 196,217 226 -0.5 151 4.77.107 %/min 119 1 . 7 6 ~ 1 0%/min ~ -0.5 182 1.2.lo8min" 182 5.9.107 min-' 217 lo6- lO'~-lm~.'moI-~.~ 0.7
R, = 0.9919, R, = 0.9784 +
From the above comparison one can see that the activation energy varies between 105 and 270 kJ/mol. Most values for E lie between 180 and 270 kJ/mol and the parameters obtained in this investigation are well within this range. The reaction order obtained is 39
also similar to the values found in literature, eventually among the lower values. These data will be further discussed in Fig. 7.
H2 INHIBITION EFFECT The experiments show that the presence of hydrogen inhibits the steam gasification reaction, as presented in Fig. 5. 1.OE-02
I
1.OE-03
3-
Ch
z 1.OE-04
700
750
800
850
900
950
1000
TW)
Fig. 5 Inhibition effect of Hz as a function of temperature and H2 partial pressure. (PHzo=0.1 bar, filled symbols: beech, hollow symbols: birch)
The equations 6 and 6.1 have been used to model the reaction. Table 5 shows the kinetic parameters obtained in this investigation. In spite of the high uncertainty of the model parameter calculation, the model fits well the experimental results (See Fig. 6). Table 5 Kinetic constants for HzO/Hz gasification of birch and beech char. Wood species Beech Birch
Elf
lblf
Elb
(kJ/mol) 199 214
(i'bar-l) 2.0.10' 7.6.107
(kJ/mol) 146 284
kolh
(i'bar-') 1.8*106 2.1-10'2
E3
(kJ/mol) 225 273
ko3
(s.')
8.4.107 1.6.10''
Table 6 compares these results with the few kinetic parameters found in the literature. Although there is a certain agreement in the value of E3, the other values are somewhat different. This could be explained by the high uncertainty of the calculation, as also mentioned by Hansen et al.m or by the differences in char origin.
40
1 .OE-02
-
2l.OE-03 P
3
-a -m
E l .OE-04
1 .OE-05 1 .OE-05
1 .OE-04
1 .OE-03
1 .OE-02
R measured (s-1)
Fig. 6 Calculated reactivity values versus experimental values.
Table 6 Comparison of kinetic parameters for HzO/H2 gasification.
‘Recalculation from Hansen et a1.” experiments.
Fig. 7 compares all the kinetic parameters obtained with those found in literature, with and without hydrogen inhibition, by means of a kinetic compensation diagram. A solid line has been drawn for each of the Langmuir-Hinshelwood constant. Most of the values lie within the same line what might be a sign of consistency in spite of the disparity in activation energies. The kinetic parameters according to n* order kinetics are somewhat more scattered although still aligned. The differences between kinetic parameters can be also due to parameters not studied in this investigation like the number of active sites or the effect of temperature on the active sites behaviour. The three sets of kinetic parameters for birch (Langmuir-Hinshelwood kinetics) represent three valid numerical solutions in the model fitting.
41
\
barley (41) 5.OE+03
1.OE+04
1.SE+04
wood (12)
2.OE+04
2.5€+04
3.OE+04
3.5E+04
4.OE+04
WR (K')
Fig. 7 Kinetic com ensation diagram for HzO and H20/Hz gasification. ti! (*:n order, +: klf, A:klb, B: k3, tw: this work)
EFFECT OF FUEL TYPE With respect to reactivity, the results have shown no large differences between birch and beech. There are however certain differences regarding the effect of temperature on the reactivity profile, and the shape of the profile itself. Fig. 8 shows several reactivity profiles, normalised with respect to their reactivity at 20% conversion to allow comparison. The final increase in gasification rate is more drastic for beech than for birch, especially noticeable for beech at lower temperatures. Moilanen and his c o - ~ o r k e r s ' ~also * ~ ' obtain increasing reactivity profiles with conversion, except for peat. They expect such increasing reactivity because of pore development structure, enhanced by the catalytic effect of the ash, since the ratio catalystkarbon increases with char conversion. Stoltze et a1.I6 obtain similar profiles with barley straw. Rensfelt et a1.' find as well increasing reactivity with conversion, and a characteristic shape of the reactivity profile for each fuel, having each fuel the same curve independent of temperature. However, for washed barley chars, Sgrensen et aL4' find a decreasing reactivity as a function of conversion. The ash analysis presented in Table 2 shows very similar values for the potassium content of both woods, but there is some variation regarding other ash components. It cannot be known from the experiments whether the differences in the reactivity profiles are due to these other ash components or to a different porosity evolution as the conversion proceeds.
42
16.5
x 5
2
14.5
-
12.5
-
/ 750°C
I
8.5 6.5 4.5 2.5 0.5 0.1
0.2
0.3
0.4
0.5 X
0.6
0.7
0.8
0.9
Fig. 8 Reactivity profiles for H20 experiments. (0:beech experiments).
INFLUENCE OF REACTIVITY DEFINITION In this section, six different procedures are used to select a representative reactivity value (rc) from the same experiments, using the reactivity definition (eqn. 8). All the definitions are explained in Table 7. Table 7 Representative reactivity definitions compared in this section.
The kinetic parameters for the n* order kinetic model have been obtained using these definitions of reactivity for the pure steam gasification experiments of birch. All the activation energies lie between 228-238 kJ/mol and the reaction orders between 0.54 and 0.58, apart from definition 3. The frequency factors are somewhat more scattered, lying between 5.10' and 3.10'. Regarding the uncertainty of the calculation, definitions 2, 5 and 4 seem to give more precise results and it is interesting to notice that the error of the reaction order calculation does not depend on how a representative reactivity value is defined. It is very important to analyse the influence of the reactivity definition (eqn. 8 and 9) on the kinetic parameters. Since all representative reactivity definitions are related to a fixed degree of conversion (or a fixed interval), the difference between r and rw will be a multiplying factor, independent of temperature and pressure, and therefore absorbed in the frequency factor. This means that whether equation 8 or 9 is used, the activation energy and the reaction order calculation will give the same result.
43
CONCLUSIONS (1)
(2) (3) (4) (5)
The kinetic parameters according to the n" order reaction model for steam gasification of wood char are E= 237 kJ/mol, b= 2.62.108 and n= 0.57 for birch, E= 211 kJ/mol, k,,= 1.71~10~ and n= 0.51 for beech char. Hydrogen inhibits the steam gasification reaction. The char gasification reaction with steam and hydrogen can be modelled based on Langmuir-Hinshelwood kinetics. The model fits well the results. The type of wood affects very little the kinetic parameters but shows some influence on the reactivity profile. The definition of the reactivity will not affect the activation energy or the reaction order calculation. The method to select a representative reactivity value from one experiment has more influence on the frequency factor than on the activation energy and reaction order. The accuracy of the calculation might also be affected.
ACKNOWLEDGMENTS The work was supported by the Norwegian Research Council, the Danish Ministry of Energy and EK Energy A.m.b.A. The authors want to thank Torben D. Pedersen, Peter Mork and Torben Lyngbech for their help during the experimental work.
REFERENCES 1. Gil, J., Aznar, M.P., Caballero, M.A., Frances, E. & Corella, J. (1997). Biomass gasification in fluidized bed at pilot scale with steam-oxygen mixtures. Product distribution for very different operating conditions, Energy and Fuels, Vol. 11, pp. 1109-1118. 2. Rensfelt, E., Blomkuist, G., Ekstrom, S . , Espenas, B.G. & Liinanlu, L. (1978). Basic gasification studies for development of biomass medium-btu gasification processes, Energy from Biomass and Wastes, IGT, 14-18 August 1978. Paper No. 27, pp. 466-494. 3. Bilbao, R., Garcia, L., Salvador, M.L. & Arauzo, J. (1998).Steam gasification of biomass in a fluidized bed, effect of a Ni-A1 catalyst, Biomass for Energy and Industry. 10" European Conference and Technology Exhibition, 8-1 1 June 1998, Wiirzburg, Germany, pp. 1708-1711. 4. Rapagnli, S., Jand, N. & Foscolo, P.U. (1998). Utilisation of suitable catalysts for the gasification of biomasses, Biomass for Energy and Industry. 10" European Conference and Technology Exhibition, 8-1 1 June 1998, Wiirzburg, Germany, pp. 1720-1723. 5. Evans, R., Knight, R. A,, Onischak, M. & Babu, S.P. (1987). Process performance and environmental assessment of the Renugas process, Energy from Biomass and Wastes X, IGT, pp. 677-696. 6. Paisley, M.A. et al. (1999). Commercial demonstration of the Battelle/FERCO biomass gasification process: Startup and initial operating experience,Proceedings of the 4" Biomass Conference of the Americas, Oakland, CA,USA, pp.1061-1066. 7. Zschetzsche, A., Hofbauer, H. & Schmidt, A. (1998). Biomass gasification in an internal circulating fluidized bed, Proceedings of the 8" EC on Biomass for Agriculture and Industry, Vol. 3, pp. 1771-1777.
44
8. Richard, J.R., Cathonnet, M. & Rouan, J.P. (1982). Gasification of charcoal: Influence of water vapor, Fundamentals of Thermochemical Biomass Conversion, pp. 589-599. Elsevier Applied Science Publishers. 9. Li, J. & van Heiningen, A.R.P. (1991). Kinetics of gasification of black liquor char by steam,Industrial & Engineering Chemistry Research, Vol. 30, No, 7, pp. 15941601. 10. Timpe, R.C. & Hauserman, W.B. (1993).The catalytic gasification of hybrid poplar and common cattail plant chars, Energy from Biomass and Wastes XVI, Institute of Gas Technology, March 2-6, 1992, pp. 903-919. 11. Kojima, T., Assavadakorn, P. & Furusawa, T. (1993).Measurement and evaluation of gasification kinetics of sawdust char with steam in an experimental fluidized bed, Fuel Processing Technology, Vol. 36, pp. 201-207. 12. Capart, R. & GClus, M. (1988). A volumetric mathematical model for steam gasification of wood char at atmospheric pressure, Energy from Biomass 4. Proceedings of the 3rd contractors'meeting,Paestum, 25-27 May, pp. 580-583. 13. Hemati, M. & Laguerie, C. (1988). Determination of the kinetics of the wood sawdust Steam-gasificationof charcoal in a thermobalance, Entropie, No. 142, pp. 29-40. 14. Moilanen, A,, Saviharju, K. & Harju, T. (1993). Steam gasification reactivities of various fuel chars, Advances in Thermochemical Biomass Conversion, Blackie Academic & Professional, 1993, pp. 131-141. 15. Moilanen, A. & Saviharju, K. (1997). Gasification reactivities of biomass fuels in pressurised conditions and product gas mixtures, Developments in Thermochemical Biomass Conversion, Blackie Academic & Professional, 1997, pp 828-837. 16. Stoltze, S., Henriksen, U., Lyngbech, T. & Christensen, 0. (1993). Gasification of straw in a large-sample TGA, Nordic Seminar on Solid Fuel Reactivity, Chalmers University of Technology, Gothenburg, Sweden, 24 November 1993. 17. Stoltze, S., Henriksen, U., Lyngbech, T. & Christensen, 0. (1994). Gasification of straw in a large-sampleTGA, Part I1 Nordic Seminar on Biomass Gasification and Combustion, NTH, Trondheim, 21 June 1994. 18. Groeneveld, M.J. (1980). The co-current moving bed gasifier, Ph.D.thesis, Twente University of Technology, Enschede, Netherlands. 19. Whitty, K.J. (1997). Pyrolysis and gasification behaviour of black liquor under pressurized conditions, Academic Dissertation, Report 97-3, Abo Akademi, Department of Chemical Engineering. 20. Hansen, L.K., Rathmann, O., Olsen, A. & Poulsen, K.( 1997). Steam gasification of wheat straw, barley straw, willow and giganteus, Risgi National Laboratory, Optics and Fluid Dynamics Department, Project No. ENS-1323/95-0010 21. Moilanen, A. & Miihlen, H.J. (1996). Characterization of gasification reactivity of peat char in pressurized conditions. Effect of product gas inhibition and inorganic material, Fuel, Vol. 75, No. 11, pp. 1279-1285. 22. Whitty, K.J. (1993). Gasification of black liquor char with H20 under pressurized conditions, Report 93-4, Department of Chemical Engineering, Combustion Chemistry Research Group. 23. Barrio, M.& Hustad, J.E. C02 gasification of birch char and the effect of CO inhibition on the calculation of chemical kinetics, This conference. 24. Muhlen, H-J., van Heek, K.H. & Jiintgen, H. (1985). Kinetic studies of steam gasification of char in the presence of H2, C02 and CO, Fuel, Vo1.64, July, pp. 944-949. 45
25. Meijer, R., Kapteijn, F. & Moulijn, J.A.( 1994).Kinetics of the alkali-carbonate catalysed gasification of carbon: HzO gasification,Fuel,Vol.73,No.5, pp.723-730. 26. Hiittinger, K.J. & Merdes, W.F. (1992). The carbon-steam reaction at elevated pressure: formations of product gases and hydrogen inhibitions, Carbon, Vol. 30, NO. 6, pp. 883-894. 27. Weeda, M., Abcouwer, H.H., Kapteijn, F. & Moulijn, J.A. (1993).Steam gasification kinetics and burn-off behaviour for a bituminous coal derived char in the presence of Hz, Fuel Processing Technology, Vol. 36, pp. 235-242. 28. Linares-Solano, A., Mahajan, O.P. & Walker, P.L. (1979). Reactivity of heattreated coals in steam, Fuel, Vol. 58, May, pp. 327-332. 29. Liliedahl, T. & Sjdstrom, K. (1997). Modelling of char-gas reaction kinetics, Fuel, Vol. 76, NO. 1, pp. 29-37. 30. DeGroot, W.F. & Shafizadeh, F.(1984). Kinetics of gasification of Douglas Fir and cottonweed chars by carbon dioxide, Fuel, Vol. 63, February, pp. 210-216. 31. Kannan, M.P. & Richards, G.N.(1990). Gasification of biomass chars in carbon dioxide: dependence of gasification rate on the indigenous metal content. Fuel, Vol. 69, June, pp. 747-753. 32. Illerup, J.B.& Rathmann, O.( 1995).C02 gasification of Wheat straw, barley straw, willow and giganteous, Department of Combustion Research, RIS0 National Laboratory, 12th December. 33. Espenas, B.G. (1993). Reactivity of biomass and peat chars formed and gasified at different conditions, Advances in Thermochemical Biomass Conversion, Blackie Academic & Professional, 1993, pp. 142-159. 34. Moilanen, A. & Kurkela, E. (1995). Gasification reactivities of solid biomass fuels, Preprints of papers, American Chemical Society, Division of Fuel Chemistry, Vol. 40(3), pp. 688-693. 35. Sgrensen, L.H.(1994). Fuel reactivity as a function of temperature, pressure and conversion, Ph.D. thesis, Risa National Laboratory, Denmark. 36. Chen. G., Yu,Q. & SjBstrGm, K.( 1997). Reactivity of char from pyrolysis of birch wood, Journal of Analytical and Applied Pyrolysis, Vol. 40-41, pp. 491-499. 37. Zanzi, R., SjiSstrom, K. & Bjombom, E. (1995). Rapid pyrolysis of agricultural residues at high temperature, Proceedings of the 2ndBiomass Conference of he Americas: Energy, Environment, Agriculture, pp. 630-636. 38. Bandyopadhyay, D., Chakraborti, N. & Ghosh, A.(1991). Heat and mass transfer limitations in gasification of carbon by carbon dioxide, Steel Research, Vol. 62, NO.4, pp. 143-151. 39. Sgrensen, L.H., Gjernes, E., Jessen, T. & Fjellerup, J. (1996). Determination of reactivity parameters of model carbons, cokes and flame-chars, Fuel, Vol. 75, No. 1, pp. 31-38. 40. Rathmann, O., Stoholm, P. & Kirkegaard, M. (1995). The pressurized thermogravimetric analyzer at the Department of Combustion Research,, RisB: Technical description of the instrument, Roskilde: Rise National Laboratory, Risg-R-823(EN), Denmark. 41. Sgrensen, L.H. et al. (1997). Straw - H20 gasification kinetics. Determination and discussion, Nordic Seminar on Thermochemical Conversion of Solid Fuels, 3d December, 1997, Chalmers University of Technology, Sweden. 42. Risnes, H., Sgrensen, L.H. & Hustad, J.E. C02 reactivity of char from Danish wheat, Norwegian spruce and Longyear coke. This conference.
46
COZ gasification of birch char and the effect of CO inhibition on the calculation of chemical kinetics M. Barrio, J.E. Hustad Norwegian University of Science and Technology, Department of Thermal Energy and Hydro Power, 7491 Trondheim, Norway
ABSTRACT: Reactivity experiments have been performed in a TGA in the temperature range from 750 "C to 950 "C in steps of 50 "C. The C02 partial pressure has been 0.05,0.1,0.2,0.5 and 1.0 bar. Reactivity profiles have been obtained as a function of conversion for all temperatures and partial pressures and kinetic expressions have been calculated. An important feature of this investigation is that it compares the chemical kinetics according to two models: n* order kinetics and Langmuir-Hinshelwood kinetics. The n* order model is widely used and allows comparison among researchers since more results are available. The activation energy obtained according to #order kinetics is 215 W/mol, the frequency factor 3.1.106and the reaction order 0.38. Since the n* order model does not include the effect of CO inhibition, the parameters for Langmuir-Hinshelwood kinetics have also been obtained. In spite of the large discrepancies for these kinetic parameters among researchers, the model fits very well the experimental data presented here. The ratio CO/CO;! appears to be a relevant factor for reactivity. INTRODUCTION Among the thermochemical conversion processes that take place in a gasifier, the gasification reducing reactions are the limiting step. The reactions of C02 and H2O with the char to produce CO and H2 are considerably slower than the drying, pyrolysis or combustion reactions. A deep knowledge of the chemical reactions involved, as well as the heat and mass transfer mechanisms, would allow an effective enhancement of the gasification process. The residence time and temperature requirements for complete reaction are crucial factors regarding reactor design's2. On the other hand, chemical ~*~. information is required as part of any gasification m ~ d e l ~ *Thermogravimetric analysis allows isolation of the chemical information. The inhibition effect of CO is widely accepted and reasonably well documented for coal char gasification but not for biomass. The lack of extensive literature for wood char C02gasification kinetics including CO has strongly motivated this investigation. This paper shows the results of birch char gasification experiments with C 0 2 and CO. Kinetics for both n* order model and Langmuir-Hinshelwood kinetic model have been obtained. The evolution of reactivity with degree of conversion is also studied, as well as the relevance of the ratio CO/CO2. A parallel investigation has been conducted with the same char but gasified in H20/H2 mixtures6.Future work is planned to combine these investigations.
47
REACTION MECHANISMS It is widely a c ~ e p t e d ~ ~ *that * ~ ~the ' ~ *char ~ ' gasification reaction with CO2 can be represented by the following reaction path:
where C'represents an available active site and C(0) an occupied site" also called a carbon-oxygen complex" or a transitional surface oxide7. The inhibiting effect of CO consists of lowering the steady-state concentration of C(0) complexes by the backwards reaction lb". The evolution of the carbon sites and porosity during the char gasification reaction has been subject of several s t ~ d i e s ~ * ' ~Plante ~ ' ~ *and ' ~ .his co-workers7present a summary of the different models used to take into account the changes in reactivity due to the degree of conversion: changes in the carbon structure as the reaction proceeds and variations in the mineral content. In this investigation, as suggested by Sprensen's work16, the variation of reactivity as a function of the degree of conversion is represented by a function f ( X ) , called the structural profile.
With such definition, the chemical kinetics, r,, do not depend on the degree of conversion. The structural function value for 50% conversion is assumed to be 1. Using Langmuir-Hinshelwood kinetics and the steady state assumption for C(O), the reaction rate will be defined as:
Although some researchers do include dual-site absorption", most investigations prefer the above and a further simplification to n' order kinetics is often ad~pted~~'*''~: rc = k pcoz" Ergun" reduces however the reaction rate to a different expression in order to show the importance of the ratio CO/CO2:
48
k,
Cerfontain et a1." also show that the gasification rate for activated carbon only depends on the ratio PcD/pc02 and not on PCo2or Pco.
EXPERIMENTAL SECTION The wood used in this investigation is Norwegian birch. The proximate and ultimate analysis of the wood are presented in Table 1. The wood sample was cut in cubes (lOx1Ox1Omm) and then pyrolysed in a Macro-TGA at 24 "C/min, up to a temperature of 600 "C. The sample was held at 600 "C during 30 minutes and then cooled naturally. The gasification experiments include a further pyrolysis up to 1000 "C with a heating rate of 30 "C/min. The pyrolysis temperature and heating rate were chosen equal to those used in a previous work for steam gasification of the same char, for the sake of coherence. De Groot et al." also heat the char to 1000 "C to drive off any adsorbed species and desorb any surface oxidation products.
Table 1 Proximate and ultimate analysis of birch. Proximate analysis
Moisture
Fixed carbon
Ash
15.76% C
Volatile matter* 93.3%, mf H
Birch wood Ultimate analysis
6.3%, mf N
48.7
6.4
0.078
0.37%, mf 0 (by diff.) 44.45
Raw wood (%, mf)
* Pyrolysis conditions: Heating at 24 "Umin until 600 "C, held for 30 min, natural cooling. More than 50 experiments have been conducted, all of them isothermal and under a total pressure of 1 bar. Since the reactivity not only varies with reactant partial pressure and temperature but also with the degree of conversion (X), it was found more satisfactory to avoid the variation of two parameters simultaneously. Several researches have modelled the dependence of the reactivity on the degree of conversion with a function called structural profile13,or reactivity factor". It is then possible to run nonisothermal experiments by including the dependency of reactivity on X in the calculations. Also Narayan and Antal" contribute to this discussion regarding nonisothermal experiments. During an endothermic reaction (biomass pyrolysis, in their case) the sample will hold a constant temperature due to the heat demand of the reaction and therefore experience a thermal lag. This will lead to an underestimation of the activation energy and the frequency factor. Once pyrolysed, the char was sieved to 32-45 pm. This particle size was found to avoid heat transfer limitations. Based on previous research on the same equipmentz3,a sample of about 5 mg and a gas flow of 200 mYmin during gasification were used for all experiments. The char sample was first dried in Nz (99,999%) at 110 "C for 30 min, then heated up at 30 "Chin to 1000 "C, kept at 1000 "C for 20 minutes and then brought to the gasification temperature of the experiment. After 10 minutes for 49
stabilisation of the temperature, the gas was switched to the gasification gas: a defined mixture of N2 (99,999%), C02 (99,2%) and CO (995%). The gasification reaction continues until complete char conversion. The thermogravimetric analyser is a SDT-DTA from TA Instruments, supported by an H p PC and software for control and data handling. The system consists of a dual beam horizontal balance. Each arm holds one cup and there is one thermocouple under and in contact with each cup. One cup contains the char sample and the other cup is empty, used as a reference for temperature effects. Detailed description of the instrument can be found somewhere else23.Ceramic cups were used for most of the experiments. The apparatus has been recently upgraded and it was possible to operate in a TGA-DSC mode. Therefore, not only the temperature and the weight have been registered but also the heat demand of the process. Table 2 and Table 3 show the experimental matrix for this work. Table 2 Experimental matrix for C02 gasification experiments
I
I Pco2(bar) 0.05 0.2
0.5
750
I
Tf "C!) -, -I
800
I
850
I
I 900 1
950
.. ... ... ... a .
4
Table 3 Experimental matrix for C02 gasification with CO inhibition
0.2
. O
It was observed that the mass loss rate during pyrolysis at 1000 "C did not decrease with time, but continued stable and, given enough time, consumed totally the char sample. This phenomenon is not uncommon. De Groot et al." observed some pyrolytic gasification during the preheating of the chars under flowing nitrogen. They attribute it to the gasification of the oxidised surface species formed during storage and handling of the chars. Also Tancredi et al." encountered a weight loss of about 28% during pyrolysis up to 1000 "C for carbon already pyrolysed at 800 "C. Mackay and Roberts25 suggest that a prolonged exposure of the char to high temperature in N2 produces rearrangement and shrinkage of the char structure, hindering the diffusion of gas reactants and products. Rathmann et aI.l7 experience the same problem and refers to Whitty26,who suggests to use a weakly reducing gas, 1-2% CO in N2while reaching the gasification temperature to avoid this unwanted reaction. After an exhaustive maintenance control of the instrument, the mass loss rate decreased but did not disappear. In order to avoid the reactions that consumed the char during pyrolysis it was found necessary to increase the N2 flow during pyrolysis to 50
500 ml/min. Under these new conditions, the mass loss during rest pyrolysis was about 20% for all the experiments and mostly while reaching the pyrolysis highest temperature, allowing a better reproducibility of the following gasification experiment. The original gas line installation for the TGA was modified in order to operate safely with CO (See Fig. I). The outlet of the TGA, initially open to the room, was connected to a steel tube of 6mm internal diameter for ca. 400 mm, followed by a plastic tube of 10 mm i.d. that conducted the gasses to the suction system. This change in the installation did not seem to affect the results significantly. As a precaution, a CO detector was used during all the C02/C0 experiments. During the experiments with high concentrations of CO, small carbon deposition was observed on both arms and cups of the thermobalance.
Inert gas
Reactive mixture Safety purge
Fig. 1 Schematic drawing of the experimental set-up.
TREATMENT OF THE RESULTS The reactivity is calculated in this work as:
where
r m(t)
mf
reactivity, (s-') mass of char at the time t, (8) mass of char at the end of the gasification reaction, at the gasification temperature, (g)
51
Several authors prefer the expression:
where m, represents the mass of char(g) at the beginning of the gasification reaction, at the gasification temperature. This expression, however, implies a continuous decrease in reactivity as the reaction proceeds. The degree of conversion, X(t), is obtained from:
Defined in this way, X will be 0 right before the gasification starts, and 1 when the gasification reaction is finished. This degree of conversion does not include then any pyrolysis reaction but only gasification and is independent of buoyancy effects and changes in flow. Moilanen et al.27consider for example m, as the initial mass of the experiment, previous to drying and pyrolysis. This implies that the gasification reaction takes place at degrees of conversion of 80-100%. Combining the above expressions for reactivity and degree of conversion one obtains:
r=
1 dX (1-X) dt
and this expression, discretised by backwards differencing, gives the reactivity as a function of the degree of conversion, r(X), i.e. the reactivity profile. The profile is then fitted with a 5" order polynomial function, R(X), and for each experiment, R(X=0.5) is selected as a representative reactivity value.
R(X = 0.5) = R, * f ( X = 0.5) where f ( X = 0.5) = 1
RESULTS AND DISCUSSION KINETIC PARAMETERS FOR COz GASIFICATION Fig. 2 shows the reactivity for all the experiments with C 0 2 and the kinetic models used to fit the results: n" order (continuous line) and Langmuir-Hinshelwood (dashed line). The kinetic model based on Langmuir-Hinshelwood kinetics fits better the results, especially at low temperatures.
52
1 .OE-02
1.OE-03
750°C 1 .OE-04
0 800°C
A 050°C 0 900°C
950°C 1.OE-05 0.01
0.1
1
P COJbar)
Fig. 2 Influence of COzpartial pressure and temperature on char reactivity,
The parameters have been calculated with the statistical program Sigma Plot@and are shown in Table 4, together with results from other references for comparison. The activation energy obtained according to n* order kinetics compares quite well with the other references except Plante et al.' and Illerup and Rathmann2*. Most values lie between 196-250 kJ/mol. The order of reaction comparison shows however larger variation. The reaction order obtained in this investigation is relatively low; only Henrich et al.30and Risnes et aL31 have obtained similar values. Table 4 Comparison of kinetic parameters for COz gasification of char.
Henrich et aL3'
Risnes et
Graphite Municipal w. soot Spruce
220
53
1.3.109min-'
0.36 0.27 0.6 0.36
INHIBITION EFFECT OF CO It is clear from the investigation that the addition of CO inhibits the COz gasification reaction. Two calculation methods to obtain the kinetic constants in the LangmuirHinshelwood based model have been compared: the two-steps calculation and the direct calculation. In the two-steps calculation, klf (klfoand Elf) and k3 (k30and E3) are first calculated from the C 0 2 gasification experiments only, according to Eq. (12), based on Eq. (4):
1 1 -=-Pcoz rc klf
-1
+- 1 k3
and only klb (klboand Elb)are obtained from the C02/C0gasification experiments with linear regression analysis, following the equation below:
The direct calculation obtains the six constants simultaneously from all the experiments, with and without CO. It is worth noting that the second method requires a powerful statistical tool since it is not possible to simplify the equation for linear regression analysis. Furthermore, a considerable number of experiments and reasonable initial values are required to allow a trustful calculation. Table 5 compares the kinetic constants obtained with these two methods and Fig. 3 allows further comparison between the calculation methods. The Arrehnius diagram shows very similar values for klb and k3 but some discrepancies for klf. Both methods have a high level of uncertainty in the calculation of klboand Elb.
Table 5 Kinetic constants for COz/CO gasification of char, this work. Langmuir-Hinshelwood kinetics Method Two-steps calculation Direct calculation
Elf (kJ/mol)
161,
Elb
kolb
E3
(s-lbar-l)
(M/mol)
(s-lbar-')
(kllmol)
161
9.2.104
36.4
1.91.10°
233
2.3.107
165
1.3.16
20.8
3.6.10-'
236
3,23.107
54
163 (s-l)
-2
I
-3 -4 -5 Y
-c
-6
-7
-8 -9
10 7.5 E-04
8.OE-04
8.5E-04
9 .OE-04
9.5 E-04
l/T(K-1)
Fig. 3 Arrehnius diagram of the Langmuir-Hinshelwood kinetic constants.
The fitting of the results is very satisfactory, as shown in Fig. 4 and Fig. 5 for the CO2/CO experiments, in spite of the uncertainty of the kinetic constants calculation -in particular the frequency factor-. This was also observed by Rathmann et al.”. The lines in Fig. 4 and Fig. 5 and represent the reactivity according to Langmuir-Hinshelwood kinetics, using the parameters obtained by direct calculation.
1 .OE -02
1 .OE-03
1 .OE -04
1 .OE -05
0
0.05
0.1
0.15
0.2
0.25
P co (bar)
Fig.4 Reactivity as a function of CO partial pressure for PcoZ=0.5bar.
55
0.3
A 850'C
v 900'C 0
I
1 .OE -03
_.
950'C
t
h
v v)
P
1 .OE -04
1 .OE -05
0
0.05
0.1
0.15
0.2
0.25
0.3
PW(b)
Fig. 5 Reactivity as a function of CO partial pressure for Pco2 = 0.2 bar.
Table 6 compares the results with data from other references where CO inhibition has been studied. Although the differences are 'remarkable, there are some similarities in the values of Elf(between 100 and 165 kJ/mol in most cases). Most authors also agree that E3 is larger than Elfand therefore that the gasification of the solid carbon -reaction (2)- is the limiting step of the reaction mechanism. Finally, the activation energy Elb is very low for several authors, including this work. This is however not the case for Bandyopadhyay and Gosh3'. Since the activation energy cannot be negative, the results could be understood as a zero activation energy, i.e. the backward reaction (1) does not depend on temperature. Table 6 Comparison of kinetic constants for CO2/COgasification of char.
+
Fixed bed reactor.
The ratio CO/C02as a relevant parameter As already referred by several authors'0'12'20,the ratio Pcdpco2 appears to be an important factor for reactivity. This investigation shows the same fact, as reflected in Fig. 6. This figure is very similar to the one shown by Freund" for carbon. 56
1 .OE-02
I
el-1.25 00.4-0.5 A0.2-0.25
00.1 7 8E-04
8 OE-04
8 2E-04
8 4E-04
8 6E-04
8 8E-04
9 OE-04
1/T (K ')
Fig. 6 Influence of the ratio PcdpcoZon char reactivity.
REACTIVITY PROFILES Fig. 7 shows the reactivity profile obtained for some of the experiments with COz, all of them with a C02 partial pressure of 0.2 bar. The reactivity profiles of experiments with CO (0.05 bar), for the same C 0 2 partial pressure, are also shown. To allow comparison, reactivity values have been normalised by the function:
The reactivity increases always as the char is consumed, especially during the last stages of conversion. All experiments have shown similar profiles. Other researcher^'^.^^ have also observed this type of profiles. The slope and shape of the reactivity profile depends to a certain extent on the gasification temperature. This trend has been observed in all the experiments, both with and without CO addition. However, the differences between reactivity profiles at the same temperature are relatively large and therefore do not suggest any further conclusion. Regarding the concentration of reactants, there seems to be no influence on the shape of the reactivity profiles. The reactivity profiles of the experiments with CO do not show a systematic deviation from the experiments without C02 and no conclusions should be extracted to this respect. Finally, the shape of the reactivity profile does not show any correlation with the ratio C02/CO.
57
5
4
2 s55 3 II
U
2
1
"
I
I
-1
0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
X
Fig. 7 Reactivity profiles as a function of degree of conversion, temperature and presence of CO. (+ 750°C + 800°C, V 850°C, 900°C.A 950°C, 8 1000°C, thick line: (2% experiments, thin line: COz/CO experiments).
CONCLUSIONS (1)
(2)
(3) (4)
It has been found that the activation energy is 215 kJ/mol, the frequency factor 3.1.106 bar^.^' and the reaction order 0.38 for the CO, gasification experiments, according to the n* order reaction model. Langmuir-Hinshelwood kinetics give a better fit to the results. CO addition has an inhibition effect on the C 0 2 gasification reaction. The Langmuir-Hinshelwood kinetics model fits well the results. The ratio CO/C02 appears to be a relevant factor for reactivity. The char reactivity increases with the degree of conversion in a very similar manner for all experiments. To a certain extent, the reaction temperature affects the shape of the reactivity profile. Neither the gasification agent composition nor the ratio CO/C02 seem to affect the shape of the reactivity profile.
ACKNOWLEDGEMENTS This investigation has been supported by the Norwegian Research Council. The authors want to thank also Silke Hubner and Lena N. Berg for their contribution to the experimental work.
REFERENCES 1. Kojima, T., Assavadakorn, P. & Furusawa, T. (1993). Measurement and evaluation
of gasification kinetics of sawdust char with steam in an experimental fluidized bed, Fuel Processing Technology, Vol. 36, pp. 201-207.
58
2.
3. 4. 5. 6.
7. 8.
9. 10 11.
12. 13. 14. 15. 16. 17. 18.
Miihlen, H-J., van Heek, K.H. & Jiintgen, H. (1985). Kinetic studies of steam gasification of char in the presence of Hz, COz and CO, Fuel, Vo1.64, July, pp. 944-949. Koss, L.I. & Walker, L.P. (1988). An iterative kinetic model for the updraft gasification of wood, The 8" Miami International Conference on Alternative Energy Sources. Whitty, K.J., Backman, R. & Hupa, M. (1993). Empirical modeling of black liquor char gasification, Abo akademi, Report 93-8. Department of Chemical Eng. Combustion Chemistry Research Group, Finland. Ggbel, B. et al. (1999). Dynamic modelling of the two-stage gasification process, Proceedings of the 4" Biomass Conference of the Americas, Oakland, CA, USA, pp. 1025-1032. Barrio, M., Ggbel, B., Risnes, H., Henriksen, U., Hustad, J.E. & Sgrensen, L.H. Steam gasification of wood char and the effect of hydrogen inhibition on the chemical kinetics, This conference. Plante, P., Roy, C. & Chornet, E. (1988). C 0 2 gasification of wook charcoals derived from vacuum and atmospheric pyrolysis, The Canadian Journal of Chemical Engineering, Vol. 66, April, pp. 307-312. Kubiak, H. & Miihlen, H.J. (1998). Gas and electricity production from waste material and biomass via allothermal gasification, Biomass for Energy and Industry, 10" European Conference and Technology Exhibition, 8-1 1 June 1998, Wiirzburg, Germany, pp. 1693-1695. Hansen, L.K., Rathmann, O., Olsen, A. & Poulsen, K. (1997). Steam gasification of wheat straw, barley straw, willow and giganteus, Risg National Laboratory, Optics and Fluid Dynamics Department, Project No. ENS-1323/95-0010. Cerfontain, M.B., Meijer, R., Kapteijn, F. & Moulijn, J.A.( 1987). Alkali-catalyzed carbon gasification in COIC02 mixtures: An extended model for the oxygen exchange and gasification reaction, Journal of Catalysis, vol. 107, pp. 173- 180. Capart, R., Gelus, M., Lesgourgues, M. & Li, Z. (1989). Study of biomass gasification under pressure, Pyrolysis and Gasification, London (UK), Elsevier Applied Science, pp. 593-597. Freund, H. (1985). Kinetics of carbon gasification, Fuel, Vol. 64, May, pp. 657660. Sgrensen, L.H. (1994). Fuel reactivity as a function of temperature, pressure and conversion, Ph. D. thesis, Risg National Laboratory, Denmark. Della Rocca, P.A., Cerrella, E.G., Bonelli, P.R. & Cukierman, A.L. (1999). Pyrolysis of hardwoods residues: on kinetics and char characterization, Biomass and Bioenergy, Vol. 16, pp. 79-88. Ye, D.P., Agnew, J.B. & Zhang, D.K. (1998). Gasification of a Sough Australian low-rank coal with carbon dioxide and steam: kinetics and reactivity studies, Fuel, Vol. 77, NO. 11, pp. 1209-1219. Sgrensen, L.H., Gjernes, E., Jessen, T. & Fjellerup, J. (1996). Determination of reactivity parameters of model carbons, cokes and flame-chars, Fuel, Vol. 75, No. 1, pp. 31-38. Rathmann, 0. et al. (1995). Combustion and gasification of coal and straw under pressurized conditions. Task 2: Determination of kinetic parameters in PTGA, Roskilde: Risg National Laboratory, Risg-R-819 (EN), Denmark. Groeneveld, M.J. (1980). The co-current moving bed gasifier, Ph.D. thesis, Twente University of Technology, Enschede, Netherlands.
59
19. DeGroot, W.F. & Shafizadeh, F. (1984). Kinetics of gasification of Douglas Fir and cottonweed chars by carbon dioxide, Fuel, Vol. 63, February, pp. 210-216. 20. Ergun, S.(1956). Kinetics of the reaction of carbon dioxide with carbon, Phys. Chem., Vol. 60, pp. 480-485. 21. Gjernes, E. et al.( 1995). EFP91. Theoretical and experimental investigation of coal and biomass combustion and gasification properties at high pressure and temperature. Final report, Roskilde: Risa National Laboratory, Ris@-R-859(EN), Denmark. 22. Narayan, R. & Antal, M.J. (1996). Thermal lag, fusion and the compensation effect during biomass pyrolysis, Industrial & Engineering Chemistry Research, Vol. 35, NO, 5, pp. 1711-1721. 23. Granli, M. (1996). Theoretical and experimental study of the thermal degradation of biomass, Ph.D. thesis, Norwegian University of Science and Technology, NTNU. 24. Tancredi, N., Cordero, T., Rodriguez-Mirasol, J. & Rodriguez, J.J. (1996). C 0 2 gasification of eucalyptus wood chars, Fuel, Vol. 75, No. 13, pp. 1505-1508. 25. Mackay, D.M. & Roberts, P.V. (1982). The influence of pyrolysis conditions on the subsequent gasification of lignocellulosic chars, Carbon, Vol. 20, No. 2, pp. 105-1 1 1 . 26. Whitty, K.J. , Sorvari, V., Backman, R. & Hupa, M. (1993). Pressurized pyrolysis and gasification studies of biomasses, Combustion Chemistry Research Group, Report 93-7, Abo Akademi. 27. Moilanen, A. & Kurkela, E. (1995). Gasification reactivities of solid biomass fuels, Preprints of papers, American Chemical Society, Division of Fuel Chemistry, Vol. 40(3), pp. 688-693. 28. Illerup, J.B. & Rathmann, 0. (1995). C02 gasification of Wheat straw, barley straw, willow and giganteous, Department of Combustion Research, RIS0 National Laboratory, 12" December. 29. Bandyopadhyay, D., Chakraborti, N. & Ghosh, A. (1991). Heat and mass transfer limitations in gasification of carbon by carbon dioxide, Steel research, Vol. 62, No. 4, pp. 143-151. 30. Henrich, E. et al. (1999). Combustion and gasification kinetics of pyrolysis chars from waste and biomass, Journal of Analytical and Applied Pyrolysis, Vol. 49, pp. 221-241. 31. Risnes, H., Sorensen, L.H. & Hustad, J.E. C 0 2 reactivity of char from Danish wheat, Norwegian spruce and Longyear coke. This conference. 32. Bandyopadhyay, D. & Ghosh, A. (1996).Validity of rate equation based on Langmuir-Hinshelwood mechanism for gasification of carbon - a reappraisal, Steel research, Vol. 67, No. 3, pp. 79-86.
60
C 0 2 reactivity of chars from wheat, spruce and coal H. Risnes, L. H. S@rensen*and J. E. Hustad Department of thermal energy and hydropower, The Norwegian University of Science and Technology, Trondheim, Norway * ReaTech c/o Center for Advanced Technology (CAT), Frederiksborgvej 399, Postbox 30, DK-4000Roskilde
ABSTRACT: Measurement and modelling of C02-gasification reactivity is presented for two biomass char species: Danish wheat straw (d,,
61
Char conversion is typically the rate-limiting step during combustion and gasification of solid fuels. The magnitude and variation of char reactivity during gasification are, therefore, of primary concern when comparing results from different solid fuels such as coal and biomass. Although this literature provides valuable information related to specific aspects of char reactivity, there is a need for further discussion of the validity range of the determination and extrapolation of reactivity data for coal and biomass. Experimental data for C02 gasification of chars have been reported using TGA, fixed bed, laminar flow (drop tube), entrained flow or fluidised bed reactors. The mechanism of this reaction is not yet fully understood. It is generally believed that COz adsorption on the surface of char (oxygen transfer) followed by CO desorption (carbon removal) are the main steps determining the apparent rate3-6.This mechanism can be further simplified by proposing the Langmuir-Hinshelwoodtype rate expression'. For the following discussion a simple nth order reaction is sufficient. The reviewed literature reports both fractional values of n '-15 and alternative reaction scheme^^^'^"^ (e.g. 1" order and Langmuir-Hinshelwood type expressions). Most of the experimental studies4.8.9,16-23.31.32 are directed towards a acked-bed gasification, with very limited work on single particle behavi~ur''~~'~~'~~~~~~~~. Kinetic parameters from selected publication on C 0 2 reactivity are shown in Table 1 . A wide span of ap arent activation energies is published: for biomass chars, 80.3 kJ/mols -318 kJ/mol and for coal chars, 79 kJ/mol 38 -359.5 kJ/mol 39. Some of the discrepancy can be explained as due to different experimental procedures (such as sample load, particle size and sample preparation) and the application of different analysing equipment. Concerning the latter one, the potential role of systematic errors in temperature measurements among various thermobalances is evident4'. Variations might also be explained by different extraction procedures, lack of accuracy caused by the approximations used in the different computational methods and the kinetic compensation effect. Many researchers have tried to correlate the kinetic rates of carbon gasification to the many physical and chemical properties of the different materials used. Despite of this a universal rate expression does not exist. Furthermore it is difficult to find data operational for reactor simulations in the relevant temperature and partial pressure ranges. In particular the variation of the reactivity with conversion due to structural variations is not dealt with, i.e. the structural profile is seldom explicitly given. In the present paper investigation are made in order to produce operational data for the gasification reactivity for biomass chars and coal chars. Furthermore the differences between chars derived from biomass and coal is illustrated. Thermogravimetric analyses have been used to obtain information about the kinetic values of the C-CO, reaction.
B
THEORETICAL For a solid fuel, the char reactivity towards a reactive gas is usually defined 'in terms of the conversion rate per remaining mass7: 1 dm =-- 1 dX R = --m dt l - X dt
where m is the mass of the organic portion of the sample, d d d t is the conversion rate and X is the degree of conversion, X = (m-mo)/mo.The reactivity is measured versus 62
Table 1 Summary of kinetic parameters available in the literature. Experiments conducted under conditions (fixed bed of particles, particle size, temperature and pressure ranges etc.) comparable to this work. Char origin ln(k0) E, n pc02 dp 4 [Urnin] [kJ/mol] [-] [atm] [pm] [mg] Wheat straw* 19.67 205.6 0.59 -1 el50 5 Spruce* 20.96 219.9 0.36 -1 <63 5 Longyearcoke* 18.66 233.1 0.51 -1 <45 10 19.04 215 0.38 -1 Birch'' 32-45 5 Eukalyptus 230-233 1 45-53 10 haft lignin16 *** Coalz8 22.73 189 1 105-74 5 (pp) 22.96 212 coat3 16.45 212.9 0.89 0.21-1 75-125 2 213-251 - * * * loo****
I***
conversion and may in order to obtain kinetics be sought split into a chemical hnetics term (rJ, and a structural profile 0). The structural profile is often implicitly or explicitly with success assumed to summarise the effect of available internal surface, available activeheactive sites and pore-evolution and to be invariant over the relevant T-Pd~rnain~"~'. For COI gasification one then finds R = rc(T,Pc02)-fix). To illustrate that the invariant structural profile assumption is not always valid one may instead write
Where cf is a structural parameter that counts available reactive carbon sites and c, is a coefficient that account for distribution of reactive carbon sites types and catalytic effects and thus a variation in c, may change kinetic parameters. The gas composition vector (P)inflX) may beside the COz partial pressure also include such gas partial, pressures as KOH since the likelihood that a catalwc site is activated is a function of the partial pressure of the catalyst, the site-catalyst attraction forces and the temperature. Since from Eq. (2) the structural profile invariance 'SPI' assumption is by no means obvious we suggest that a temperature and partial pressure range are always given for the validity of the structural profile invariance assumption. Only if the invariant structural profile assumption is approximately valid a reference profile (Rrer> can be used to eliminate the structural profile to form a normalised reactivity (R,) and to determine kinetics up to a constant
63
The proposed Langmuir-Hinshelwood kinetics43expression for the gasification reaction (Cot, N2 and CO gas mixture) is given by’: rc =
kl f pc02 1+ upcoz + bPco
(4)
This more complex rate expression can describe both, saturation by the reactant gas and inhibition by the product gas. However, for to obtain data in the present work no external CO was added to the reactant gas and the simpler nth order reaction scheme was applicable rc = k P&,
, k = k, exp(-E I RT)
(5)
Both E and n were calculated using the Marquart-Levenberg least square algorithm. The sum SS = Z w (R,t,-R,)2 was explicitly used to minimise effects of non-uniform errors. By applying the weight function w = 1/RZap all data points are considered equally significant.
Char reactivity versus conversion Comparing rates at a given conversion (eg. initial rate) is frequently done in kinetic studies. In most cases the initial 10% or 50% of the conversion is used for evaluating the kinetic parameters. For most chars of coal, lignite and peat, the reactivity decreases with increasing conversion or time, whereas for most chars of biomass it increases. Reactivity can also exhibit a maximum or a minimum. Tancredi et al.m and RodriguezMirasol et explains the steep reactivity increase at high conversion levels, often seen for biomass chars, as a result of an increasing catalytic effect of the metallic constituents of the inorganic matter present in the char. McKee et al.45 found the activity of alkali and alkaline earth metal salts to decrease in the order Cs>K>Nu>Li>>Bu>Sc=Cu,in catalytic gasification. Alkali metals being approximately ten times more active than alkaline earth metals. (Also observed for a lignite char and graphite). Moilanet~~~ correlated gasification rates and ash composition, and found that especially the rates at higher fuel conversions seemed to decrease with increasing silica content in the fuel. This indicates that catalytic active ash components can loose their activity due to reactions with silica, or due to sintering behaviour. To derive the kinetics for a char or coke, with respect to a specific reactant gas, the normalised reactivity R,, must be constant and hence the structural profile AX,P,T,c,) invariant at all the relevant experimental conditions.
64
Defining f ( X )
In the following the structural function is simplified and described as a function of X only (i.e. Ax)).Ax) is calculated from the experimental values in the range 0.1S10.9 and extrapolated for Xc0.1 and X>0.9. The structural rofile is modelled through a polynomial Ax) = A O + A , X + A ~ X ' + A ~ X ~ + A & + A & ~ + AR,~ is defined as the average value of the normalised profile times the reactivity of the reference profile at 50% conversion. Thus the reactivity is modelled as:
, whereAX=OS) is defined to be 1.0. The uncertainty in measured reactivity is often large for low degrees of conversion (x), before the particles pore structure has opened up and experimental conditions have settled. The uncertainty in the measured reactivity will also increase as m goes to zero. Kinetics determinations are therefore limited to values of the normalised reactivity (R,) within the conversion interval from X = 0.2 to 0.8.
EXPERIMENTAL SECTION Reactivity profiles were measured vs. conversion over a broad range of temperatures and carbon dioxide partial pressures for all three chars. A reference reactivity profile (Rref) was chosen and used to eliminate the structural profile in the relevant experiments. Kinetic parameters are calculated by using both the normalised reactivity (R,) and the reactivity at 20, 40, 60 and 80% conversion (RXZO,R x ~ &RX60and Rx8o). Calculated reactivities are compared.
INSTRUMENTAL In the atmospheric experiments an "SDT 2960, TA-instruments simultaneous differential and thermal gravimetric analyser" were used to determine the COzgasification reactivity under isothermal conditions. The samples were kept isothermal for 10 minutes in N2 at 150"C, before the sample temperature was increased (heating rate 30"C/min) to the defined reaction temperature and switching to the N2/C02 mixture. The gas flow used was 400 ml/min. Ash was determined by increasing the temperature to 1000°C for complete conversion.
SAMPLES The investigations are made on two biomass chars derived from wheat straw and spruce, and a dense metallurgical coke derived from Longyear coal. Sample loading was around 4 mg for the biomass samples and 10 mg for the Longyear coke. Wheat samples were prepared from a batch of wheat straw bales. The straw was grounded to a particle size < 150 pm and pyrolysed in a pressurised entrained flow reactor (2-3s at 900°C and 15 bar N;'). The char derived from spruce (d, c 63 pm) was produced at atmospheric conditions. The Longyear coke (d, c 45 pm) was prepared from cooking a high volatile coal at approximately 1000°C for several In all cases nitrogen was used for the devolatilisation. Kinetics are determined based on gasification experiments made at atmospheric conditions, and C02-partial pressures equal to 0.03, 0.20, 0.50 and 1.0 bar. The 65
gasification temperature varied between 700 and 1O0OoC,in steps of 50 degrees. For each char the reference reactivity profile was measured over a broad conversion interval.
RESULTS EXPERIMENTAL EVALUATION OF GASIFICATION REACTIVITY Isothermal experiments Reactivity profiles obtained from all three chars are evaluated by using the normalisation procedure. The quality of this procedure is indicated by the calculated standard deviation (aj relative to the reference profile: a ) Wheat straw: 11 out of 14 experiments shows CT<5.8%, b) Spruce: 9 out of 12 shows Q< 5.9% and c) Longyear coke: experiments below 1000°C shows an extremely good fit, C T < 2.4%. At 1OOO"C AX) changes dramatically for Longyear coke, as the CO2 partial pressure is reduced from 1.0to 0.03bar. (The calculated standard deviations are based on R, (0.21X 10.8) and a step size of AX = 0.002.) Wheat straw: A total of 14 isothermal and isobaric runs were made at atmospheric conditions. Gasification temperatures were 700, 750, 800 or 850°C and C02-partial pressure were 0.03, 0.20, 0.50 or 1.0 bar. Rathmann et investigated C02gasification reactivity for wheat straw at pressurised conditions and found the effect of the total pressure to be insignificant, at a constant C02 partial pressure. S@rensen3' obtained the same conclusion from combustion (02)reactivity studies. However, external addition of carbon monoxide, resulted in decreased gasification reactivity and a change in the reactivity profile. In the present study the same char sample as used by Rathmann4' was investigated at atmospheric conditions. Spruce: Kinetic parameters for a Norwegian spruce char are calculated based on 12 isothermal experiments. No marked effect of CO-inhibition or C02-saturation seems to be present between 750-850°Cand 0.03-1.O bar C02. Longyear coke: Because of the dramatically change of reactivity profile at 1000°C. experimental data obtained at 1000°C and 0.20 and 0.03bar C02 are not used to determine the kinetic parameters but added to the plot in order to show the error that may otherwise be introduced.
Evaluation of the structuralprofde Fig. 1 shows a comparison of the reference profiles normalised to the reactivity at X = 0.2. This reveals a pronounced difference in the development of the overall reactivity as a function of conversion. For the spruce char the overall reactivity increases with a factor of 2.8 between X = 0.2and 0.8,while wheat straw char and Longyear coke gives 1.0and 1.2 respectively. For the Longyear coke the reactivity profile at 1000°C showed an unexpected behaviour. At 1 .O bar COz, AX) approximately equals the reference profile observed at lower temperatures (increasing with conversion, Ro.dRo.2 = 1.2),while the profile at 0.03
66
3.0 1 2.5
14
--
0.00
0.20
0.40
0.60
0.80
1.00
Conversion, X [-I
Fig. 1 Comparison of different reactivity profiles from the analysed chars. bar shows a marked decrease (Ro.$Ro.2 = 0.4). In the former case activation and the latter case deactivation of the char is observed (Fig. 2). The profile obtained at 1OOO"C and 0.2 bar is also shown in Figure 2. The reactivity profiles obtained at 1000°C were validated. The C 0 2 reactivity profile at 0.03 bar C02 and 1000°C is fairly similar to the O2 reactivity profile obtained in air at 600°C by S@rensen4'.This indicates that the same sites may eventually be non-catalytic and be responsible for the pore and site evolution in both these cases. Alternatively a time dependent deactivation may take place for T21000"C.
1.40
1.20
I.Oh
-.
1.00 - ry
6 0.80 -.
3
?5 0.60--
0 Zhs
0,
%
0.40 --
0.004
0.00
I
0.20
0.40
0.60
0.80
1.oo
Conversion, X [-I
Fig. 2 Changing reactivity profile for the Longyear coal, at 1000°C. 67
DISCUSSION The results obtained from the kinetic determination are shown in Table 2 and Fig. 3. Experimental values and the results from nth order kinetics are shown in Fig. 4. Taken into account the wide span of E,, (79 kJ/mol3*-359.5W/mol 39) and n (0-1) reported in the literature the results reported in this study is found to be within a relatively narrow range. Our results are in general agreement with the data shown in Table I (205.6233.1 vs. 189-260 kJ/mol and. 0.36-0.59 vs. 0.38-0.89 respectively). Miwas’, in his review, reported activation energies for coal chars in the range 190-210 kJ/mol. The structural profiles of both biomass chars show consistency within the studied temperature and pressure range. This result is in agreement with earlier findings”. Our reactivity profiles obtained under different reactants (e.g. CO2 and 02) ives reactant specific reactivity profiles. Which is consistent with Salatino13and Floess .
8
Table 2 Reference profile AX), and kinetic parameters for nth order kinetics. Standard deviation is calculated, assuming 95% confidence interval and given in %(k).Experimental data are assumed to be normal distributed. ln(ko)[l/min] E[kJ/moll n [-I JlX) Wheat straw (9 (+) Ao 1.2334E+00 x=0.2 18.21 193.8 3.8 0.52 0.02 A1 -3.7472E+00 X=0.4 19.68 205.9 0.2 0.60 0.02 A2 2.1309E+01 X=0.6 20.03 208.6 0.3 0.62 0.03 A3 -7.2306E+01 20.41 X=0.8 213.1 0.6 0.55 0.04 Aq 1.4679E+02 ~ ~ , ~ ~ ~ ( ~ =19.67 0 . 5 ) 205.6 0.3 0.59 0.02 A5 -1.5311E+02 A6 6.1511E+01 ~~
(+I
Spruce x=0.2 X=0.4 X=0.6 X=0.8 ~,,,(~=0.5)
20.07 21.03 20.65 21.11 20.96
215.3 222.0 215.4 215.6 219.9
0.4 0.3 0.7 0.7 0.4
Longyear x=0.2 X=0.4 X=0.6 X=0.8 ~c.n,,(X=0.5)
18.81 18.98 18.29 18.95 18.66
234.2 236.4 228.8 233.5 233.1
0.3 0.3 8.3 0.3 0.6
0.35 0.37 0.40 0.34 0.36
(+I
68
0.52 0.51 0.51 0.52 0.51
(+) Ao 1.5399E+00 0.02 A1 -1.965OE+Ol 0.02 A2 1.5876E+02 0.04 A3 -5.9680E+02 0.04 Aq 1.1513E+03 0.02 A5 -1.0900E+3 A6 4.0332E+02 (+) Ao 5.7762E-01 0.02 A , 8.9778E+00 0.03 A2 -6.2449E+01 0.03 A3 2.0354E+02 0.02 A4 -3.4601E+02 0.02 A5 2.9850E+02
850°C
800°C
750°C
700°C
E
e
P
t
t 4
I
ow1 0 01
0.1
1
PC02 [bar]
Fig. 3 Reactivity (rc,mm(X=0.5) = rc(T, Pcoz) .flX=O.5)) for a) wheat straw char, b) spruce char and c) longyear coke, presented in a log(R)-log(Pcoz) diagram (A = experiments not included in kinetic analysis). Solid line: reactivity as calculated by the nth order kinetic model.
69
\
le-5 -:
le-6 -
\
\
\
\ Longeyar I
The observed shift in the reactivity profile for the Longyear coke was highly unexpected. A similar effect has not been found in the reviewed literature. Introduction of mass transfer effects would be expected to result in the reverse effect, i.e. increasing char reactivity at higher degrees of conversion. We suggest that the observed effect is caused either by a shift in reactive site distribution, by thermal annealing or a time dependent deactivation of the catalytic sites. Thus, the parameters in cf and c, in Eq. 2 can not be ignored This observed behaviour illustrates that the reactivity profile and thus kinetic data for the Longyear coke cannot be extrapolated to beyond Iz100O"C. A more precise description of this phenomena demands further investigations. A comparison of the reactivity of wheat, spruce and Longyear coke is shown in Fig. 4. The two biomass chars exhibit almost identical reactivity while the reactivity for the much more dense metallurgical coke is clearly lower. This is in general agreement with the reviewed literature.
CONCLUSION For all fuels tested the nth order kinetic models gives a good representation of the thermogravimetrically determined C02-reactivity. Since normalised reactivity was approximately independent of conversion, for each char, one reaction order, and activation energy could be estimated for the whole range of temperature and C02 partial pressures. Applying nth order kinetics the following kinetic parameters were obtained (atmospheric experiments): apparent activation energies ( E ) for wheat straw char; 205.6 [kJ/mol], spruce char; 219.9 [kJ/mol] and Longyear coke; 233.1 [kUmol], the respective reaction orders equals 0.59, 0.36. and 0.51. In the investigated temperature (7") and C02-partial pressure (Pi) range, using normalised reactivity (R,) and the reactivity at 20, 40, 60 and 80% conversion (RXZO,R x ~ oRXM) , and RXSO)for kinetics determination gives similar results, as expected in the case of an invariant structural profile AX). O2 and C 0 2 give reactant specific and different reactivity profiles. For the Longyear coke partial pressure dependent reactivity profiles were
70
measured and discussed. A simple nth order reaction and an invariant structural profile assumption was used under the evaluation of experimental data. (see Eq. 6). However, the results indicated that a more complex mechanism, e.g. Eq. 2, must be considered if the results, even for a single pure gas, are to be extrapolated to their limits for the purposes of modelling reactor data under realistic conditions.
REFERENCES 1. TERES (1994) The European Renewable Energy Study; Prospects for renewable energy in the European Community and Eastern Europe up to 2010 (ALTENER PROGRAM). Main report made for the "Commission of the European Communities, Directorate General for Energy (DGXVII). Brussels, Belgium 2. Bridgewater A. V. (1995) Fuel, 74, p. 631-635 3. Freund, H. (1985) Fuel, 64, p. 657-660 4. Ergun, S. (1956) Phys. Chem., 60, p.480-485 5. Walker Jr., P. L., Rusinko Jr. F. & Austin L.G. (1959) Advan. Catalysis., 11, p. 133 6. Matsui, I., Kunii, D. & Furusawa, T. (1987) Ind. Eng. Chem. Res., 26, p. 91-95 7. Laurendeau, N-M. (1978) Prog. Energy Combust. Sci. 8. Plante, P., Roy, C. & Chornet, E. (1988) The Canadian Journal of Chemical Engineering, 66, p. 307-3 12 9. DeGroot, W. & Shafizadeh F. (1984) Fuel, 63 10. Standish, N. & Tanjung A.F.A. (1988) Fuel 67 11. Groeneveld, M. J. & van Swaaij, W.P.M. (1980) Chem. Eng. Sci. 35, p. 307-313 12. Barrio, M & Hustad, J. E. C02 gasification of birch char and the effect of CO inhibition on the calculation of chemical kinetics. This conference 13. Salatino, P., Senneca, 0. & Masi, S. (1998) Carbon, 36, p. 443-452 14. Hampartsoumian, E., Murdoch, P.L., Pourkashanian, M., Trangmar. D.T. & Williams, A. (1993) Combustion science and technology, 92, p. 105-121 15. Roberts, D.G. & Harris, D.J. (2000) Energy & Fuels 14, p. 483-489 16. Rodriguez-Mirasol,J., Cordero, T. & Rodriguez, J. J. (1993) Carbon, 31, p. 53 17. Dutta, S., Wen, C. Y. & Belt, R. J. (1977) Ind. Eng. Chem., Process Des. Dev., 16, p. 20-30 18. Salles, J. E. F., De Castro, L. F. A. & Tavares, R. P. (1985) Metallurgia-ABM, 41, 15 19. Li, J. & van Heiningen, A R. P. (1990) Ind. Eng. Chem. Res., 29, p. 1176 20. Tancredi, N., Cordero, T., Rodriguez-Mirasol,J. & Rodriguez, J. J. (1996) Fuel, 75, p. 1505-1508 21. DeGroot, W. & Richards, G. (1989) Carbon 27, p. 247-252 22. Turkdogan, E. T. & Vinters, J. V. (1970) Carbon, 8, p. 39 23. Blackwood, J. D. & Ingeme, A. J. (1960) Austral. J. Chem., 11, p.194-209 24. Capart, R., Gelus, M., Lesgourgues, M. & Li, Z. (1989) Pyrolysis and gasification, London (UK) Elsevier Applied Science, p. 593-597 25. Dasappa, S., Paul, P. J., Mukunda, H. S . & Shrinivasa, U. (1994) Chem. Eng. Sci., 49, p. 223-232 26. Whitty, K. J. (1997) Ph.D. dissertation, Abo Academy, Combusiton Research Group Report, R97-3 27. Zamalloa, M., MA, D. & Utigard, T.A. (1995) ISIJ International, 35, p. 458-463 28. Kovacik, G., Chambers, A. & Ozum, B. (1991) Canadian Journal of Chemical Engineering, 69, p. 811-815 71
29. Wu, P.-C., Lower, W. E. & Hottel, H. C. (1988) Fuel, 67, p. 205-214 30. Radovic, L. R., Jiang, H. & Lizzio, A. A. (1991) Energy & Fuels, 5, p. 68-74 31. Kasaoka, S., Sakata, Y., Shimada, M. & Matsutomi, T. (1985) Journal of Chem. Eng. of Japan, 18, p. 426-432 32. Mulhen, H. J., Heek, K. F. & Juntgen, H. (1985) Fuel, 64, p. 944-949 33. Mukunda, H. S., Paul, P. J., Srinivasa, U. & Rajan, N. K. S. (1984) In: Proceedings of the 20th International Symposium on Combustion, pp. 1619-1628,the combustion institute 34. Reyes, S. & Jensen, K. F. (1986) Chem. Eng. Sci. 41, p. 333-343 35. Reyes, S. & Jensen, K. F. (1987) Chem. Eng. Sci. 41, p. 345-354 36. Morell, J.I., Amundson, N.R. & Park, S.K. (1990) Chem. Eng. Sci. 45, p. 387-401 37. Calemma,V. & Radovic, L. R. (1991) Fuel, 70, p. 1027 38. Kwon, T.W, Kim, S.D & Fung, D.P.C. (1988) Fuel vol. 67 ,no. 4, p. 530-535 39. Walker Jr., P. L. Rusinko Jr. F. & Austin L.G. (1959) Advan. Catalysis. vol 11, s. 133,1959. 40. Granli, M. G., Antal, J. A, Jr. & VLhegyi, G. (1999) Ing. Eng. Chem. Res., 38 41. Sarensen, L. H. (1996) Ph.D. disstertation, Risa national laboratory, Danmark, Risgr-R-838(rev.)(EN) 42. Sarensen, L. H., Gjernes E., Jessen T. & Fjellerup J. (1996) Fuel, 75, p. 31-38. 43. Rief, A. E. (1952) Journal of Phys. Chem., 56, p. 785-788 44. Rodriguez-Mirasol, J., Cordero, T. & Rodriguez, J. J. (1993) Energy & Fuels 7, p. 133-138 45. McKee, D. W. (1983) Fuel, 62, p. 170-175 46. Moilanen, A. & Kurkela, E. (1995) Preprints of papers, American Chemical Society, Division of Fuel Chemistry, 40, p. 688-693 47. Rathmann, 0. & Illerup J.B. (1996) Dep. of Combustion Research, RIS0 National Laboratory, Denmark, RIS0 R-873 (EN) 48. Hustad, J.E., Aho, M.J., Hupa M., Noopila, T., Sarensen, L.H., Clausen S., Kiarboe, L., Gromulski, J., Bengtsson M. & Leckner B. (1990) La Rivista dei Combustibili, 10,257. 49. Hustad, J. E. (1990) Ph.D. dissertation, Norwegian Institute of Technology 50. Miura, K., Hashimoto, K. & Siverston, P. L. (1989) Fuel, 68, p. 1461-1475 51. Floess, J. K., Longwell, J. P. & Sarofim, A. F. (1988) Energy & Fuels 2, p. 18-26
72
Gasification reactivity of charcoal with C 0 2 at elevated conversion levels R. P. W. J. Struis 0, C. von Scala #, S. Stucki 0 and R. Pins * j Laboratory for Energy and Materials Cycles, Paul Scherrer Institute, 5232 VilligenPSI, Switzerland * Laboratoly for Technical Chemistry, ETH, 8092 Zurich, Switzerland # Now at Sulzer ChemTechAG, Dpt. SRT 0600, Post box 65, 8404 Winterthur, Switzerland $3
ABSTRACT The lunetically controlled charcoal reactivity with C02at 8OOOC can very well be described over the entire conversion range when extending Bhatia and Perlmutter's random pore model derivation with two additional parameters only. With untreated charcoal, the extension addresses mainly non-porous phenomena associated which the gradual disintegration of the particle structure at the higher conversions, but the extended kinetic relations are also well suited to describe reactivity effects dominated by metal catalyst accumulation (or re-activation) in the charcoal with progressing conversion.
INTRODUCTION Gasification of carbonaceous materials, especially coal, is an advanced and developed subject. The relative reactivity of different carbons, or of a carbon which has been modified, depends on surface area, catalysis, and inhibition effects. This study addresses surface effects in untreated, catalytic effects in sodium impregnated, respectively, inhibition effects in chloride treated wood char during gasification with C02 at 800°C in the lunetically controlled regime. Special focus is given to a mechanistically coherent description of the charcoal reactivity over the entire gasification stage in terms of both, structural and catalytic, effects. This study was initiated by our wish to understand and describe the observed negative d u e n c e of toxic waste wood contaminants like heavy metals and chloride, for which little is known in the literature.' It is known that the effect of the surface area in the gasification of charcoal is intimately related to the very broad pore size distribution of this material. Random pore structure models accounting for the effects of pore growth and coalescence have been proposed by various authors'.' and have often shown satisfactory agreement between theory and experiment, but none of the proposed kinetic relations describes the charcoal reactivity in the conversion range beyond X=0.7 satisfactorily. For the latter conversion
73
range we observed that the charcoal reactivity passes through a maximum, whereas the pore models predict a systematic drop to zero. Although this deviating behaviour is ignored by most workers, some authors334 pointed out that it emerges from the sudden overall disintegration of the porous structure. The anticipated behaviour, however, contrasts strongly with the gasification results reported below. For the metal catalysed gasification, several authors report the occurrence of a But, up reactivity maximum around Xa0.7 with sodium or potassium enriched to date, no efforts have been undertaken to extend the lunetic relations provided by state-of-the-art random pore models, to account for the occurrence of the maximum around XzO.7. We believe that this late reactivity maximum results from catalyst accumulation (but not saturation) in the charcoal. Inhibition effects induced by chlorine and reactivation by hydrolysis have been reported in the literature, but mainly from a phenomenological point of view in alkali metal catalysed steam gasification studies.' However, a description of the charcoal reactivity in the presence of chlorine over the entire gasification stage is lacking. This study utilises the capability of acid washing to remove mineral matter from charcoal to separate structurally from catalytically determined contributions to the charcoal reactivity. We present a kinetic relation, whch we derived by extending the pore model derivation of Bhatia and Perlmutter' with two additional parameters to account also for additional effects, such as the gradual creation of new surface area by the particle disintegration process, respectively, catalyst accumulation or re-activation effects in the alkali metal catalysed gasification. The resulting relation is found to describe our gasification results very satisfactorily over the entire conversion range. EXPEFUMENTAL GASIFICA TION
To accomplish gasification it is inevitable to pass through a pyrolysis step. We performed the pyrolysis separately and used the resulting charcoal for the gasification experiments reported below. As charcoal is mostly composed of solid carbon, C(s), we chose CO2 as gasifying agent in the well-known Boudouard reaction: C(s) + COZ + 2 CO. The progress of this reaction was monitored by a thermo-gravimetric analyser (TGA) recording the weight loss of the charcoal during gasification. Moreover, the charcoal was characterised with respect to its initial reactive surface area and metal content. Pre-pyrolysis Wood pieces (ca. 50~30x20mm, about 12 g) from one single fir lath were pyrolysed in a tubular furnace under an flowing argon atmosphere. The pyrolysis temperature of 600°C was reached at a heating rate of 20"C/min. The samples were kept at 600°C for 2.5 h to ensure complete pyrolysis, then the h a c e was turned off. Some untreated charcoal samples were pyrolysed further for 1 h at 8OO"C, respectively, 900°C. The argon flow was maintained until the sample had cooled down below 200°C to ensure that no reaction with air would occur.
74
Post-pyrolysis and gasification in the TGA The charcoal samples were milled in a mortar to very fine particles prior to each gasification experiment. In all cases, about 10 mg (+-0.5) mg of ground charcoal was gasified in a thermogravimetric analyser (Model TGA 51 from TA Instruments). To ensure that the pyrolysis stage has come to full completion, we first introduced a postpyrolysis stage by using a linear temperature ramp of 20"C/min to 900°C under a helium flow (220 ml,/min). During the heating ramp there is a typical weight loss of about lo%, which stabilises at 900°C. This temperature was held during 20 minutes, then lowered to the gasification temperature of 800°C. After stabilisation of thls temperature, the gasification agent (C02, 40 mldmin) was introduced to the purge gas. Figure 1 shows a typical run.
loo
80
p-h .
I
. ...
............
:
60
c
40
t I
E M
.-
20
200
0
0 0
50
100
150
200
250
300
350
time (min)
Fig.1 Typical TGA run with 10 mg of pulverised charcoal. ~ ~ , )serve as The time required for complete gasification of the charcoal ( T , , ~ ~ ~may a tool to quantify the overall gasification rate. However, a more revealing source of dormation is the reactivity, r, being the time derivative of the mass conversion degree, X, vk. r=dX/dt,' with ~ ( t ) = [ k f , - ~ / [ & k f o s h ] , M, the initial charcoal mass, kf its mass at time t, and kf& the mass of ash remaining after complete gasification of the sample. The conversion degree and the reactivity as a h c t i o n of time are calculated from the weight losses recorded during the gasification stage (see Figure 1) by appropriate subtraction and normalisation and by setting the time to zero at the start of the gasification.
CHARCOAL TREATMENTS
Impregnation Wood pieces taken from a single fir lath were impregnated with an aqueous Na2C03, respectively, NH&l solution (2 wt.-%) by stirring in the respective solutions for 6 h at room temperature, then boiled until complete evaporation of the water, and dried overnight at 120°C under ambient pressure. The salt impregnation procedure was 75
repeated several times and from the metal content analysis of the resulting charcoals we estimate an impregnation reproducibility of better than 10%.
Acid washing Literature reports that indigenous metal ions can be removed from coal char by acid washing." In order to lower the metal ion content in our charcoal, we stirred some charcoal in an aqueous HC1 solution (1 8 wt.-%) for 52 h, filtered and dried it. The metal removal efficiency was checked by metal content analysis which is discussed further down in the text.
CHARACTERISATION OF THE CHARCOAL Surface area determinations The BET surface area of several charcoal batches was measured with a Micromeritics ASAP 2000 automated gas adsorption apparatus. As adsorbent gas we used C 0 2 at 273 K as it may reveal more of the surface area residing in the micropore domain of the charcoal than, e.g., N2 at 77 K."
Metal content analysis About 100 mg of each charcoal sample was wet ashed with nitric and sulphuric acid according to the method described by Bajo et a1.12 and the metal contents of the samples (Na, K, Ca, Mg) were determined with ICP-AES (Inductive Coupled Plasma Atom Emission Spectroscopy, Varian, Model Liberty 110).
REPRODUCIBILITY To test the reproducibility of our charcoal sampling and that of the gasification experiment, two wood batches were pyrolysed separately and from each batch nine charcoal samples were gasified in the TG analyser. With the first batch we obtained an average overall gasification time of 194f13 min, with the second batch 200320 min. We conclude that the overall gasification time is acceptably well reproduced by both charcoal batches within an estimated uncertainty of 10%.
MASS TRANSFER LIMITATIONS In order to check if under given gasification conditions we might have external mass transport or pore diffusion limitations with the pulverised ~harcoal,'~.'~ experiments were performed under variation of the gas flow rate, the initial charcoal amounts and the charcoal particle sue, and the results are discussed below. External mass trznsfer limitations can be diminished by increasing the flow rate through the reactor as the latter decreases the stagnant gas region around the charcoal particle.' Two experiments were performed by raisied total flow of the gas in the TGA by 50%. The overall gasification times were 176 and 195 min, respectively, and, with both times being in the same range as the experiments with lower flow rates, no mass transfer limitations were observed. Another check is to change the initial weight of the sample (see Table 1A). There seems to be a slight influence of the initial weight on the 76
gasification time, probably due to the relatively poor contact between the gas and the char sample in the TGA geometry, but the values are still near to the experimental uncertainty level of 10%. To determine the possible influence of pore diffusion on the reactivity, the effect of the particle size on the overall gasification time was also examined as the size determines the length over which reactants have to d i f f i ~ eTable .~ 1B shows a typical particle size distribution of the pulverised charcoal determined by sieving. The particle size distribution varied somewhat between different samples, but the particles were never larger than 2 mm. The overall gasification time for various particle size fractions are shown in Table 1C. (The latter results were obtained with copper acetate impregnated charcoal the reactivity behaviour of which is very similar with that of the untreated charcoal.') Here again a small effect of the particle size on the gasification time is not excluded, but these effects are irrelevant compared with the overall gasification times resulting from added sodium catalysts or chlorine treatments reported in this study. Therefore, we conclude that our experiments revealed neither mass transport nor pore diffision limitations, in accordance with the predictions of other authors (see references cited in [l]). But, to guarantee a proper comparison between different samples, we gasify the same amount (10k0.5 mg) of pulverised charcoal (treated or not) under, else, identical gasification conditions.
Table I A Variation o f the charcoal weight. initial weight (mg) 5.22 10.1 15.7
Table IB Typical particle size distribution.
Toverall
166 198 230
< 125
unsieved
< 63
63 - 125 125 - 250
8
250 - 500
26
500 - 1000
20
> 1000
9
ToVera11 ( d n ) 152 184, 188 212 197, 159, 186
125 - 250 > 250
Wt-YO
15 22
Table 1C Variation of the charcoal particle size. particle size (pm)
size (pm)
EVALUATION OF THE KINETIC PARAMETERS From the characteristics of our reactivity curves (presented later), we selected the random pore model developed by Bhatia and Perlmutter' as the model can represent the behaviour of a system that shows a maximum in the reactivity curve as well as that of a system that shows no maximum. The maximum arises from two opposing effects: the growth of the reaction surface associated with the growing pores and the loss of surface as pores progressively collapse at their intersections (coalescence). In the lunetically controlled regime, the model equations derived for the reaction surface variation (S/S,) with conversion and conversion-time behaviour are given by:
-_S s,
-
I-x (l-tlcT)3
il-yln[
]
I-x
(1 - t I cr)3
77
with
x
= 1 - (1 - 7 /
v/ r / 4 ) )
~ 7e x ) p~ (-.r(l+
(2)
and T the dimensionless time and w, (J,structural and particle size parameters, respectively:
v =
4
7r
L 0 (1 -
Eo
(4)
)
so’
In Eq. (3) the initial reactivity is given by the parameter A,,. In Eqs. (3-5) the label o refers to the initial charcoal structure which is characterised by the reaction surface area per unit volume, So, the total length of the pore per unit solid volume, Lo, the particle radius, &, and the porosity, co. The surface reaction is characterised by the reaction rate constant K, and the reaction order n with respect to the reactant gas concentration C. Differentiating Eq. (2) with respect to T for (J + 00 (i.e., the reaction on the outer particle surface is neglected) one obtains d X - (1 --
dr
+ -)Y
r
2
v e x p (- r (1 + -))
r
4
The right-hand side of Eq. (6) can be rewritten entirely in terms of the conversion degree and after replacing dT with (A,,.dt) one has dX r =-=A dt
0
(I
- x )J1-
ln[l-X ]
(7)
RESULTS AND DISCUSSION 0VERAL.LREACTIVITY OF UNTREATEDAND TREATED CHARCOALS Table 2 shows the overall time needed to complete the gasification, for differently treated charcoals. The standard deviations were calculated fiom at least 3 gasification experiments. In the case of untreated charcoal, two different wood batches were pyrolysed and fiom each charcoal batch 9 different samples were gasified in the TG analyser.
78
Table 2 Overall gasificationtime (Tovedl) for untreated and treated charcoals (see text). Charcoal sample
Toverau
HCI washed NH&l (i)
580 (se) 348 & 197 f 16 201 f 13 76 f 6 33f2
*
Untreated Boiled NaCl (i)
Na2COi(i) (i = impregnated; se = single experiment) To verifL that the differences in the overall gasification times shown in Table 2 are not induced by the impregnation procedure itself, we impregnated the charcoal just with water. This "boiled" sample reacts as fast as the untreated charcoal. As expected from the charcoal enriched with sodium carbonate reacts much faster than the untreated one; under our conditions a factor of 6 faster. On the other hand, with chloride as the anion we observe a pronounced negative effect on the charcoal reactivity, but one may note that the NaCl impregnated charcoal still reacts much faster than the untreated charcoal. The negative influence of chlorine is not restricted to the metal chloride form as the charcoal impregnated with W C 1 (which decomposes in NH3 and HC1 around 34OoC)I6 needs ca. twice as long to react completely as the untreated one. It is therefore clear that the chloride species itself influences the gasification negatively. The cation exchange capability of acid washing, on the other hand, lowers the indigenous mineral matter content of the charcoal and, hence, their activity, with the effect that the HC1 washed charcoal reacted even 3 times more slowly. ELIMINATION OF EXPLANA TIONS
One possible explanation for the deviating reactivity with the treated charcoals is that the deposited salts lead to pore blocking, thereby reducing the huge reactive surface area located within the charcoal, or that the impregnated salt or the chloride treatment leads to a more porous charcoal considering the fact that in the production of high surface area activated carbons alkali and earth-alkali metal salts in various forms (carbonate, hydroxide or chloride) are added to the carbonaceous matter for this purpose." To check these hypotheses, the BET surface area of several untreated and treated charcoals was determined using C02at 273 K (See Table 3): The area of the untreated charcoal after wood pyrolysis at 600°C is somewhat lower than the C02 areas reported in the literature which, for Werent wood chars, range between 430 and 520 m2/g.' One may also note that the surface area value changes with the heat treatment temperature and passes through a maximum for pyrolysis temperatures around 800°C. This trend is explained in the literature."
79
Table 3 BET surface area of differently treated charcoals under variation of the pyrolysis temperature. ~~
Pyrolysis temperature:
600°C
800OC
900°C
BET surface area (m*/g)
Charcoal sample :
~~
NaCl impregnated
266 f 6
Untreated
310 f 27 (a)
HCI washed
328 f 13
(a)
@)
~
3 5 6 f 16 457 f 17
364 f 5
Standard deviation from three experiments; Other deviations reflect experimental Uncertainty from single experiments.
Compared with untreated charcoal, the other results obtained after pyrolysis at 600°C do not reveal any substantial difference in the measured surface area, except with the NaCl impregnated charcoal which showed a ca. 15% lower surface area. This may imply a moderate blocking of pores. Pore blocking phenomena have indeed been observed with coal char, but mainly at low temperatures where the impregnated alkali metal salts exhibit a negligibly small m~bility.'~ At sufficiently high temperatures, however, added alkali catalysts are known to achieve a very high dispersion in carbons, irrespective of the method of catalyst introduction, including that of dry mixing and cation exchange." Our BET results confirm this expectation as the difference in surface between the NaCl lmpregnated and the untreated charcoal disappears after prolonged pyrolyis at 900°C. Probably, NaCl distributes itself more evenly in the charcoal as the pure salt melts at 801"C.'6 Another possible explanation for the deviating reactivity is that the impregnation procedure expels the indigenous mineral catalysts from the wood by cation exchange. This possibility is not unlikely as indigenous metal species can be removed by acid washing." The metal composition of the studied charcoals are given in Table 4. The results are average results obtained from, at least, three different ICP-AES determinations made with the same wet-ashed charcoal solution. With untreated charcoal, we even analysed three different charcoal batches. From Table 4 it is seen that the metal content in untreated charcoal is constituted mainly by the indigenous K, Ca and Mg species. The estimated uncertainty in the indigenous metal content lies in the order of 20-30%, except with Na (100%) for which we suspect that its low amount (factor of 10 compared to potassium)' is susceptible to contamination from the glassware or solutions used. The metal composition in the charcoal which was just impregnated with water ("boiled") closely resembles that of the untreated one, supporting the idea that the indigenous metal content is not altered by the impregnation procedure itself. Washing the charcoal with HCl removes the alkali metals, Na and K from the charcoal, whereas the earth-alkali metals, Ca and Mg, remain. With Na2C03,NaCl, or NH4Climpregnated charcoal we can not clearly detect cation exchange effects, that is to say, with the latter two samples the respective indigenous metal contents are ca. 30% lower than with untreated charcoal, but also still close to the lower accuracy limit determined with the untreated sample. But, the apparent lower indigenous metal contents can, by no means, account for the large spread in the overall gasification times (see Table 2).
80
Table 4 Alkali and alkali-earth metal content (in MoVkg charcoal) of differently treated charcoals. (i = impregnated) K
Na
Untreated Boiled HCl washed
0.060 f 0.010 0.054 0
NH4CI (i) NaCl (i) Na2C03(i)
0.047 0.040
0.008 f 0.009 0.012 0.0004 0.0060 0.927 1.010
0.054
Ca 0.092 f 0.031 0.095 0.082 0.049 0.076 0.135
Mg 0.023 f 0.005 0.026 0.028 0.012 0.018 0.029
Summarising, we conclude that (1) hardly no indigenous cations are exchanged by salt impregnation, except by acid washing, and (2) no relevant pore blocking, nor increased porosity resulted from impregnation or HC1 washing. We expect therefore that the accessibility for the gasifylns agent, C02, will be the same for all charcoal samples studied.
CHARCOAL REACTIWTY DURING GASIFICATION The left plot in Figure 2 shows the experimentally derived reactivity course with untreated, respectively, HCl washed charcoal as a function of the conversion degree; the right plot those of untreated, respectively, Na2C03impregnated charcoal on a 14 fold larger vertical scale axis.
,
0.1
-.- 0.006 E
E
I
0.004
-3
NGO, impregnated charcoal
0.002
0.02
i,' J
0 0
0.2
0.4
0.6
0.8
1 0
Conversion degree, X
0.2
0.4 0.8 Conversion degree, X
0.8
1
Fig. 2 Left plot: Gasification reactivity of untreated, respectively, HC1 washed charcoal with C 0 2 at 8OOOC as a function of the conversion degree. The dashed lines represent model fits obtained with the original random pore model of Bhatia and Perlmutter. A and B indicate regions of mismatch between experiment and model. Rightplot: Gasification reactivity of untreated, respectively, Na2C03impregnated charcoal versus conversion degree. With untreated charcoal we obtain a maximum reactivity around X=0.3; a phenomenon which has often been obtained in the literature. With the untreated charcoal, the kinetic relation given by Bhatia and Perlmutter, Eq. (7), reproduces the initial part and the position of the maximum reactivity ( X m ) very well (see dashed line), provided that the model parameters (A, and w) are extracted from a limited experimental data range. Using the data between 0 5.X I 0.6 we obtain y = 5 and A, = 0.0057 min-', in line with 81
literature value^.^ For X > 0.6, however, the model fit clearly fails to describe the pronounced acceleration in the reactivity (see region "A"),as the pore model predicts a systematic drop to zero for X > XMm.By the same token, the model overestimates the time needed for complete gasification by a factor of 1.5.' We note that also other lunetic relations from other, either structurally based or empirical, models failed to describe our experimental reactivity results above X s 0.7.' ascribe the discrepancy between experiment and pore model to Some the collapse of the charcoal structure in the higher conversion regime. Before discussing the peculiarities associated with the overall collapse of the charcoal structure, we note that the observed acceleration in the untreated charcoal reactivity beyond XsO.6 is not only due to structural disintegration effects, but is partly the result of an increasing indigenous catalyst activity with progressing gasification. This picture emerges after comparing the two experimental reactivity courses shown in the left plot of Figure 2: The absence of indigenous alkali species in the HCl washed charcoal leads to a markedly lower initial reactivity than with untreated or the HC1 washed charcoal, but it does not affect the relative reactivity course up to Xs0.6. However, beyond X=0.6 the deviating reactivity behavior is much less prominent than with the untreated charcoal (see region "B")and, therefore, might be related to the particle disintegration effects only. The opposite case is encountered with the Na2C03 impregnated charcoal (see right plot in Figure 2), showing a much larger initial reactivity value (as expected) and a reactivity course with conversion degree which differs entirely from that of the untreated charcoal. The marked maximum reactivity around X 4 . 7 seems typical for alkali catalysed ga~ification,~.~ and it can not be denied that the deviating reactivity behaviour with the Na2C03impregnated charcoal is merely an extrapolation of a trend the onset of whch we already encountered with the untreated charcoal in the higher conversion regime. The pore model is unable to describe this late reactivity maximum around X=0.7 as it foresees a possible maximum reactivity to occur only between 0 IX < 0.393. In the literature, the late occurrence has been explained by intercalation of alkali metal species into the carbon structure, leading to a gradual release of active centres with conversion.6722 We note, however, that intercalation effects have seldom been reported for charcoals (in contrast to graphite). In our opinion the cause for the "anomalous" reactivity behaviour stems from a combination of structural and catalytic phenomena emerging from the reaction mechanism involved. The most important mechanism proposed nowadays is the oxygen transfer mechanism in which the oxygen is extracted from the reactant gas (C02) by the catalyst, which then supplies it in an active form to the carbon. The principal embodiment of such a process is a redox cycle involving carbothermic reduction of the alkali carbonate to the metal, followed by oxidation of the latter by the gaseous reactant (see, e.g., [l]). Hence, one may then also expect that the catalyst amount per unit available carbon area increases during gasification, since the carbon is removed continuously by the gasification reaction. Supporting evidence for catalyst accumulation may be found in our observation that the indigenous potassium concentration in untreated charcoal increased by 6% after 30 min of gasification, during which the charcoal weight lowered by lo%.'
82
EXTENDED REACTIVITY MODEL In the pore model developed by Bhatia and Perlmutter, the rate of the gasification reaction per unit pore surface area is characterised by the reaction rate constant, K,. As the original work addresses structurally based effects only, K, may well be assumed constant throughout the gasification stage and, under kmetic control, the char reactivity is then a direct measure of the available surface area. To allow the description of additional (i.e., non-porous) phenomena, we follow a semi-empirical approach in which we assume that K, can vary with time, the cause of which can either be structural or Ltalyst(t).Strictly catalytx in nature. Accordingly, we define K,(t) = Ksrmcture(t) to be time dependant when the gasification is a speaking, we only expect Ltalyst catalytically dominated reaction in which the catalyst concentration (or activity) changes with time. In such case, &atalyst(t)describes the relative change in the catalyst concentration with time. For phenomena such as particle disintegration, both K, and K;catalystwill likely remain constant, but the amount of reactive surface area may rise as fractures may reveal new surface areas with time. Hence, also for the latter case we assume K, to be time dependent, but now merely as a indirect mathematical tool to compensate for the exposure of new surface areas by fractures. The selection of an adequate time functionality for K,(t) was done working backwards from the experiment. As an example, for the HCl washed charcoal Figure 3 shows the ratio of the experimental to the fitted reactivity as a function of time. The fitted values were calculated with the original kmetic relation, Eq. (7), using fitted parameter values for A,, and y~which were extracted from the data range below X=0.7, i.e., the data range where Eq. (7) holds (see Figure 2). As Eq. (7) converges to zero for t + T ~ (a580 ~ min), ~ ~ the uratio n(t) = rexper~enc/rm~el becomes mfiitely large, hence, Figure 3 only shows Q(t) values for times up to Tend=530min (ie., for 97% of the conversion range) to avoid misleading interpretations. Figure 3 reveals that sZ(t) effectively equals unity for t < t* a 330 min, as expected, and that n(t) increases rapidly thereafter. Such a behaviour can very well be reproduced by a power law-lke time functionality of the form: S2(t)=1+B(t/Tend)P, where constant B fmes the end value of n(t) at time Tendand constant p positions the onset of the particle disintegration process around time t*, which we arbitrarily estimated from the equality 1+B(t*/Tend)P=1.05. The latter property is illustrated in Figure 3 by the other two dashed curves. Both curves converge to the same end point, but for curve (a) with p=l, t* a 0, for curve (b) with p=3.5, t* = 200 min, whereas from the fit one has p=7 with t* = 330 min. For completeness’ sake, we note that Kasaoka et al.23 introduced a comparable, twoparameter time functionality, i.e., r’(t)=atb, which, in combination with the entirely empirical relation (dX/dt)=r’(t)( 1-X), served the sole purpose of reproducing the inflection point in the char conversion with time. (With untreated charcoal this point marks the reactivity maximum around X=0.3). Therefore not surprising, their relation could not reproduce our reactivity data with untreated charcoal for X 2 0.7, but it also failed to describe the initial reactivity as their relation approaches zero for t 0.
83
Conversion degree, X
-
0.50
0.25
0
0.75
0.97
2
I
Y
1.5 h Y .4-
c: 1
0.5
I
I
0
I 100
I
I 200
I
I 300
Gasification time, t
I
I 400
I
I 500
[min]
Fig. 3 Experimentally derived and simulated time course for n(t) (p=l, 3.5, and 7 for the curves labelled a, b, and fit, respectively).
r
Tend
=
rexpe~ent/rmo~e~
Returning to the extension of the original pore model, we then introduced the time functionality, K,(t)=[ l+(bt)p], into the original model derivation where K, appears for the first time. This happens in the basic equation for the radial pore growth with time, viz., dr,,Jdt = K$". In all the following derivation steps we observed no violations of any other basic assumptions or boundary conditions set in the original derivation. The resulting kinetic relation for the conversion X reads
X = 1
-
exp(-z'(1
+
v/
z'/4))
(8)
Eq. (8) closely resembles the originally derived one, Eq. (2) for o+m, but now the correlation time, T, is replaced by a new one, T', which is given by the product of the original correlation time T defined by Eq. (3), times the dimensionless time functionality [ l+(bt)P],viz.,
r ' = r[l+(bt)P]= A , t [l+(bt)P]
(9)
In Eq. (9), b is a constant with dimension [time-l] and p a dimensionless power law constant. The resulting reactivity is then given by
84
Process information contained in the extended model parameters, b and p With uncatalysed gasification of charcoal, the functionality [l+(bt)p] appearing in Eq. (9) basically simulates the gradual emergence of new surface area (i.e., other than that already constituted by the charcoal pores) by the particle disintegration process, whereas the description of its decline with time by the gasification reaction is implicitly handled by the boundary conditions already set in the original derivation. The volumetric excess reactive surface area involved at time t can be estimated from Kstmcme(t) = [spore(t) + sexcess(t)l/ Spore(t) = 1 + (bt)P
(1 1)
The total amount of excess surface area exposed by the disintegration process can be estimated by replacing Sp0,(t) by Spore(0)) and extrapolating Eq. (1 1) to t+ToveraU, viz., Sexcess(Toverall)/Spore(O) = (bToVeral1)’. For gasification dominated by catalysis effects, the fimctionality [l+(bt)p] mamly describes the relative change in the catalyst concentration with time, viz., Kcatalyst(t) = [Cat(t>l/ [Cat(O>l= 1 + (bt)P
(12)
Fitting results In Figure 4A and 4B we present experimental and fitted reactivity curves of differently treated charcoals as a function of the conversion degree. To illustrate the descriptive potential of the extended kinetic relation, Eq. (lo), we include experimental and fitted reactivity results for NH4Cl, respectively, NaCl lmpregnated charcoal. The various results underline the very satisfylng agreement provided by the extended lunetic model. b and p), which we extracted Table 5 contains the fitted parameter values (A,,, simultaneously from the entire experimental data range, together with the overall gasification time, T ~ and ~ X*, ~ which ~ characterises ~ ~ ~the ,onset of particle disintegration, respectively, catalyst accumulation effects on the conversion scale. The X* values were estimated indirectly by solving the corresponding critical time, t*, from the equality [1+(bt*)P]= 1.O1. As mentioned before, with untreated charcoal the late acceleration in the reactivity likely comprises particle disintegration and catalyst accumulation effects. We expect the latter contribution to diminish with increasing demineralisation of the charcoal. Support for this is provided by the NH4Cl impregnated charcoal, which reactivity values lie well between those obtained with the untreated, respectively, the HC1 washed charcoal (see Figure 4B). Because of the similarity in the respective reactivity courses, we conclude that the HCl gas released by NH4Cl (around 34OoC)l6reacted with a part of the indigenous mineral matter of the charcoal, hereby increasing their volatility, hence, promoting their evaporation fiom the charcoal during the post-pyrolysis stage (T I 900OC) performed in the TG analyser.’ The advocated demineralisation is also indicated by the systematically lower initial reactivity values (A,,) and, by the same token, larger overall gasification times. These trends are accompanied by structural parameter values, which are substantially higher than with the untreated sample. We have no reason to believe that the hgher values indicate differences in the initial charcoal structure (see BET surface analysis) and we suspect that the higher w values merely depict a shift in the position of the maximum reactivity, from X = 0.3 (with the untreated sample) towards 0.4.This suggestion makes sense considering the fact that in the kinetic formulation of Bhatia and Perlmutter,’ the upper limiting value for
w,
w,
w
85
X~m=0.393is reached for y~ + 00. The release of additional surface area during the disintegration process is described by the dimensionless time functionality, [ l+(bt)P]. Figure 4C shows the fitted course for [l+(bt)p] as a function of the corresponding conversion degree. With untreated, NH4Cl impregnated, respectively, HC1 washed charcoal, the reactivity course is not dominated by catalysts accumulation effects and with all these charcoal samples, the reactivity contribution by the particle disintegration process, typically, starts around X*=0.8, in line with literature values reported for untreated charcoa1.4~z'~24 From the fitted functionality values (see Figure 4C) we estimate that the total amount of excess surface area released by the fragmentation process lies in the order of 40% of the pore surface area available at the start of the gasification. Saturation
Molar NdCarbon atom ratio in Na,CO, impregnated charcoal 0.012 I
0.024 I
(WC=O.l)
0.048 I
C
I
*
:'
3
0.008
.$
0.006
I
0.004
a
0.002
HCI washed & with NH,CI 01 0
I
,
0.2
i
,
I
,
I
,
0.4 0.6 0.8 Conversion degree, X
I 0.25
I
1 0
I
0.5
I 0.75
0.5 1
Conversion degree, X
Fig. 4A and 4B Experimental reactivity of various treated charcoals. as a function of the conversion degree. The dashed lines are model fits. Fig.4C Experimentally derived correction factors to the specific reaction rate constant KS(see text). Figure 4A shows that NaCl itself is not catalytically active during the early stages of the gasification as the initial reactivity closely resembles that of the untreated charcoal, and certainly not that of the NazC03 impregnated sample, the initial reactivity value of which is ca. five times higher. The deactivating influence of the chloride cation have been noted by many authors. Also known from literature is the reactivation of the alkali chloride by steam."z0 With both NaCl and NaZCO3the catalyst accumulation process typically starts after X*=0.2, but with NaC1, the catalyst accumulation factor, [l+(bt)p], rises less fast, suggesting that NaCl partly evaporates from the charcoal, whereas the remainder is transformed into the active carbonate From the results shown in Table 5 it is apparent that with the NazC03 impregnated charcoals the value of the structural parameter, y ~ differs , substantially with that obtained with the untreated charcoal. This is not surprising as the gasification 86
reactivity is determined almost entirely by the catalyst accumulation effect, hereby rendering the rise resulting from the pore growth (described by w) almost superfluous.
w,
Table 5 Extended model parameters {&, b, p } and related quantities. Estimated errors (in %) are given between the round brackets.
HCI washed a
0.0012
19.2
0.0015
7
590
0.72
with N h C I a
0.0026
13.3
0.0029
15
347
0.88
Untreatedcharcoal
0.0056 (23)
5.4 (4)
0.0051 (8)
14 (11)
197 (8)
0.82 (6)
with NaCl
0.0076 (15)
6.3 (27)
0.0156 (8)
5 (24)
77 (9)
0.23 (35)
with NazCO3
0.054
2.9(129)
0.070 (21)
3 (14)
33 (5)
0.19 (9)
(13)
Physical mechanism underlying the particle disintegration process Percolation theory concepts seem adequate for describing char-gas reactions as it provides a natural framework for modelling pore opening, enlargement and coalescence, in the evolution of the porosity and internally accessible surface area with solid conversion. In the kinetic regime, particle shrinkage due to chemical reaction on the external surface can often be neglected and, then, porosity and carbon conversion are ) . particle disintegration phenomenon is introduced as a related by X=(E-E,)/(l - ~ ~The percolative fragmentation process with the idea that, as the solid phase within the porous particle disappears by reaction, the porosity increases to a critical value, E*, where the particle core disintegrates into fine fragments (uniform percolation). Some a ~ t h o r s ~apply ~ , ~ 'percolation theory concepts to model the gasification of porous particles under different reaction regimes. For the chemically controlled reaction regime, however, these studies show that the disintegration phenomenon will hardly affect the reactivity course with conversion degree compared to the model predictions given by Bhatia and Perlmutter. We do not contest these findings, because they are justifiable within the framework of the assumptions made, but we doubt the adequacy of the key model assumption that no new surface areas are being created by the breaking up (disintegration) of the solid phase. Being the key issue at stake here, we note that the physical mechanism underlying the disintegration process is not well understood. Under kinetic fixed-bed conditions, Fuertes and MarbanZ7attribute the percolative fragmentation to topological features rather than to a mechanical phenomenon by considering the disintegration fragments as small particle regions which become isolated from the main core as a consequence of pore growing. We note that this type of isolation will not bsclose new surface areas constituted by fractures. Moreover, the percolation theory concept itself does not provide for the necessarily isolated porosity when X* > XMM.Reyes and JenseQ2* on the other hand, attribute the fragmentation process to particle-spanning fractures in the solid phase, but they assume the topological mechanism to dominate. From our gasification results with charcoal, however, the picture emerges that new surface areas are gradually being exposed by the disintegration process as soon as the critical conversion degree has been reached. Hence, our results plea more in favour for 87
a gradual breaking up of the solid phase (embrittlement) than for a cleavage-like fragmentation taking place instantaneously. We suspect that the postulated critical porosity value is not determined by topological features only, but also by material properties like structural stability. Our idea is that also stress (e.g., those induced by its own weight, gas pressure, or temperature gradients) may turn critical as the structural stability of the charcoal is undermined progressively by the carbon consumption. The observed gradual rise in the excess surface area then merely reflects the fracturing of the charcoal structure into fragments of various sizes in addition to fines. This process continues with the fragments, but gradually at higher conversion levels as their inherent lower weights create less stress, until eventually only fines will remain. To estimate the potential amount of surface area exposed by fracturing of the porous charcoal structure, we considered charcoal as a 3-dimensional matrix made up of elementary unit cells. Each cell consists of carbon and contains one void space within the cell to represent “average” structural quantities, like, porosity and reactive surface area. The unit cell has cubic symmetry, which is characterised by the cube length acell.The void space located concentrically within each cell has also cubic symmetry with cube length avoid(I a&. The assumed cubic symmetry has been chosen for mathematical convenience only and does not necessarily reflect the true symmetry, as with real charcoal the pores are interconnected, whereas in the model they are not. However, this difference is not crucial for the present estimation. At the start of the gasification, we have E~ = 0.7, hence, the ratio avoido/acel,” = ( E ~ ) ”= ~ 0.89. (The subscript O indicates gasification start.) The weight of the unit cell is determined by the non-porous carbon space, the density of which we estimated by p&on=1900 kg/m3.” Per cell, the reactive pore surface area is then given by Sporeo=6(avoid0)2 and the weight by pw~on[(ace~)3-[(avoi~)3]. Adjusting these relations to the experimental BET surface area of 310 m2/(g charcoal), it follows that acel; = 850 nm. With progressing gasification, the void space increases (pore growth) on cost of the surrounding solid carbon. In our approach, this process is simulated by increasing avoidwhile keeping ~ I constant. I The overall particle disintegration process starts at a critical conversion level (X*=0.7) and with the relation X(E)=(E-E,)/( l - ~ ~we ) ,estimate ~*=0.91.In terms of our simple structure model, this means that avoid*/acel;= = 0.97. At this stage of the gasification, the amount of surface area exposed by fracturing is estimated by cutting the cell in one dimension. The newly exposed surface area then equals Sexcess*= 2[(aCe1;)’-( avoid*)’] and constitutes ca. 2.5% of the pore surface area accessible at the start of the gasification. From our extended kinetic model fit with the HCl washed charcoal, we estimate that at the end of the particle disintegration process, the total amount of reactive surface exposed by fracturing equals ca. 40% of the pore surface area accessible at gasification start (see Fig. 4C). This suggests that in total each unit cell “suffered” around 16 (= 40%/2.5%) cuts in random directions and that we are left with much smaller particulates (sizes 50-200 nm)which are likely no longer porous (“fines”). Of course, the above estimates are very rough. For example, one may argue that in reality the surface area exposed by fracturing is likely (much) larger as the fracturing front does not propagate in one dimension necessarily, and that the carbon and fracture surfaces may be fractal in nature.” Notwithstanding these, the second maximum in the experimental reactivity marks most likely the end of the gradual disintegration process. With untreated charcoal, this second maximum occurs at X = 0.95 (see Fig. 2). The
-
88
solid nature of the resulting fines is also hinted at by the reactivity data with the untreated charcoal, as the final drop of the reactivity decays non-exponentially to zero when plotted against the gasification time." Moreover, we observed that the latter reactivity course with time shows a marked resemblance with the kinetic relation provided by the shrinking core model,I4 whch describes the reactivity progress of solid spherical particles as r = W d t ( l - t / ~ , , ~ ~ ~ ~ l l ) ~ .
-
Mechanisms accompanying the alkali catalyst accumulation process With respect to the dependence of the charcoal reactivity with the alkali metal content, the literature reports that the initial reactivity of chars impregnated with alkali carbonates increases systematically with the metal-to-carbon atom ratio (M/C) up to a saturation level, typically, around M/CzO. 1.29 From our metal content analysis (Table 4) and by assuming that the charcoal consists of carbon only, the initial atomic M/C ratio in the Na2C03 impregnated charcoals is ca. 0.012, and this value lies well below the saturation threshold mentioned earlier. In the literature it has been suggested that the late reactivity maximum around Xz0.7 (see Figure 4A) results from the saturation of the carbon surface area with catalytically active alkali species. (See, e.g., Hamilton et al.7 This explanation, however, is not supported by the catalyst accumulation factors (= [l+(bt)p]) derived by us as we find them to rise steadily with increasing conversion degree (See Figure 4C). Catalyst saturation may be defined as a state where the charcoal surface area is covered entirely by a mono-layer of catalytic species.30If we assume the extreme case that carbon, but not the added alkali species, is being removed from the charcoal, then, from the initial atom ratio it follows that saturation effects may be encountered, but not before ca. 88% of the carbon has been consumed by the gasification reaction. In equivalent terms, h s means that the saturation threshold may be reached for catalyst accumulation factors in the order of ten, and larger. Seen the much lower accumulation factor values obtained by us below XxO.88, we conclude that the added alkali species and carbons are both being removed from the charcoal during the course of the gasification, albeit each to a markedly different extent. Catalyst saturation effects, on the other hand, are indicated in the conversion range above Xz0.95,where the conversion of the remaining 5% carbon takes ca. 50% of the overall gasification time. In closing, wemote that the above outlined role of the catalytic species bears some similarities with the so-called auto-catalytic reactions, where one of the products of In catalysed gasification of chars, reaction acts as a catalyst (see, e.g., Le~enspiel).'~ however, the catalyst is not a reaction product, but its concentration rises likewise with the char conversion as the alkali species recycle.
ACKNOWLEDGEMENT We thank A. Schuler (PSI) for performing the ICP-AES analysis and F. Geiger (PSI) for the BET surface area analysis.
89
REFERENCES 1
von Scala C. (1998) The influence of contaminants on the gasification of waste wood, DPhil Thesis, ETH-Ziirich, CH, Thesis Nr. 12665. 2 Liliedahl T. & Sjostrom K.(1997) Fuel, 76(1), 29-37. Magnaterra M. Fusco J. R., Ochoa J, & Cukierman A. L. (1993) In: Advances in 3 Thermochemical Biomass Conversion, (Ed. by A. V. Bridgwater), pp. 116-130 Blackie Academic & Professional, Glasgow. Standish N. & Tanjung A. F. A. (1988) Fuel, 67,666-672. 4 Hamilton R. T., S a m D. A. & Shadman F. (1984) Fuel, 63, 1008-1012. 5 Wigmans T., Haringa H. & Moulijn J. A. (1983) Fuel, 1983,62, 185-189. 6 7 Wood B. J. & Sancier K. M. (1984) Catal. Rev.-Sci. Eng., 26 (2), 233-279. 8 Bhatia S. K. & Perlmutter D. D. (1980) AIChe Journal, 26 (3), 379-386. 9 Kapteijn F. M. & Moulijn J. A. (1986) In: Carbon and Coal gasification. Science and Technology, (Ed. by J. L. Figueiredo & J. A. Moulijn), pp. 291-360 Martinus Nijhoff, Dordrecht, NL. 10 DeGroot W. F. & Richards G. N. (1988) Fuel, 67,352-360. 1 1 Lopez-Peinado A, Rivera-Utrilla J., Lopez-Gonzales J. D. A. & Mata-ArJona A. (1985), Absorption Science & Technology, 2,31-38. 12 Bajo S., Suter U. & Aeschliman B. (1983), Anal. Chimica Acta, 149,321-335. 13 Szekely J. Evans J. W. & Sohn H. Y., (1976) In: Gas-Solid reactions, Chapter 4. Academic Press Inc., New York. 14 Levenspiel 0. (1972) In: Chemical Reaction Engineering, (Ed. by 0.Levenspiel), Ch. 12, John Wiley & Sons, New York. 15 Mims C. A. (1991) In: Fundamental Issues in Control of Carbon Gasification Reactivity, (Ed. by J. Lahaye & P. Ehrburger), pp. 383-407, NATO AS1 Series Vol. 192, Kluwer Academic Publishers. 16 Knacke O., Kubaschewski 0. & Hesselmann K. (1991) In: Thermochemical properties of inorganic substances, 2" edition, Springer Verlag, Berlin. 17 Greenwood N. N. & Earnshaw A. (1984) In: Chemistry of the elements, Chapter 8, Pergamon Press, New York. 18 Tseng H. P. & Edgar T. F. (1984), Fuel, 63,385-393. 19 Ohtsuka Y. & Asami K. (1996), Energy & Fuels, 10,431-435. 20 Mims C. A. & Pabst J. K. (1980), Am. Chem. Soc., Div. Fuel Chem. Preprints, 25 (3), 258-268. 21 Dasappa S., Paul P. J., Mukunda H. S. & Shrinivasa U. (1994) Chem. Engng. Sci., 1994,49,223-232. 22 Wigmans T., Goebel J. C. & Moulijn J. A. (1983) Carbon, 21,295-301. 23 Kasaoka S., Sakata Y., Kayano S., & Masuoka Y., (1983) International Chemical Engineering, 23 (3), 477-485. 24-26 From the present authors: PART I: Gasification reactivity of charcoal - A phenomenon revisited-, PART 11: Metal catalysed gasification of charcoal with CO,, and, PART 111: Influence of chlorine on the gasification reactivity of charcoal with CO,, sent to Fuel. 27 Fuertes A. B. & MarbinG. (1994) Chem. Engng. Sci., 49,3813-3821. 90
28 29
30
Marban G . & Fuertes A. B. (1997) Chem. Engng. Sci., 52, 1-1 1. Reyes S. & Jensen K. F. (1986) Chem. Engng. Sci., 41,333-343. Rensvelt E., Blomkvist G., Ekstrom C., Engstrom S., Espeniis B.-G. & Liinanki L. Symposium paper in: Energy from Biomass and Wastes, pp.465-494, 14.-18. August 1978, IGT, Washington D.C. Walker P. L. Jr., Shelef M. & Anderson R. A. (1968) In: Chemistry andphysics of Carbon, Vol. 4 (Ed. by P. L. Walker Jr.), p. 287, Marcel Dekker, New York.
91
Dynamic modelling of char gasification in a fixedbed B. Garbel, U. Henriksen, B. Qvale and N. Houbak Department of Energy Engineering, Technical University of Denmark, Nils Koppel AIIk, DTU - Building 403, DK - 2800 Kongens Lyngby, Denmark
ABSTRACT A dynamic one-dimensional model of char gasification in a fixed-bed reactor has been developed. The model is designed to serve as a tool for optimisation of the operating conditions of gasification plants. T h s involves the prediction of the dynamic behaviour of the gasification reactor where the load of the plant is varied. The model is based on conservation of mass and energy together with chemical equilibrium in the gas phase (using the water gas shft reaction). To determine the kinetics of the char used in the experiments (wood chps fiom beech), a series of TGA (Thermo Gravimetric Analysis) experiments were carried out. Based on the TGA measurements, analytical expressions for the reactivity were developed talung into account the reactivity of H 2 0 and COz, the inhibiting effect of H2 and CO, and the dependence on the conversion rate. Two different idealised models (shrinlung particle and porous particle) for the particle conversion were investigated. In order to validate the model, the results from the model were compared with experimental data. The data were obtained during tests on the 100 kW gasification plant at the Technical University of Denmark (DTU), using wood chips (beech). A test was carried out in whch the operating conditions of the plant were reduced from full load to quarter load in one step, in order to investigate the dynamic behaviour of the gasification system. The model was able to predict satisfactorily the temperatures, the gas production and the gas composition as a function of time and position in the char bed, and the height of the char bed as a function of time. INTRODUCTION Operations of multi-stage fixed-bed gasifiers have shown that gasification of char is the slowest process during the gasification of biomass. This has also been seen both in the present gasifier and in open-core gasifiers where the char gasification zone makes up the largest part of the reactor”” and determines the performance of the entire reactor3. An appropriate model that predicts the response of the gasifier, for example in terms of bed height and temperatures, to changes in operating parameters, will also constitute a good base for the development of a control strategy.
92
An important part of the description of the char bed gasification is the chemical reaction kinetics of the char. In this area limited attention has been paid to inhibiting effects on the reaction kinetics of H2 and CO in the gas. Experimental work4 has demonstrated that presence of, for instance, 10 % H2 in the reactant gas idubit the char reactivity with about 90 % compared with no content of H2. These effects are taken into account in the presented model. THE TWO-STAGE GASIFIER The plant used for the validation of the model is a 100 kW two-stage (Figure l), developed at Department of Energy Engineering, DTU. The plant is presented in Figure 1 and consists of an externally heated pyrolysis unit, an oxidation zone, where a partial combustion of the pyrolysis gas takes place, a down-draft char gasifier using air and possibly steam as gasification agent, a gas cleaning system and an internal-combustion engine with an electric generator.
h-
w -
I
I'r
MGlNE
C P
Fig. I The 100 kW Two-Stage Gasifier (top) at DTU, with Gas Cleaning (bottom)
93
This plant is based on the two-stage gasification process’, where the pyrolysis and the gasification are physically separated in two distinct reactors. This gives a well-defined gasification zone that is ideal for studying the char gasification at downdraft conditions (Figure 2). PREHEATED
Pyrolysis
Fig. 2 Char gasification reactor fitted with measuring equipment ASSUMPTIONS
The assumptions are: The ideal gas law is valid. Plug-flow, no radial gradients of concentration and temperature (l-D model) No accumulation of gas in the char bed. (The mass density of the gas is about 300 times smaller than the density of the char) No limitations from conductive and diffusive transport inside the char particles are included in the model. This assumption has been confirmed experimentally, using a “makro-TGA” *. The char consists of pure carbon. An ultimate analyse of wood chips (beech) pyrolysed to 600 “C (initial conditions) showed a low content of H and 0 remaining in the char. (C~.zsOo.04) There is no tar in the gasification zone. The tar content in the gas has been determined experimentallygto less than 2 g/Nm3above the char bed and 20 ppm. below. 94
The gas consist of N2,H20, H2, C02, CO, CH4. The methane is formed in the preceding pyrolysis process and remain unaffected through the char bed. The water gas shift reaction is at equilibrium: Measurements on the two-stage gasifier at DTU indicate that the gasses H20, H2, C02, CO was in equilibrium during the flow through the char bed”. In order to validate these results, a shortcut worst-case calculation made on the basis of the reaction kinetics presented by Biba et. al.” showed that the gasses were close to equilibrium. Several “An authors have observed equilibrium in the gas phase during ga~ification*’~~’~. error in this assumption would introduce only a small error in the energy balance because of the slight exothermicity of the reaction, and the gasification rate will not be affected” Constant atmospheric pressure in the char bed. The pressure drop over char bed is measured to be in the range 500 - 3000 Pa, which is to low to have a significant influence on the energy balance. Thermal equilibrium bet\l’een solid and gas locally. Calculations including energy transferred between char and gas by radiation and convection support this statement. The equations are illustrated in the section: “Temperature differences between the gas and solid phases”
DESCRIPTION OF THE MODEL A dynamic one-dimensional model of char gasification in a fixed-bed reactor has been developed. The model is based on conservation of mass and energy together with chemical equilibrium in the gas phase between H20, H2, COz, CO, using the water-gas shift reaction. Methane is assumed inactive in the char bed. The basic equations are:
MASS CONSERVATION FOR CHAR dMLhar dt
+
&char
=
char
MASS CONSERVATION FOR GAS
ENERGY CONSERVATION FOR ALL COMPONENTS U) +C-=Q’ a(m h) c-d(M‘. dt dx *
(3)
THE GASES ARE IN EQUILIBRIUM AFTER THE WATER-GAS SHIFT REACTION14 Ka (TI = pH,’ “O ‘ = (1 303. PH, ’ PCO,
T + 7.17. 10-4).T - 1.3006
95
(4)
CHAR REACTIYITY The reactivity of a solid fuel is commonly described as a function of the fractional (Burnon), X: con~ersion'~
where m is the mass of organic material during the conversion and X is the conversion ratio, defined as: m o- m XE(6)
mo
The reactivity, R, is d u e n c e d by the temperature, gas composition, total pressure, and other fie1 specific conditions as catalytic material and porosity. Describing the reactivity according to (5) & ( 6 ) and assuming that the kmetics are invariant to the conversion (X), it is possible'6 to split R into a chemical kinetics term rc and an independent structural term "a structural profile", f(X):
R = rc .f(X)
(7)
Reactivity for the biomass under study, wood chips of beech, is determined by TGA measurements carried out at the Norwegian University of Science and Technology (NTNU), Norway, and Riser National Laboratory, Denmark. Based on these, an expression for the calculation of the conversion rate as a function of the tested parameters has been established. 1000 900 5 800 e 700 2 600 500 400 300 200 100 0
100 90
80 70 60
$
0
50
50 40 30 20 10 0 100 150 200 250 300 350
-8
5 ' 3
time [minutes] Fig. 3A Measured temperature and weight during an experiment.
96
0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 Conversion ratio, X [-I
Fig. 3B Obtained structural profile on basis of the measured weight curve.
Experiments using H20/H2 for gasification have been carried out in Riser's PTGA" (Pressurized Thermo Gravimetric Analyzer) and experiments" using C02/C0 for gasification on two identical SDTs (Simultaneous Differential Scanning Calorimeter-TGA) at NTNU and Ris0, respectively. The following expression based on Langmuir-Hishelwood kinetics is used to describe the kinetic parameters in the gasification of mixtures of H20/H2 and CO2/CO, respectively. The effect of H2 inhibition and of CO inhibition is taken into consideration in these expressions: kl.fw
kl,fc
r- - =
'
PH-0
. Pco,
(9)
where kl,fc,kl,bc,kl,h,kl,bwand k3 are Arrhenius expressions of the form:
k = A . exp(
5)
The expressions rc,c and rc.w are combined in one expression describing the reactivity as a function of the temperature (T), the partial pressure of the reactant gasses ( pH,O, pH2 ,pco, ,pco ) and the degree of conversion (X):
kl,fw
k1,fw
1+-.
k3
'PH20 + kl,fc 'PC02
k l fc
PHIO+L' k3
Pco2 +
k l bw
k3
PH,
kl,bc -ip' k3
(11)
Pco
The constants have been established and are presented in Table 1 & Table 2:
97
Table 1 Activation energy for H20/H2/C02/C0gasification of beech char.
Activation energy [kJ/mol]
El,fw
El,bw
El,fc
199
146
248
El,bc 217
E3
225
Table 2 Rate coefficients for H20/H2/C02/C0gasification of beech char. Al,fw
Rate coefficients [s-'.am-'] 2.0.107
Al,bw
Al,fc
Al,bc
1.8.106
3.3.108
2.0.108
8.4.107
STRUCTURAL PROFILE, F G )
The structural profile, f(X), is determined on basis of the measured reactivity as a fbnction of conversion ratio, X, and modelled through a 6*-degree polynomial (Figure 3B). Figure 3B shows further more the increasing reactivity during the conversion. The experiments" showed that the structural profile was rising more slowly if the concentration of H20 is low and the concentration of H2is high. At the bottom of the gasification reactor, where the degree of conversion of the char, X, is high, the H20 concentration under normal gasification condition has fallen to a concentration below 30% and, at the same time, the H2 concentration has increased to a concentration of about 20% or more. According to the PTGA experiments, all measurements made under these conditions ( pH,O 2 0 .3 atm. and pH, 2 0.2 atm. ) have a lower value than the average at X = 0.8, described by the following polynomial of 6th degree, used in the model, obtained during data fitting to the observed profile (Fig. 3B): fa,,(X)=32.17.X6 -57.17.X' +46.10*X4-16.04*X3-I-2.92.X2 +0.297*X+0.529
DETERMINATION OF THE CONVERSION RA TI0
The conversion ratio, X, of the char is defined by (6). It is simple to determine the conversion ratio in a batch process but quite complicated to determine in a continuously operated reactor where non-converted char (X = 0) is introduced and the char later on are transported downwards during the gasification. In the presented work the conversion ratio for char (X) is calculated by using the ash content in char as an indicator for the conversion. During the gasification, the ash content in the char will increase (under the assumption of an inactive ash) and the degree of conversion can be calculated. The conversion ratio of the char is calculated as the ratio of the present ash content of the char to the ash content of the char at the start of the gasification process:
98
where: Yash =- mash and Yash,O
=-
mchar,O
mchar
w,h and m&, w,h,O and
m&,o,
mash,O
is the present mass of char and ash. is the mass of char and ash at the start.
In order to find the present ash content of the char, a mass balance for ash is established:
PARTICLE MODEL The conversion of char can be described using two different physical particle models, the shrinking-particle model and the porous-particle model, where p is the density of the particle: The porous-particle model presumes a reaction in the entire particle:
The shnnlung-particle model only presumes reactions on the outer surface:
TEMPERATURE DIFFERENCES BETWEEN THE GAS AND SOLID PHASES One of the assumptions prior to establishing the system of equation is that the temperatures of the solid and gas phases remain the same in any cross section, A, of the reactor. The validity of th~sassumption is examined by separating the gas and char temperatures and calculating the exchange of heat between these two phases. Radiation and convection transfer the energy between the char and the gas. Thls leads to the following expressions for radiation and convection. Radiation
where:
E~~~ = 0,019
&,.ha,. = 0,86
T = 0,981
Convection
99
p = 0,14
CT = 5,670
10- 8
where:
bloc =
-
0,9 1 Re-0'5'.
\y
.c
~.Go ,
~
~
~
Subscript f, means that the values are to ,e evaluated at film temperature: T f = %.(Tcbar+ TgaJ This leads to the following expression for the total heat transfer between the gas and the char.
L
1
[
m2 m reactor volume and A . dx is an incremental reactor volume. The separation of the solid and gas temperatures exerts a direct influence on the water-gas-shift reaction and the reactivity expression which is why the expressions are determined on the basis of the gas temperature and the char temperature, respectively. Simulations show only a minor temperature difference between the char and the gas phase. For the sake of calculation time and model simplicity, the temperature of the char and gas phase are assumed equal in the following.
where a is the specific heat-transfer area
HEA TING VALUE
The heating value of the producer gas is calculated from measured quantities using the following expression:
COLD GAS EFFICIENCY
Once the heating value of the biomass and of the producer gas are known, the cold gas efficiency of the plant can be determined
100
INPUT MODEL An input model for the determination of the conditions at the entry to the char bed is
constructed. The input model is based on conservation of mass, conservation of energy, and equilibrium between the gases as given by the water-gas shift reaction
SIMULATION The model is implemented in the simulation language SIL"
OUTPUT FROM MODEL The equations are established for a differential control volume. They are solved by integration from the top of the reactor to the bottom and the time variation of the following quantities are calculated: Temperatures locally down through the char bed. The gas composition locally down through the char bed. The mass of char in the bed and the height of the bed. Gas volume produced down through the bed. The reactivity down through the char bed as well as the averagehotal reactivity. The gas production is calculated either on dry or wet basis. Upper and lower heating values. The cold gas efficiency of the plant.
COMPARISON OF MODEL AND TEST In order to validate the model, results from the model are compared with experimental data. The data are obtained during tests on the 100 kW gasification plant (Fig. 1 & 2) at the Technical University of Denmark (DTU), using wood chips (beech). Th~splant is based on the two-stage gasification process, where pyrolysis and gasification are physically separated. Observations were made and measurements taken while the plant load was instantly reduced from full load to quarter load in order to investigate the dynamic behaviour of the gasification system. Table 3 Operating conditions during quarter load operation in the time period from: 00.00 to 07.30 [h] Operating conditions thermal power: rs25 kW 6.2 k g h fuel consumption moisture content of fuel 19 % mass flow of air 9.3 k g h mass flow of steam 0.0 k g h ~3384 "C preheating temperature of air preheating temperature of steam ~ 3 1 4"C
101
TEMPERATURE PROFILE AND GAS COMPOSITION In Figure. 4 the measured temperatures are compared with the calculated temperature profiles. The model overestimates temperature and temperature drop in the top of the char bed, but determines the outlet temperature in good agreement with the observed temperature. The best estimate is obtained using the shrinking-particle model. Only results based on the shrinking particle model are shown in the following figures.
u
900
2
950
1
time elapsed after change [h] exp. data: 00.50 0 exp. data: 00.55 + exp. data: 01.OO exp. data: 01.05 x exp. data: 01.10 -theory: shrinking particle -theory: constant particle size
b
R I f 700 I I I I I I I 720 660 600 540 480 420 360 300 240 180 120 60 1
I
I
I I
I
I I
I
I
I I
0
distance from grate [mm]
Fig. 4 Comparison of the measured temperature profiles (points) and temperature profiles (thermal equilibrium between char and gas) calculated from the model assuming 1 hour elapsed time (lines). Figure 5 shows a good agreement between the observed gas composition and composition predicted by the model. The major change in gas composition is observed in the upper part of char bed, whle only minor changes are observed in the lower and colder part of the bed. BED HEIGHT
After a step change in the operating condition to a quarter of the load, the mass of char in the bed moves towards a new level where char conversion in the reactor corresponds to the actual amount of char being introduced into the reactor under the new operating conditions. The height of the char bed is measured, using the obtained temperatures in the bed. Changing the height of the bed by changing the rate of feeding the reactor with biomass is normally a very slow process as illustrated clearly by Figure 6 where there is hardly any change after an hour, even though the load has been reduced by 75 %. After another 6 hours in operation the bed has only increased by about 12 cm which corresponds to an accumulation of char of about 1.9 kg.
102
+ co A C02
x cH4
720 660 600 540 480 420 360 300 240 180 120 60
0
Distance from grate [mm] Fig. 5 Comparison between the measured gas composition (points) and the model gas composition using the shrinking model for particle conversion (lines). 0.90 0.80 0.70
1 I ~
::
T
1
r
-0.60 E v 30.50 .-
2 0.40 -
B
0.30
-
0.20 0.10
-
HEA TING VALUE During the quarter-load operation part, the heating value was not determined regular intervals because gas measurements were made down through the char bed. At 07:20 hours the producer gas was sampled and analysed and the lower heating value was calculated to be 5.1 MJ/Nm3. At this point in time the model calculation showed a lower heating value of 4.9 MJ/Nm3.
103
AL TERh'A TIVE CONTROL STRATEGY
In the search for a more effective control strategy, the consequences of controlling the addition of air (rather than the addition of biomass) in order to maintain a constant bed height, even where operating conditions are changed markedly is investigated. This is done in order to reduce the largest time constant of the system. This control strategy is demonstrated in a mathematical simulation where the plant is taken from full to half load one hour after the start of the simulation. The addition of air is reduced from 24.89 kg/h to 14.87 kg/h. The air addition is chosen so that the bed height will remain constant. The simulation study shows that a constant bed height is maintained when going from full load to half load (change less than 1 %) (not illustrated), with the chosen reduction of air flow. , . I "
745
.
I
740 1
g 735 -i $ 2
730
-i
725 -1
2 720
i
s 5 715 -1
9
\710 -705
-:
700 120
Fig. 7 Bottom temperature (thermal equilibrium between char and gas) as a function of time
104
30
-7 25 0
-
Y
G
20
5
15
." .-
10
5-
:
CH4
CH4
Fig. 7 and 8 show that both the bottom temperature and the gas composition reach a new equilibrium within three minutes. During operation over a long period of time, gas composition or bed height remains unchanged. This means that the entire char-bed system has a short response time when aiming at maintaining a constant bed height. Therefore the plant can be brought to a stationary operating level witlun three minutes, even after a drastic change in operating conditions. CONCLUSION The aim of h s project is to construct mathematical models for char conversion in a fixed-bed gasifier. The understanding of the conditions during char gasification will facilitate a better control, design and the scaling up of gasification plants. In order to determine a control strategy for a gasification plant an understanding of height of the bed changes as a function of a given change in operating parameters are required. Thereby, an appropriate bed height can be maintained at any given time. Furthermore, test runs of the mathematical dynamic model shows that by choosing an operating strategy aimed at maintaining a constant bed height by changing the rate of addition of air rather than the rate of addition of biomass the large time constant of the char bed could be reduced. Following t h ~ salternative strategy, it was possible to bring the char bed to a stationary level quickly, even after a drastic change in operating conditions. Bed height, bottom temperature and gas composition reached a stationary level in just three minutes. The model is verified with measurements obtained on a 100 kW demonstration plant during operation. In order to focus on the dynamic aspects of the model the study was carried out when the operating conditions of the plant was reduced from full load to quarter load in one step. The model was able to predict satisfactorily the change in bed height as a function of time, as well as the time-variations of temperature and gas composition.
105
ACKNOWLEDGEMENTS The research was funded by The Danish Energy Agency.
NOMENCLATURE
k thermal conductivity kl,k3 reaction rate
Al, A3 reaction rate coefficient A.dx bed volume (solid plus fluid) a external char surface per unit bed volume cp heat capacity at constant pressure dQ energy per area El, E3 activation energy f(X) structural profile Go superficial mass velocity (empty reactor) h specific enthalpy Lo standard enthalpy bloc local heat-transfer coefficient of a cross section of the char bed K, equilibrium constant of the water-gas reaction
M' m m
p Q
0 R C r Re t T U
X X
Y
Greek letters f Source: char converted to gas during gasification Source per unit length E emissivity q efficiency v stoichometric coefficient,
mass per unit length mass mass flow rate partial pressure heat exchanged per unit length heat exchanged reactivity, gas constant kinetic factor Reynolds number time temperature [K] internal energy conversion ratio of char rectangular coordinate mass ratio
kinematic viscosity mass density, reflectivity Stefan-Boltzmann constant transmissivity empirical coefficient (particle shape)
Subscripts 0 initial conditions ash ash av. average b backward bc backward, CO2/CO- system bw backward, H20/H2 - system CO2/CO- system c CO carbon monoxide C02 carbon dioxide
char
char
dry
dry
f fc fw fuel gas
forward, film conditions forward, C 0 2 / C 0- system forward, H20/H2 - system fuel gas phase hydrogen steam H20/H2 - system
H2
H20 W
106
REFERENCES 1.
2. 3.
4.
5.
6.
7. 8.
9.
10. 11.
12. 13. 14.
Milligan, J.B., Evans, G.D. & Bridgwater, A. V. (1993) Results from a transparent open-core downdraft gasifier. Advances in Thermochemical Biomass Conversion, Proceedings of the International Conference on Advances in Thermochemical Biomass Conversion, 1 1-15 may 1992, Interlaken, Switzerland. Di Blasi, C. (2000) Dynamic behaviour of stratified downdraft gasifiers. Chemical Engineering Science 55, pp 293 1-2944. Manurung, R. K. & Beenackers, A. A. C. M. (1993) Modeling and Simulation of an Open Core Down-draft Moving Bed h c e Husk Gasifier. Advances in Thermochemical Biomass Conversion, Proceedings of the International Conference on Advances in Thermochemical Biomass Conversion, 1 1-1 5 may 1992, Interlaken, Switzerland. Barrio, M., Grabel, B., Risnes, H., Henriksen, U., Hustad, J.E. & Srarensen, L.H. (2000) Steam gasification of wood char and the effect of hydrogen idubition on the chemical kinetics. Proceedings of Progress in Thermochemical Biomass Conversion, 17-22 September 2000, Tyrol, Austria. Bentzen, J. D., Henriksen, U. & Hansen, C. H. (1999) Investigation of a TwoStage Gasifier. 2nd Olle Lindstrom Symposium on Renewable Energy, Bioenergy, 9-1 1 June, 1999, Royal Institute of Technology Stockholm, Sweden Grabel, B., Bentzen, J. D., H e d s e n , U. & Houbak, N. (1999) Dynamic Modelling of the Two-stage Gasification Process. Proceedings of the Fourth Biomass Conference of the Americas (vol. 2), Overend, R. P. & Chornet, E. (Ed.), Elsevier Henriksen, U. & Christensen, 0. (1994) Gasification of Straw in a Two-stage 50 kW Gasifier. Proceedings of the 8th European Conference on Biomass for Energy, Environment, Agriculture and Industry. Vienna. Stoltze, S., Henriksen, U., Lyngbech, T. & Christensen, 0. (1994) Gasification of Straw in a Large-Sample TGA. Part 11. Further Studies Introducing Lower Steam Concentration, C02-Gasification, and Gasification of Wood Chips. Nordic Seminar on Biomass Gasification and Combustion". NTNU, Norway. June 1994. Bentzen, J. D., Brandt, P., Grabel, B., Hindsgaul, C. & Henriksen, U. (1999) Optimering af 100 kW totrinsforgasningsanlzg p i DTU - Resultater fra forsrag i uge 37, 1998. (in Danish) (ET-ES 99-02), Department of Energy Engineering, Technical University of Denmark. Grabel, B. (2000) Dynamisk modellering af forgasning i fixed koksbed. Ph.D.dissertation (ET-PhD 99-04) (In Danish). Department of Energy Engineering, Technical University of Denmark. Biba, V., Maciik, J., Klose, E. & Malecha, J. (1978) Mathematical Model for the Gasification of Coal under Pressure. Ind. Eng. Chem. Process. Dev., Vol 17, No 1, 1978, pp 92 - 98. Yoon, H., Wei, J. & Denn, M. M. (1978) A Model For Moving-Bed Coal Gasification Reactors, AIChE Journal, Vol. 24, No 5, pp 885 - 903. Groeneveld, M. J. (1980) The Co-Current Moving Bed Gasifier. Ph.D.-thesis, TU Delft. Bentzen, J. D. & Grabel, B. (1995) Dynamisk model af totrinsforgasningsprocessen. (In danish) (PE 95-13). Department of Energy Engineering, Techmcal University of Denmark.
107
15. Laurendeau, N. M. (1978) Heterogeneous Kinetics of Coal Char Gasification and Combustion. Prog. Energy Combust. Sci., Vol4, pp. 221-270, Pergamon. 16. Ssrensen, L. H. (1996) Fuel Reactivity as a Function of Temperature, Pressure & Conversion. Ph.D.- dissertation, (R-838), Riss National Laboratory, Denmark. 17. Rathmann, O., Stoholm, P. 8z Kirkegaard, M. (1995). The Pressurized Thermogravimetric Analyzer at Department of Combustion Research, Ris0: Technical Description of the Instrument. (Riss-R-823(EN)), Riss National Laboratory, Roskilde, Denmark. 18. Barrio, M.& Hustad, J.E. (2000) C 0 2 gasification of wood char and the effect of hydrogen inhibition on the calculation of chemical kmetics. Proceed. of Progress in Thermochemical Biomass Conversion, 17-22 September 2000, Tyrol, Austria. 19. Houbak, N.(1987). SIL - a Simulation Language, User’s Guide, Lecture Notes in Computer Science, ed. by G. Goos and J. Hartmanis, Springer-Verlag, Berlin
108
Biomass treatment in supercritical water. The way from total oxidation to the gasification N. Boukis, J. Abeln, M. Kluth, A. Kruse, H. Schmieder and E. Dinjus Forschungszentrum Karlsruhe, Institut fur Technische Chemie, Karlsruhe, Germany
ABSTRACT In the supercritical water oxidation process (SCWO), biomass or organic contaminants in aqueous wastes are rapidly and quantitatively oxidised to harmless species at temperatures up to 600°C and pressures up to 30 MPa, avoiding an expensive off-gas treatment. Typical conversion efficiencies are higher than 95 % even at reactor residence times of less than 60 seconds. The process is self-sustaining for feeds with about 10 wt. YO organic matter. Plugging of the reactor by precipitating salts and corrosion of the reactor material are the two major technical challenges of the SCWO process. To avoid these problems, a new double pipe reactor with a porous inner pipe has been constructed and successfully tested. A tubular reactor made of Inconel 625 and titanium has been developed and proved in long time (1000 h) experiments with hydrochloric acid and oxygen containing feeds to be corrosion resistant. When the same process is performed without an oxidising agent, the organic matter is converted to gaseous products. Gasification of biomass with a high water content (up to 90 wt YO)in hot compressed water has several advantages compared with combustion or the traditional gasification. Under optimised conditions, the TOC of the liquid product and the CH4 and CO concentration in the gas product are low resulting in a high hydrogen yield (experimental results: 62 vol % H2, 32 vol % C02, 4 vol % CH4 and only 0.5 vol % CO). The gasification efficiency and the split of water by CO can be drastically increased by addition of KOH (<0.01 moV1). A first corrosion test has been performed, the analysis showed no worth mentioning corrosion. INTRODUCTION
Thermal treatment of high moisture biomass or aqueous streams loaded with organics require, under ambient pressure, the energy to evaporate the high water content. This energy can be easily recovered by thermal treatment under pressures higher than the critical pressure of water. The high density of the fluid under supercritical conditions enhances the space - time yield of the process. Further advantages of the high-pressure processes are different reaction pathways leading to a high hydrogen yield. The main part of carbon of the organic material is converted to C 0 2 and CH4.
109
OXIDATION (SCWO) Research on the applications of supercritical water started in the Institute for Technical Chemistry in the early 90's. At this time interest was focused on the oxidation of hazardous aqueous wastes in supercritical water (SCWO). Many industrial wastewaters containing highly toxic substances, which cannot be destroyed by biological treatment or disposed of, can be completely converted into harmless products [ 1-51. The SCWO process is operating at pressures and temperatures above the critical data for water (Pc = 22.1 MPa, T, = 374 "C), typically at 25-35 MPa and 450-650 "C. Under these conditions, water, oxygen (or air), C02 and most of the organic compounds, form a single, fluid phase, and oxidation rates are not limited by transport processes across phase boundaries. Consequently, SCWO is a process with high spacetime yields. Nitrogen containing organics form N2 and only traces of N20, NO/N02 formation has never been observed due to the relatively low oxidation temperatures. Other heteroatoms are mineralised to the corresponding acids (HC1, H2SO4 and H3P04) or salts. The formation of acids may lead to corrosion, the formation or the presence of salts to plugging, and many attempts have been made to solve the corrosion problem, by constructive or material means or by process control. Current R&D work is focused to the most promising engineering solution of the salt problem, a double pipe reactor with a porous inner pipe. The flow diagram of the installed SCWO bench scale plant (Figure 1) shows that feed, water and air are pumped and compressed separately, typically to 26-30 MPa. After preheating and mixing, the reactants are fed into the pipe reactor (PR) or double pipe reactor where oxidation takes place. After cooling and gas-liquid separation, the aqueous product and the off-gas are analysed. The whole SCWO bench scale apparatus is controlled automatically.
I
gasanalysis
airloqgen compressor
1 emuent feed system
preheater
readers
Fin. 1: Flow diaeram of the SCWO bench scale atmaratus.
110
aroling. depressurizing
The pipe coil reactor is made of Inconel 625, 15 metres in length with an inner diameter (i.d.) of 8 mm and an outer diameter (0.d.) of 14.4 mm. It is submerged in a fluidised sand bath, which is electrically heated and acts as a thermostat. The TWR used is a double pipe reactor, about 0.95 m in length, with an outer pressure bearing tube made of stainless steel (1.4989,o.d. 140 mm, i.d. 80 mm) and in inner porous tube of sintered stainless steel (1.4404, 0.d.: 66 mm, i.d.: 60 mm, average pore width 35 pm). The coil tube type reactor was first installed and most experiments until now have been performed with this reactor. OXIDATION OF MODEL COMPOUNDS
The oxidation of the model substances, ethanol, toluene and phenol, were investigated at 26 m a . Experiments were carried out between 400 "C and 550 "C at three different air feed rates. Organic feed rate was 0.25 f 0.02 kg/h, water feed rate 10 f 0.1 kg/h, resulting in a concentration of 2.5 w%. The flow velocities are 0.4-1.5 m/s, residence times 10-35 s, and Reynolds numbers 17000-33000. These data were calculated by means of steam data for water and air together with the flow rates and the dimensions of the tube. The other constituents of the solution have been neglected for this calculation. The resulting error in the calculated data is estimated to be less than f5 %. The reduction of TOC was found to be 99 % to 99.99 % in the case of ethanol. The SCWO treatment of toluene was performed in the same way as for ethanol, varying the temperature between 400 "C and 550 "C and the excess stoichiometric oxygen supply between 2 and 6. The destruction efficiencies were between 97 % and 99.99 % for the toluene oxidation. At 550 "C the destruction efficiencies were highest. The values of the off-gases for both, the ethanol and toluene oxidation, are for oxygen 7.5 v01.-% to 19.9 v01.-%, for carbon dioxide 0.5 v01.-% to 10.5 v01.-% and < 10 ppm to 1.4 v01.-% for carbon monoxide. At temperatures above 500 "C and a twofold stoichiometric oxygen supply, the CO contents are less than 50 ppm. OXIDATION OFREAL WASTES
Some results and conditions for the oxidation of real waste waters from the pharmaceutical, chemical and paper industry and our own sewage works in the tube reactor at about 27 MPa are summarised in table 1. The residence times were between 10 and 60 s. The feeds cover a broad range with respect to TOC, salt and solid content. The paper mill and sewage works effluents are containing solids of up to 5 %wt. The mean components of these solids are paper fibres with fillers inside. The other effluents are clear solutions. At higher temperatures the destruction efficiency are higher than 97%, while the feed concentration at this level up to 2.3 % TOC is not affecting the conversion. At high salt concentrations,the reactor plugs up eventually, but this blockage can be washed out.
111
Table I SCWO treatment of real waste effluents, conditions and results.
Waste type Pharmaceutical Industry Chemical Industry Paper Mill Sewage Works
Feed-TOC PDm
20,000 23,000 4,500 2,000 1 1,000 630 5.400
Salt content %wt 3
0.1 0.2 0.1
0.1
Temperature Conversion OC YO 550 550 550 500 500 550 550
97 99.99 99.98 99 97 98 99.8
CORROSION The complex load at temperatures up to 600 "C, pressures higher than 21 MPa and fluid densities from 100 to 1000 kg/m3, on the one hand, and a - in some cases - highly corrosive aqueous environment on the other hand, can lead to fast corrosion of the construction materials of SCWO reactors. Chloride is the most important corrosive species in SCWO treatment of wastes. After the oxidation of neutral or acidic feeds, the pH of SCWO solutions is frequently rather low. To keep the chemistry as simple as possible hydrochloric acid has been used to simulate chlorine-containing wastes in SCWO environments. In the time since 1994 a lot of materials, Stainless steel [6], Ni-base alloys [6-91, numerous high performance ceramic materials [7, 9-12] and metals like Titanium [13141, Tantalum [15], Niobium [16], Gold [6] and Platinum [6] have been tested under SCWO typical conditions in the Institute for Technical Chemistry) laboratories. The stainless steel 316 has been tested in supercritical water containing oxygen. After a 150 h exposure only some pitting could be detected in the section of the reactor with a temperature of about 350 "C. In solutions with 0.05 moVkg (1800 ppm) hydrochloric acid the UNS S3 1600 exhibit strong general corrosion. The measurement of the wall thickness reveals a maximum general corrosion of 0.6 mm / 150 h in the temperature range of about 350 "C. In the same solution with 0.5 moVkg the reactor made of this stainless steel leaked after only 10 h of operation due to stress corrosion cracking [6]. Screening tests with several Ni-base alloys shows that differences of the corrosion resistance between the tested alloys are within a factor of two and all alloys are rapidly dissolved in O2 and HC1 containing conditions at temperatures near the critical temperature of water [6, 171. At higher temperatures, corrosion rates are low. The corrosion mechanism is described in [8,9]. Some ceramic materials, modifications of alumina and zirconia are well resistant under these conditions [9-121. Titanium has been measured to be also well corrosion resistant in HCI and 0 2 containing solutions under SCWO cbnditions [ 13-14, 171. The latest corrosion experiments were performed with reactors made of alloy 625 and Titanium. The reactors in form of pressure tubes are used as samples. Since titanium does not own the required mechanical strength for high temperature - high pressure applications, it is used in form of liners for the corrosion tests. The length of 112
the tubes was 1000 mm, the outer diameter 14.3 mm and the inner diameter 8.5 mm. Only the ends (each 200 mm) of the tube reactors were lined with Titanium. The liners were thin-walled tubes made of titanium grade 2. In one experiment (with 0.05 m o k g HCl) the liner was only pushed inside the pressure tube in the other experiment both ends of the liners were electron beam welded with the outer pressure tube. The middle parts of the tubes were heated up to a maximum temperature of about 600 - 650 "C in order to reach an inner temperature of 600 "C. The ends of the tubes were cooled down to room temperature. After each experiment, the whole tube reactor has been examined metallographically. To produce oxygen, hydrogen peroxide has been used. After compression to 24 MPa in the low temperature section of the apparatus, hydrogen peroxide is decomposed to water and oxygen by a platinum catalyst. The resulting oxygen concentration was 0.5 mol [O,] ikg. The hydrochloric acid concentrations used for the two experiments were 0.05 and 0.2 mol/kg. Such acid concentrations are high for typical SCWO applications. Measure of the corrosion rates has been performed metallographically by optical and electron microscopy. During experiments, only the concentration of corrosion products can be measured. Since these corrosion products come from a wide temperature range, only a rough comparison of the corrosion rates during experiment can be made. Dissolution of Ni and Fe is much higher at the beginning of the experiments. The dissolution rate of Ti is high only during the start up operation. Generally, concentration levels are in the ppm and sub-ppm range, which is a strong indication for low corrosion rates. Chromium concentrations are mostly higher than Ni-concentrations. All of the chromium is measured to be Cr(V1). These measurements c o n f m one more the corrosion mechanism of alloy 625 proposed in [8]. The metallographic analysis of the reactors with the optical microscope shows not worth mentioning trace of corrosion in the high temperature section of the reactor The elemental analysis of the exposed surface revealed a layer reach in Ni and Fe oxides and depleted in Cr and Mo. These changes affect only a thin layer. In account to the long exposure time of more than 1000 h and the high (up to 0.2 mol/kg) HCl concentration, corrosion rates of alloy 625 in high temperature - low-density supercritical water solutions are defmitely low. For such HCl and 0, solutions Ni-base alloys, similar to alloy 625, can be used. Titanium has been proved sufficient corrosion resistant in high-pressure aqueous solutions containing HCI and 02.After the two long time exposures (longer than 1000h), no change of the liners thickness (470 pm) could be observed. The protective function of the Ti-liner can be fully achieved only if no corrosive solution reaches the small gab between liner and reactor wall. In the experiment with 0.05 molikg HCI the liner was not leak-free installed into the reactor. Even in this case corrosion of alloy 625 under the liner was less than about 200 pm in 1035 h. In the experiment with 0.2 m o n g HCI the liner was stiffly connected with the Ni-base alloy of the reactor. After experiment, the metallographic analysis revealed that the welding connection has leaked Corrosive solution reaches the gap between liner and alloy 625 tube and caused minor corrosion (see also [17]). The Ti-liner is, as expected not corroded. Corrosion rates of alloy 625 under the leaked liner are several orders of magnitude lower than for the unprotected alloy 625. For the oxidation of hazardous wastes in acidic supercritical water solutions, reactors made of alloy 625 or similar Ni-base alloys can be used, if the preheater and the cooler sections of the reactor are protected with liners made of Titanium. Some
113
effort must be taken to achieve mechanical stable, leak-free connection of the liners with the Ni-base alloy. DOUBLE PIPE REACTOR
To overcome the problems of corrosion and plugging a double pipe reactor with porous inner wall has been installed. Waste and oxidant (air) are fed at the top of the reactor by
feed
d,-
<
(
$heating
I
water (T < Ti)
I
I I
I
I
I
I*
I
I I I
I
I I I
4-
I I I
water (T < Ti)
I
I I I
/ porous wall (metal, ceramics)
pressure containment
1
oxidation products, dissolved salts, water
Fig. 2 Double pipe reactor with porous inner pipe
means of pumps or a compressor. This solution or slurry is brought to the supercritical state (with respect to water) using electrical heaters and the exothermic oxidation reaction. Due to the high temperature and pressure, the salts precipitate in this reactor part. Rinsing water, typically compressed to 30 MPa and preheated to 550 O C , is
114
pumped in the annular gap and passes through the porous pipe. This stream avoids the formation of a sticky layer of salts and can improve the corrosion resistance. In the lower part of this reactor the salt slurry is dissolved again, when subcritical conditions are adjusted by feeding quench water. Figure 2 gives a drawing of the reactor. For the treatment of salt containing waste effluents it is evident to find the most suited parameters of the temperature of the upper and lower reactor part, of the flow ratios of transpiring water to quench water for a given feed flow and of the feed flow to the air flow rate. At the time the optimisation of these parameters is close to be finished and there is strong evidence that this rather complicated reactor concept works well.
SC WO: CONCLUSIONS
-
-
Highly efficient destruction of toxic organic materials with high space-time yield No nitrogen oxides are formed. No expensive gas treatment required Heteroatoms are mineralised and leave the process with the aqueous phase Plugging of the tube reactor by dissolved or produced salts is a problem C1- induced corrosion can be minimised by Ti-liners in the preheater and cooler Corrosion and plugging are reduced with the double pipe reactor
GASIFICATION The second process for biomass treatment in supercritical water, the gasification to hydrogen, methane and carbon dioxide, is performed under similar conditions and in very similar apparatus like the oxidation. The target products of the process - fuel gases - present the more valuable energy form. Based on the experience of several years with the SCWO-process we try to optimise the energy efficiency of the gasification process and to treat effluents with a considerably amount of solids. On the other side, the gasification process does not need the compression of oxygen or air like the SCWO. Negligible energy is consumed to compress the feed. For more details see also [ 18-25]. GASIFICATION EXPMMENTS
The experiments were performed in autoclaves (100 and 1000 ml internal volume) as well as in two units equipped with tubular reactors (id. 8 mm, miniature plant 350 mm heated length; bench scale plant 15 m length). The batch reactors were used for lower temperatures, longer reaction times, and the tubular reactors for higher temperatures and shorter reaction times. Real biomass can be treated only in batch reactors up to now because feeding of small flows of suspensions is difficult. Gas chromatography was used to analyse the composition of the gas phase. The concentration of organic carbon in the liquid effluent was determined by a Total Organic Carbon (TOC) analyser. For a detailed description of the experimental set-up, the reader is referred to earlier publications [3,4, 18 1. Many experiments were performed in the tubular flow reactors to investigate the influence of the main parameters - temperature, pressure, residence time and alkali salt addition - on the gasification efficiency. Target is to produce a gas with a high hydrogen yield., In these experiments, only model compounds could be used because no reliable high-pressure slurry feeding system is available on the market for low flow rates up to now. 115
GASIFICATION RESUL TS
At temperatures lower than 550 "C, soot and tar are formed. For the experiments in the miniature plant generally a closure of the carbon balance better than 96 % could be achieved. The results show that at temperatures higher than 550°C a complete gasification of glucose without the formation or only with traces of solid and oily byproducts can be obtained. Residence times are up to 2 min. Samples of the aqueous effluent (residual TOC) contained some phenol and phenol derivatives. The experiments with glucose show that the carbon monoxide concentration in the product gas decreases more than twenty - fold by the addition of KOH. A similar influence is observed in the experiments with vanillin where potassium carbonate was added instead of KOH, (Table 2). It is known [26-281that the addition of alkali metals as catalyst increases the rate of the water gas shift reaction (CO + HzO S COz + H2). Table 2 Average CO and CH4product gas content with and without addition of KOH or K2C03. Feed M 0.1
T "C 550,600
vol% 20
co
CH4 vol% 3.7
Glucose + KOH, 0.18-2x10" MA
0.1-0.6
600
0.6
10.0 f 0.1 3.3 f 0.5
Vanillin without KzCO3
0.067
600
36
9.4 f 0.25
Vanillin + KzCO3 0.067 0.7- 1Sx1Om3h.1/1
600
1.5
12.6 f1.4
Glucose without KOH
The same influence of alkali salt is found at 450°C in experiments carried out in the tumbling reactor even if the reaction time is two hours (Figure 3). In the presence of K2C03, the CO formation is suppressed and the amounts of COz and hydrogen increase. Real biomass like straw includes alkali salts; therefore, no significant influence of additional alkali salts is measurable, see also [27]. The influence of residence time on the gasification efficiency was tested in the miniature plant for both glucose and pyrocatechol at 30, 60 and 120 sec. (feed: 0.2 M; 600"C, 250 bar, KOH). At these conditions the residual TOC content of the effluent solution of pyrocatechol gasification decreases from 2% to less than 0.1 % and the methane yield increases with reaction time (Figure 4). No tar or coke are formed and the effluent is colourless and odourless clear aqueous solution. For glucose, complete conversion and constant gas phase composition is already obtained at 30 s (Figure5). At feed concentrations higher than 0.6 M and the low potassium concentrationsused, a considerable decrease of the gasification efficiency as well as small formation of soot and tar is observed at a residence time of 120 sec. A marked pressure effect is observed for the CH4 yield. The CH4 content in the product gas increases from about 3 vol% at 250 bar to about 8 vol% at 450 bar along with a decrease of Hzof about 60 vol% at 450 bar to about 50 vol% at 250 bar. For these experiments, the closure of the carbon balance is better than 96% and no or only traces 116
of solids or oily material were found. Besides methane, C2-C4 hydrocarbons with C2H4 as the major component were detected in the product gas with a yield of below 1 ~01%. The comparison with results under higher pressure in capillary reactors [29] shows well comparable results (with the used reactors).
H,
I
40
S 30
-
O >
20 10
0 TOCres,=72%
TOCres.=26%
Fig. 3 Gas composition in vol% after the reaction of vanillin with and without K2C03 addition (tumbling reactor, 10 wt% vanillin, 450°C, 35 MPa, 2h reaction time). The gas phase composition is not corrected for the formation of KHCO3. The results show that water takes part as reactant to a large proportion under these conditions. For glucose, 76 % of the theoretical hydrogen formation is found according to the equation: C6H12O6 + 6 H20 + 6 C02 + 12 H2, for perfect gas: AH=158 kJ/mole [30] (1) The experimental result corresponds approximately to the following equation: 2 C6Hi2O6 + 10 H20 + 11 C02 + CH4 + 20 H2, (2) AH=152 kJ/mole [30]
0 0-
202b
'
0-
-
40
'
60
-,
,80
,_
I
co-0 100
120
Reaction Time /s
Fig. 4 Gas composition of pyrocatechol conversion (miniature reactor, 0.2 M Pyrocatechol, 600°C, 25 MPa, 0.0018 M KOH) as function of reaction time.
117
-4
-3q -
- 50-
. o
>
0
>
40-
0
b
- co,
- 2 0 0 A
A
co 201
e I
30
*
e
"d
1
I
90 120 Reaction Time / sec. 60
Fig. 5 Gas composition of glucose conversion (miniature reactor, 0.2 M Glucose, 600"C, 25 MPa, 0.0018 M KOH) as function of reaction time.
For the gasification of pyrocatechol, 82 % of the theoretical hydrogen formation is achieved if only carbon dioxide and hydrogen are considered as products. Glycine, a model compound for proteins and polyamides, was also used in gasification experiments in the miniature plant (600"C, 250 bar, 0.0018 M KOH) [24]. A residual TOC content of 1.5 % was found in the aqueous effluent at residence times of 60 and 120 seconds. The nitrogen of the amino acid is completely converted into ammonia and a hydrogen rich gas is produced because a large part of the carbon is bound as carbonate in the aqueous effluent (about 45 %). At lower temperature (400"C), vanillin is mainly hydrolysed to pyrocatechol in the absence of alkali salts and especially at higher pressures [18]. This selectivity vanishes at higher temperatures as the gas yield increases (stirred autoclave). The achieved TOC destruction efficiency is high although the real gasification efficiency could not be determined because the traces of solid and oily product are not quantitatively recovered. In the aqueous effluent pyrocatechol, methoxyphenol and phenol were identified as the major compounds by GC-MS. Both the low feed concentration and the higher reaction time compared with the tubular flow reactor experiments could explain this finding. With real biomass and wastes, only batch experiments have been done in our laboratory up to now. Table 3 shows results of the screening experiments with straw, wood and sewage sludge at temperatures of 450 and 500°C using the tumbling autoclave. As expected, a significant increase in the TOC destruction efficiency with increasing temperature is found for straw gasification, but no difference is found between the experiments with and without addition of potassium. The high potassium content of straw can explain this finding (ash: 4.6 wt YOwith about 15 wt YOK). Noticeable is the high yield of methane compared with the miniature tubular flow reactor experiments. The lower temperature and the longer reaction time of two hours with the tumbling reactor are an explanation. Altogether, the results confirm the findings described above as well as older results reported in [31] that temperatures higher than 550°C are required for complete gasification. 118
Table
3
Screening
Feed T [TOCIppm "C Straw' 450
experiments
with
the
tumbling
TOCDE
H2 vol%
CO CO;? vol% vol%
CH4 vol%
-0.8
33-43
1.2
40-47
17.5
500
>0.9
34-42
0.3
40-45
16.2
450
-0.9
30
1.7
49
19
450
-0S2 -0.853
49 47
2.8 1
31.2 37
17 15
addition
of
KzCO3;
autoclave
[4600]
Straw' [4600]
Wood2 [4900]
Sewage [2300] 'with
and
without
' [K2C03]: 1 .7x105 M; reaction time: -7200 sec.; p: 3 15-350 bar
'without
KZCO3;
GASIFICATION CONCLUSIONSAND OUTLOOK
The observed results show that a nearly complete gasification of the used model compounds can be achieved at temperatures of 600°C. The screening experiments with the real feedstock confirm the experiments done at the University of Hawaii. Compared to the traditional gasification process the following advantages for the hydrothermal gasification of wet biomass I organic waste feedstock can be expected: -much higher thermal efficiency -a hydrogen rich gas with low CO yield can be produced in one process step -soot and tar formation can be suppressed Further experiments have to be done to optimise the process parameters (pressure, additives) especially in view of higher feed concentrations (> 10 wt % organic), which are necessary to achieve a thermal efficiency high enough to establish an economic process. Other engineering challenges are the construction and testing of a reliable high pressure feeding system for slurries, to solve fouling problems of the heat exchanger, preheater and reactor caused by salty precipitates and to test corrosion resistant construction materials especially in regard to hydrogen embrittlement ACKNOWLEGMENTS G. Franz and W. Habicht have performed all reported corrosion experiments. Dr. M. Schacht, P. Rimbrecht and D. Mayer performed some of the gasification experiments. The authors would like to thank these colleagues for there contribution.
119
LITERATURE 1 . Schmieder, H.; Abeln, J. (1999): SCWO: Facts and Hopes, Chem. Eng. Technol. 11, 903-908. 2. Modell, M.; Mayr, S.T.; Kemna, A (1995): Supercritical Water Oxidation of Aqueous Wastes, in: Walker, J. (Ed.): Proc. 56IhInternational Water Conference, Pittsburgh, Pennsylvania, 30.10.-1.11.1995479-489. 3 . Goldacker, H.; Abeln, J.; Kluth, M.; Kruse, A.; Schmieder, H.; Wiegand, G. (1996) Process Technologv Proceedings, 12, Elsevier, Amsterdam, 1996,6 1-67 . 4. Schmieder, H.; Abeln, J.; Boukis, N.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri E. and Schacht, M. (1999). Proc. 5th Con$ on Supercrit. Fluids and their Applications, 13-16 June, 1999, Garda, Italy, 347-54. 5. Abeln, J.; Goldacker, H.; Kluth, M.; Petrich, G.; Schmieder, H.(1997). The Oxidation of Hazardous Waste in Supercritical Water, 4Ih Colloquium Supercritical Fluidr and Environment, Lyon, 20-21 Jan. 1997. 6. Boukis N. ; Franz G.; Friedrich C.; Habicht W., and Ebert K. (1996) Corrosion Screening Tests with Ni-Base Alloys in Supercritical Water Containing Hydrochloric Acid and Oxygen. HTD-Vol. 335 Proceedings of the ASME Heat Transfer Division (International Mechanical Engineering Congress and Exposition, USA) Volume 4, 159-167. 7. Boukis, N.; Friedrich, C.; Habicht, W.; Schacht, M. and Dinjus, E.(1997). Corrosion screening tests in supercritical water containing hydrochloric acid and oxygen. EUROCORR'97, Trondheim,Norway, 22-25 September, 1997 Proceedings, Volume I, 617-622. 8. Kritzer, P.; Boukis, N.; Dinjus, E. (1998). Corrosion of Alloy 625 in Aqueous Solutions Containing Chloride and Oxygen. Corrosion, 54,824-834. 9. Boukis, N.; Kritzer, P.; Schacht M. and Dinjus, E. (1999). The initiation of material corrosion in semicritical and supercritical aqueous hydrochloric acid solutions. NACE, Corrosion '99, USA, Paper No. 99256. 10. Boukis, N.; Claussen, N.; Ebert, K.; Janssen, R.; Schacht, M. (1997). Corrosion screening tests of high performance ceramics in supercritical water containing oxygen and hydrochloric acid. Journal of the European Ceramic Society, 17 (l), 71-76. 1 1 . Schacht, M.; Boukis, N.; Dinjus, E. ;Ebert, K.; Janssen, R.; Meschke F.; Claussen, N. (1998). Corrosion of zirconia ceramics in acidic solutions at high pressures and temperatures. J. Eur. Cer. SOC.18,2373-2376. 12. Schacht, M.; Boukis, N. and Dinjus, E. (1999). Corrosion of alumina ceramics in acidic aqueous solutions at high pressures and temperatures. 6th Conference and Exibition of the Eur. Cer. Soc., Brighton, UK, June 21-25, 1999, ECERS 99. 13. Boukis, N.; Friedrich, C.; Dinjus, E. (1998). Titanium as reactor material for SCWO applications. First experimental results. NACE, Corrosion '98, Paper Nr. 984 17. 14. Boukis, N.; Friedrich, C.; Dinjus, E. (1999). Corrosion of Titanium under SCWOConditions. Recent Results. WissenschaJtlicheBerichte, FZKA-6271(Juni 99) Karlsruhe 10 1- 102. 15. Friedrich, C.; Kritzer, P.; Franz, G.; Boukis N. and Dinjus E. (1999). The corrosion of tantalum in oxidizing sub- and supercritical aqueous solutions of HCl, H2S04and H3P04.J. Material Science 34,3 137-3141.
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16. Kritzer, P.; Boukis, N.; Franz, G.; Dinjus, E. (1999). The Corrosion of Niobium in Oxidizing Sub- and Supercritical Aqueous Solutions of HCl and H2S04J. Material Science Letter 18,25-27. 17. Boukis, N.; Franz, G.; Habicht, W. and Dinjus, E. (2000). Corrosion in Supercritical Water containing hydrochloric acid and oxygen - Problem solutions. EUROCORR 2000 Proceedings, London UK, 10 -14 September, 2000. 18. Kruse, A.; Abeln, J.; Boukis, N.; Dinjus, E.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht M. and Schmieder, H..( 1999). WissenschaftlicheBerichte, FZKA-6271 (Juni 99), Karlsruhe, 1999, 1 1 1. 19. Boukis, N.; Schmieder, H.; Abeln, J.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Schacht, M.; Sadri, E. (2000). Gasification of High Moisture Waste Biomass in Supercritical Water. The fifth international symposium on supercritical fluids, ISSF 2000 Proceedings; April 8-12,2000, Atlanta, USA(in press). 20. Schmieder, H.; Boukis, N.; Abeln, J.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Schacht, M.; Sadri, E. (1999) Gasification of High Moisture Waste Biomass in Supercritical Water. Fifth International Conference on Carbon Dioxide Utilisation, September 5-10, 1999, Karlsruhe, Germany, 1 12. 2 1. Xu, X.; Antal, M.J. Jr. (1998). Gasification of Sewage Sludge and other Biomass for Hydrogen Production in Supercritical Water. Environmental Progress 17(4), 2 15220. 22. Xu, X.; Matsumura, Y.; Stenberg, J.; Antal, M.J. jr (1996). Carbon-catalyzed Gasification of organic feedstocks in supercritical water. Ind.Eng.Chem.Res. 35; 25222530. 23. Schmieder, H.; Abeln, J.; Boukis, N.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E. and Schacht M. (2000) Hydrothermal Gasification of Biomass and Organic Wastes J. Supercrit. Fluid 17, 145-153. 24. Kruse, A.; Abeln, J.; Dinjus, E.; Kluth, M.; Petrich, G.; Schacht, M.; Sadri, E.; Schmieder, H. (1999). Gasification of Biomass and Biomass Model Compounds in Hot Compressed Water. AIRAPT-1 7: International Conference on High Pressure Science and Technology,25-30 Juli, 1999, Honolulu, Hawaii, USA (in press). 25. Kruse, A.; Meier,D.; Rimbrecht, P.; Schacht, M. Gasification of Pyrocatechol in Supercritical Water in the Presence of Potassium Hydroxide. The fifth international symposium on supercritical fluids, ISSF 2000 Proceedings; April 8- 12,2000, Atlanta, USA (in press). 26. Elliot, D.C.; and Sealock, L.J. Aqueous Catalyst Systems for the Water-Gas Shift Reaction. 1. Comparative Catalyst Studies. Ind Eng. Chem. Prod. Res. Dev. 22, 1983, 426-43 1. 27. Elliot, D.C.; Butner, R.S.; and Sealock, L.J. Low-Temperature Gasification of High-Moisture Biomass. Research in Thermochemical Biomass Conversion, April 1988, Phonix, USA. 28. Elliot, D.C.; and Sealock, L.J. Low-Temperature Gasification of Biomass Under Pressure. Fundamentals of Thermochemical Biomass Conversion, 937-950. 29. D. Yu, M. Aihara, M.J. Antal (1993) EnergydFuels 7, 574-577. 30. Calculated by "Benson", NIST Standard Reference Database 25, Structures and Properties, Version 2.02, Jan. 1994. 31. S. Ramayya, A. Brittain. C. DeAlmeida, W. Mok, M. J. Antal, Jr., (1986). Fuel 66, 1364.
121
Characterisation method of biomass ash for gasification 4. Moilanen', L. H. Sarrensen2, T. E. Gustafsson3, J. Laatikainen-
E. Kurkela' 'Luntama', VTT Energy, P. O.Box 1601, FIN-02044 VTT, Finland ' ReaTech, c/o Centre of Advanced Technology, P. O.Box 30, DK-4000 Roskilde, Denmark VTT Manufacturing Technology, P. 0.Box 1 703, FIN-02044 VTT
ABSTRACT Tendencies for ash agglomeration and deposit formation in biomass gasification were studied in two atmospheric-pressure fluidising bed reactors together with reactivity tests carried out using a thermobalance. Various herbaceous and woody biomasses of different origin and a selection of Danish straws were tested. The measurements were carried out as a function of temperature, and H 2 0 and COa were used as gasification agents. The reaction gas pressure was varied between 1 and 30 bar, and the temperature range was 650 - 950 "C. The appearance of the ash residues after the gasification reaction was examined by microscopy. Bed agglomeration and freeboard deposit formation were monitored by collecting samples after test runs. Samples were analysed by using a computer-controlled scanning electron microscopy (CCSEM) technique developed for ash deposit analyses. The thermobalance measurements gave results comparable to the ash behaviour in the fluidised-bed reactor. The strongest ash sintering was observed for wheat straw both in the thermobalance and in the fluidised-bed reactor. The effect of additives to prevent agglomeration formation was included in the study. INTRODUCTION Many potential biomass feedstocks, such as straw, have a problematic ash composition, which causes sintering and fouling problems in combustors. In biomass combustion, ash deposit formation is a common problem and has been studied by a number of researchers [l - 51. There are also some previous observations about operational problems in fluidised-bed gasification processes, caused by ash. In pressurised steam-oxygen gasification of peat, ash deposits have been formed in the upper part of the gasifier and in the cyclones [6].Furthermore, straw ash has been found to cause both bed sintering and deposit formation in pressurised air-blown bubbling fluidised-bed gasification [7]. These problems were difficult to overcome in straw-alone gasification. In fact, the gasification temperature had to be reduced to below 800 - 850 "C, which resulted in
122
poor carbon conversion and high tar concentrations. On the other hand, co-gasification of coal and straw (up to 50 wt% straw) was carried out without any signs of ash problems in spite of high operation temperatures of the order of 950 - 980 "C. Char reactivity and ash behaviour are factors limiting possible operation conditions in all gasification processes when planning the use of new fuels. Thus, the purpose of this work was to create data that can also be used in the development projects of gasification of various biomass feedstocks. The ash sintering behaviour was characterised using a thermobalance, and larger-scale tests were performed in two atmosphericpressure fiuidised-bed reactors, a bubbling bed reactor (AFB) and a circulating fluidising bed reactor (CFB). Bed agglomeration and deposit formation in the freeboard were monitored by collecting samples after test runs.These samples were analysed by computer-controlled scanning electron microscopy (CCSEM). The effect of additives to prevent agglomeration was also included in the study. The sintering part of the study was a continuation for a work reported previously [8].
LABORATORY STUDIES The samples used in the study comprised various biomasses and a selection of Danish straws. The analyses of the fuels used in the study are presented in Tables 1 and 3, and the chemical composition of ash, together with a summary of the thermobalance sintering tests, in Tables 2 and 4. The Danish wheat straw qualities were selected on the basis of growth site, fertilising and weathering conditions [9].
ASH SINTERING The ash sintering behaviour was determined in a thermobalance according to the method described in [8]. The ash residues from the thermobalance experiments were studied by microscopy using the following classificationcriteria: 1. Non-sintered ash residue: ash structure resembling the original fuel particles, easily crumbling when touched (no asterisk 0) 2. Partly sintered ash (different degrees in this group): particles contained clearly fused ash (1 or 2 asterisks: *, ** ); 3. Totally sintered ash: the residue was totally fused to larger blocks (3 asterisks: ***). The results are given in Tables 2 and 4 completed with the data of ash chemical composition [8]. Accordingly, the results showed that ash sintering was dependent on the pressure of steam. For spruce bark, willows and alfalfa, as shown in Table 2, ash sintered stronger in pressurised conditions than at atmospheric pressure. When the ash composition data of Table 2 are compared to the observed degree of sintering, it seems to be evident that in the samples (i.e. spruce bark, willow and alfalfa), having a higher sintering tendency under pressure than at atmospheric pressure, the silica content was very low (40%) and calcium as well as potassium contents were relatively hgh. Pine bark and spruce bark are very similar with respect to ash chemical composition. No sintering was observed in pine bark, but very strong sintering took place in spruce bark under pressure. We should suggest that this be due to the high alumina content (or the ratio alumindpotassium) in pine bark ash. The behaviour observed under pressure can be due to the carbonate chemistry involved in the ash. According to phase diagrams presented in literature [ll], the mixture of calcium and potassium carbonate forms a eutectic at about 750 "C. T h ~ s behaviour provides, of course, that carbonates are formed in the ash. During gasifica123
tion, carbonates can be formed, and they decompose slowly, especially at higher COZ partial pressures. To confirm this, gasification tests were also carried out at 1 bar COZ, which is above the C02equilibrium pressure at 850 "C [8]. For willow and spruce bark samples, the sintering was pronounced when measured at 1 bar C 0 2 whde in 1 bar steam there was no sintering at 850 "C. However, alfalfa had not the same behaviour and indicated only weak sintering at 1 bar in COz, while it strongly sintered under pressure in steam even at the temperature of 600 "C (Figure 1). To see the effect of temperature on the melting of carbonates, additional tests were carried out for spruce bark and willow in 30 bar steam and at 700 "C, which was clearly below the eutectic. The results showed that the spruce bark had no ash sintering but the willow ash was clearly sintered. This observation and the behaviour of alfalfa indicate that chemical compounds other than carbonates are also involved in ash sintering. One explanation could be found in the chlorine chemistry: Below 700 "C the major part of chlorine remains in ash. The chlorine content of Danish straw ashes, which were prepared at 550 "C in a laboratory furnace, is shown in Table 4. Mixtures of compounds CaCl and KC1, and also KCl and K2C03are possible. The former mixture has the eutectic of 595 "C and the latter 640 "C [ 113. In the major part of the Danish straw samples, the sintering behaviour was fairly similar, between ** and ***, as shown in Table 4. In straws ##4,#9 and #11, the degree of sintering was less severe, from * to 0.This behaviour is related to the ratio of the potassium and silicon contents in the sample. Very low SiO2/K2Oratios cause a high reactivity and a low inherent sintering tendency, while very hlgh SiO2/K2Oratios cause a low reactivity and also a low agglomeration tendency. Medium-value SiO2/K2Oratios cause a medium reactivity and a high inherent sintering tendency [9]. Table I Analyses of various biomass feedstocks (drybasis). Sample Pine bark Pinesawdust Sprucebark Needles** Finnish willow
Ash Volatile Fixed LHV* C H N % matter% carbon% MJkg % % % 1.7 73.0 25.3 19.7 52.5 5.7 0.4 0.08 83.1 16.8 19.0 151.0 6.0 0.1 2.3 75.2 22.5 18.5 49.9 5.9 0.4 3.6 75.3 21.0 n.d. n.d. n.d. n.d. 1.2
79.9
18.9
n.d.
49.7 6.1 0.4
O(diff) 40.0 42.8 41.4 n.d.
C1 ppm 0.03 85 0 4 0 0.03 279 n.d. n.d.
42.6
0.03 n.d.
%
Swedish 1.3 80.5 18.2 n.d. 49.4 6 0.5 42.7 willow Danish wheat 4.9 77.4 17.7 17.7 47.3 5.8 0.5 41.4 straw A (-93) Danish wheat 4.5 76.9 18.5 17.5 47.1 5.9 0.6 41.9 straw B (-93) Danish wheat 4.8 76.1 19.1 17.4 47.5 5.9 0.7 41.2 straw -95 Danish wheat 6.1 75.8 18.1 17.3 '46.5 5.7 1.4 40.1 straw -97 Alfalfa 5.0 75.8 19.2 18.4 45.8 5.4 2.2 41.5 * Low Heating Value, MJkg, n.d.: not determined, **half pine half spruce
124
S
%
0.03
130
0.07 1770 0.07 3 190 0.16 5 200 0.12 4 360 0.1 3 920
Table 2 Ash chemical composition of the various biomass feedstock, and ash sintering test results (Sint750 = sintering degree at 75OoC,Sintsso= sintering degree at 850OC). Sample
Ash
Ash chemical composition, % in ash
Sintsso
Sint750
Si02A1203Fe203CaOMgOK20Na2CTi02S03 P20 1bar 30bar 1bar 30bar 5 H2O HZO H2O H20 o n.d. n.d. 1.7 1.3 5.3 0.3 40.6 4.5 7.6 0.5 0.1 2.1 4.8 o %
Pinebark Pinesaw dust Sprucebark Needles Finnish willow Swedish willow Finnish wheat straw Danish wheat straw A (-93) Danish wheat straw B (-93) Danish wheat straw 95 Danish wheat straw 97 Alfalfa
0.0 8.3 8 2.3 1.5 3.6 34.2 1.2 0.4
2
1.8 41.8 11.8 12.3 0.3 0.1 1.9 5.2
*** ***
1.1 35
13.20.250.04 0.9 2
***
n.d. n.d. o
***
***
o
**
*** ***
slow
***
3.5 9.6 0.8 0.1 1.9 7.8
7.4 68.4 0.85 0.45 4.3 2
0
*** * ***
1.1 0.1 39.25.1 7.60.4 0 1.04.1 o 0.5 0.2 23.83.32 7.7 0.1 0.02 1.5 5.0 o 0.3 0.2 30.8 5.1 26.5 0.3 0 3.0 11.5 o
1.3 7.5 1.6
o
o
0
4.9 72.7 0.7
0.5 7.3 1.6 8.1 0.4
0 n.d. n.d.
*** ***
n.d.
**
4.5 49.2 0.4
I
0 n.d. n.d.
*** ***
*
*
4.8 34.2 0.3
0.2 8.4 2.2 30.1<0.50.02 3.2 3.9
*** ***
**(*)
n.d.
6.1 29.9 0.8
1.3 10.6 5.6 27.7 1.5 0.07 5.5 10.8 ***
n.d.
**(*)
n.d.
5.0 3.0 0.3
0.7 22.4 9.0 27.7 1.9 0.02 3.0 8.0
**
***
*
***
9.7 2.2 18.1 0.3
o no sintering; * slightly sintered; ** moderately sintered; *** completely sintered; n.d.: not measured; slow: low reactivity Alfalfa 20 bar, H20 c 200
.E 3- 150 se
75
80
85 90 Fuel conversion, %
- -
95
100
6OOOC 650°C 675°C 700°C 750°C
*
*r)
n+
Fig. I The gasification reactivity and ash sintering behaviour of alfalfa in 20 bar steam at different temperatures. 125
Table 3 Proximate and ultimate analysis of Danish straws (drybasis). Sample
# I Wheat #2 Wheat #3 Wheat #4 Wheat #5Barley(W) #6 Grass #7 Barley@) #9Wheat98/rib # I 1 Wheat/rib+
Ash Volatile Fixed % matter carbon 5.4 5.4 5.3 3.5 5.7 5.2 5.5 15.5 8.0
%
%
76.5 77.1 76.2 78.2 76.2 75.9 76.2 68.0 75.3
18.1 17.5 18.5 18.3 18.1 18.9 18.3 16.5 16.7
LHV,
C
H
N
0
MJkg
%
%
%
%
S %
17.27 17.36 17.13 17.53 17.16 16.89 17.07 15.76 17.13
46.9 46.4 46.9 47.2 46.4 46.7 46.5 42.1 46.1
5.8 5.9 5.8 5.9 5.7 5.9 5.7 5.2 5.5
0.8 0.6 0.7 0.5 0.7 1.0 0.5 1.2 0.9
40.9 41.6 41.2 41.5 41.4 41.1 41.7 35.9 39.4
0.18 0.13 0.13 0.13 0.13 0.15 0.12 0.14. 0.10
Table 4 Ash chemical composition (wt?? dry), and ash sintering test results (Sintno = sintering degree at 750 OC, 1 bar of steam, Sintsso= sintering degree at 850 "C, 1 bar of steam). Sample
SiO2Al2O3 Fez03CaOMgOKzONazCTiOz SO3 PzO CI S h t a Sintss
#1 Wheat #2 Wheat #3 Wheat #4 Wheat #5 Barley(W)
38.5 29.9 27.8 16.3 34.2
5
0.2 0.2 0.3 0.2 0.6
0.3 0.2 0.4 0.6 0.5
9.1 10.4 10.4 15.4 12.7
3.8 4.1 4.5 4.1 3.2
27.7 32.5 34.9 41.0 30.1
0.4 0.5 0.4 0.4 2.4
0.03 0.03 0.03 0.03 0.05
6.2 6.0 5.2 7.5 5.0
4.6 3.9 3.7 1.8 5.0
#6 Grass 29.9 0.1 #7 Barley(S) 32.1 0.2 #9Wheat98/rib59.9 1.8
0.2 8.3 3.8 36.1 0.3 0.03 5.0 5.7 0.3 11.9 2.7 30.1 2.4 0.03 4.7 3.4 1.1 7.7 2.5 14.5 0.5 0.17 3.2 3.9
#11 Wheathib+
0.3
70.6 0.3
8.1 3.3 9.0 0.5 0.03 2.7 4.6
3
3
IO**(u) **(u) 11 ** *** 12**(*) *** 12* * 6. **(*) *** 3 11 **(u) *** 13 **(u) *** 4. n.d. *(*) 1 0. n.d. ( 0 ) 4
n.d.: not determined, u: unburned carbon present BEHA VIOUR OF POTASSIUM Ih' GASIFICATION
The behaviour of the most abundant alkaline metal in straw, i.e. potassium was studied during gasification in the thermobalance. Another abundant element is silicon. Potassium and silicon react easily to form potassium silicate (glass), which is waterinsoluble. The objective of this task was to detect the relationship between watersoluble potassium and insoluble potassium, and also possible evaporation of potassium during gasification. In the tests, the gasification was stopped at certain conversion levels in the thermobalance. Afterwards, the samples were extracted by water to analyse the water-soluble potassium and the total potassium (the total alkaline metal analysis was carried out by neutron activation analysis MAA). Steam, C02 and their mixture were used as gasification agents in these tests. The difference between the total potassium content and water-soluble potassium described the amount of insoluble potassium, which is in form of potassium silicate. The results of this study are summarised in Figure 2. The results were calculated so that the amount of potassium was normalised to one either in the char (immediately after pyrolysis, i.e. at 0% char conversion), or in the original straw. Hence, the watersoluble potassium and the total potassium were compared, as shown in Figure 2. 126
Danish Wheat straw 95 soluble potassium, potassium in char after pyrolysis= 1 gasifcation in TG at 800°C
1 Y
0.8 c02+H20
t
H20.total K (INAA)
p o.6 0.4 0.2
0
0
20
40 60 Char Conversion
80
100
Danish Wheat straw 95 soluble potassium, potassiumin straw gasifmtionin TG at 800°C
=1
1 Y
H.20 (repeat)
0.8
0
.a 0.6
Ep
WO,total K (INAA)
0.4
0.2
0
0
20
40 60 Char Conversion
80
100
Fig. 2 The amount of potassium in straw char as a function of char conversion. Symbols without line: amount of water-soluble potassium (normalised values); sym-bols with the solid line: amount of total potassium. In the upper graph, potassium is de-noted as 1 in char after pyrolysis; in the lower graph in straw dry matter, respectively. Accordingly, it shows that the repeatability of the determinations was acceptable, and no dependence on the gasification agent was observed. The amount of soluble potassium was relatively constant to the conversion level of 50%, after which it decreased to the value of about 15% of the origin. The total potassium content decreased by about 10% as a fimction of char conversion The results indicated that the major part of potassium had reacted to the water-insoluble form, i.e. potassium silicate, and this reaction started after the char conversion of 50%. In addition, the total potassium level apparently was reduced during pyrolysis: at the char conversion of 0%, the total potassium content was 10% and the soluble potassium content 20% lower than that in straw, respectively. This indicated that during pyrolysis, a part of potassium would have been reacted to insoluble potassium, and moreover a part was evaporated during pyrolysis. However, due to the relatively large scatter of the measuring points, a more detailed study is needed to venfy this observation.
127
BENCH-SCALE FLUIDISED-BED GASIFICATION TESTS
The behaviour of straw ash was studied further in a bench-scale atmospheric-pressure fluidised-bed reactor (AFB). This study completed the test series reported in [8]. The Danish wheat straw -95 (Tables 1 and 2) was used as the fuel in the 7 set points AFBB24-3 1 (Tables 5 and 6). In these test series, the effect of bed material mixtures, a different limestone grade, freeboard cooling and fuel additives were tested. A selection of Danish wheat straws was also tested (set points 9811 - 98/5, Table 7). Each test was camed out in one day. TESTING OF DIFFERENT BED MATERIALS AND ADDITIVES Previously [8], different bed materials (A1203,dolomite and coke) were tested, and this work was continued by carrying out four additional tests, in which another limestone and a dolomite as well as mixtures of limestone and sand were used. Data for these tests are given in Table 4. The bed material mixtures were tested at two set points: AFBA324-25. Two mixtures (50/50 wt?h basis) were used: a) sand and limestone Parfil, b) sand and dolomite Myanit. The test temperature at these set points was 800 "C. As soon as the temperature signal deviated from the steady state indicating poor fluidisation, the test was stopped. The test periods are given in Table 5. However, in these tests no large agglomerates were found in the sampled bed material. Still the bed contained agglomerates having diameters of a few millimetres, which were too large to be fluidised. This caused temperature gradients in the bed. The amount of freeboard deposits was also somewhat hgher in these tests than at the following set points. The visual inspection showed that the deposits contained agglomerates formed by molten ash, similarly to the previous test series [8]. Table 5 AFB test runs with straw -95 carried out with different bed materials. Set point Feedstock Bed material T(bed, average), "C T(freeboard), "C C-conversion to dry gas, wt% mg freeboard deposits/ g feedstock in the freeboard
AFB/B24 AFB/B25 AFBh326 Straw Straw Straw Limestone + sand Dolomite+ sand Limestone (Orsa) 800 800 800 800 800 800 80.7 80.0 75.1 1.7 1.4 1.9 some molten ash some molten ash some molten ash particles present particles present particles present 125 170 252 Total test time, min (1 15) 160 252 Initiation of bed sintering, min
In order to compare the effect of limestone type, one test (AFB 26) was carried out with Orsa limestone, which is of somewhat harder grade than Parfil. Thls test resulted in a similar sintering behaviour as at set points AFB/24-25. The aim of the test series summarised in Table 6 was to study whether the ash sintering problems of straw gasification could be prevented by using special additives. Two additives were selected: kaolin and magnesium oxide (MgO). The reason for using these materials was that they are known to affect the ash melting behaviour favourably. Ratherhgh 128
Table 6 AFB test runs with straw -95 carried out to study the effects of additives. Set point Feedstock
AFBiB28 Straw + kaolin pelletised Bed material Dolomite T(bed, average), "C 830 830 T(freeboard), "C C-conversion to dry gas, wt% 84.9 mg freeboard deposits/ g fed feedstock 0.7 Total amount of feedstock, g 3025 used bed material, g 420 240 Total test time, min Initiation of bed sintering, min --
AFBiB29 Straw + MgO pelletised Dolomite 830 830 89.2
AFBiB30 Straw + kaolin Dolomite 830 830 86.1
AFBh331 Straw + MgO Dolomite 830 830 87.5
< 0.5 3055 420 240
0.6 3070 420 240
0.6 3215 420 240
--
--
--
amounts of additives were used in these preliminary tests, which were not yet focused on optimising the use of additive. The additives were added as fine powder to ensure good mixing with the fuel. They were mixed with the straw feedstock in two ways: First, they were added among the powder fuel and the mixture was pelletised. Another way was to add them directly into the feeding tank among the crushed feedstock. The particle sizes of the additives were < 20 pm for kaolin and < 70 pm for MgO. The amount of additive used was the same as the ash content in the straw (4.8 g in 100 g fuel). However, during the pelletising and crushing procedures, material losses took place to that extent that the final amount of kaolin was 55% of the amount added, and 70% of MgO, respectively. When the additives were mixed in the feeding tank, there were no losses. In total four set points were carried out in AFB/28-31. The temperature in these tests was higher, i.e. 830 "C,than in the previous set points of straw gasification. Both additives (kaolin and MgO) proved to be effective: all tests were finalised without any signs of sintering in the bed or without deposits in the freeboard. Only small amounts of loose ash were collected from the reactor wall. The mixing procedure seemed to have no influence on sintering. A difference was observed, when the MgO additive was mixed during pelletising: the ash material collected from the freeboard contained significantly less carbon compared to addition into the tank. Probably, h s additive catalysed the reactivity of carbon when mixed intimately with the biomass material. The degree of agglomeration of the bed samples was determined by sieving the samples. The results are shown in Figure 3. Accordingly, when compared to the original bed material the intensity of agglomeration can be seen as an increase in the particle size. At the first two set-points B24 and B25 the original bed material was a mixture of sand with a particle size of 0.60 - 0.71 mm and dolomite or limestone of 0.71 -1 .OO mm in size. At all other set points, only dolomite or limestone of 0.71 - 1.00 mm was used as the initial bed material. The bed agglomeration was clearly measurable at set points AFBiB25 and AFBiB26, whde at AFBh327 the amount of larger particles was only a few percents. At the other set points no measurable amounts of bed agglomerates of >1 mm in size were observed. In all tests, the bed particle size decreased significantly. T h s is due to the easy attrition of limestone and dolomite.
129
100
80 60
40 20 n
AFB set point
Fig.3 Particle size distribution of bed materials and bed samples after the AFB setpoints. GASIFICATION OF DIFFERENT DANISH STRA WS
The gasification behaviour of different straws was also studied in AFB. The aim of these tests was: a) to confirm laboratory findings (themobalance tests) concerning the ash sintering behaviour, which sets limits to the maximum gasification temperature) and b) to assess the gasification behaviour and ash sintering/deposition tendency in a 111scale fluidised-bed reactor. Myanit dolomite was used as the bed material at all set points of this test series. Particle size of the bed material (0.71 - 1.00 mm) and the superficial velocity was selected so that good fluidisation was maintained. The operation conditions and the main results of the gasification tests are presented in Table 7. The first set point AFB98/1 was carried out with wheat straw JE3 2-3 at about 800 "C.The test run was successful and no signs of ash-related problems were met during the operation: bed temperature was constant and no signs of pressure drop fluctuation were seen. The bed material after the test contained only small agglomerates, and the freeboard and the gas outlet pipe were practically clean. Only some dust was deposited on the eeeboard wall, which contained some small agglomerates consisting of fines from bed and ash glued by molten ash. Since there were no problems at the first set point, the second set point AFB98/2 was carried out with the same straw at higher temperature, at 830 "C. This test was as successfid as the first one. No signs of ash-related problems were met during the operation. The same type of small size agglomeration was observed in the bed, as in the first set point. The freeboard deposit was also similar to the above. Set point AFB98/3 was carried out with straw #7 (Barley JB6) at 800 "C. This test run was not as successhl, as the pressure of the reactor started to rise after one hour operation and consequently, the operation had to be stopped earlier than planned. The removed bed material was fairly clean, as were also the freeboard and the gas outlet pipe. The blockage that caused the increase in reactor pressure was detected in the gas line after the cyclone.
130
Table 7 summary of operation conditions in the AFB tests with different straws Set point Feedstock
9811 M , JJ3 2-3, Wheat 800 805 5.5
T(bed), "C T(freeboard), "C Tars, glm'n amount of freeboard deposits,mglg of 0.7 fed feed-stock outlook of deposits small porous agglomerates' 240 Total test time, min Initiation sign of bed sintering, min Bed sinters after no large the test agglo.merates', small agglomerates of the same size as bed particles*
9812 M , JB 2-3,
Wheat 830 830 5.7
9813 9814 #7, JB6 Barley Danish Wheat s.-97 800 800 805 805 4.6 7.5
1 .o
5.8
1.3
9815
Danish Wheat s.-97 770 770 9.4 < 0.5
small porous
small porous agglomerates, char particles agglomerates' agglomerates' char particles 240
932
no large agglomerates, small agglomerates of the same size as bed particles
252 252
240
no large very large agglomerates, agglomerates small agglo- *** merates of the same size as bed particles
larger char agglomerates and small molten ash particles (only a few small agglomerates) Test run was stopped as the pressure of the reactor was increasing due to a blockage in the gas outlet pipe. Agglomerates comprise bed and ash particles glued by molten ash.
'
Set point AFB98/4 was carried out with Danish wheat straw-97 at 800 "C. No signs of fluidisation loss were met during the operation (bed temperatures and pressure drop remained stable). However, after the test there were very large agglomerates in the bed, but the freeboard and gas outlet pipe were practically clean. The freeboard dust contained small ash agglomerates. The bed agglomerates were quite porous and probably located at the enlargement section of the bed so that they did not affect fluidisation. Because the temperature of 800 "C resulted in ash-related problems for Danish wheat straw -97, the next test run AFB98/5 was camed at a lower temperature of about 770 "C. This test was more successful. The freeboard and the gas outlet pipe were practically clean. The bed contained some large agglomerates consisting of char particles adhered together. In addition, in the bed sample there were some single glassy particles, which seemed to be formed solely by molten ash. A few agglomerates containing bed particles were also observed in the bed. CFB GASIFICATION TESTS WITH STRAW
Gasification tests were canied out with the atmospheric-pressure Circulating FluidisedBed (CFB) gasification Process Development Unit (PDU) of VTT Energy [lo]. Two wheat straw feedstocks were also used in the test runs, i.e., Danish wheat straws -95 and -97. The composition of these two straws were fairly close to each other, both having very high potassium and chlorine contents, which can be considered typical of dry harvesting seasons (Tables 1 and 2). Both straws had also a rather high reactivity, but the ash sintered strongly already at below 800 OC in the laboratory tests.
131
RESULTS
The straw gasification trials had two main goals: a) to find such gasification conditions, at which wheat straw could be gasified without ash-related problems, and b) to study the removal of potassium and chlorine derived from straw. According to the experiences from a CFB gasifier it is also possible to use such a biomass fuel that on the basis of laboratory characterisation seem to be very problematic. Under optimum conditions a carbon conversion efficiency of 95% was reached without any signs of ash-related problems. The tar content was also clearly lower in straw gasification than in wood gasification at the same temperatures. However, the achievement of the optimum conditions requires controlled operation of the gasifier, otherwise sintering and agglomeration problems are met, and moreover, carbon conversion is clearly lower [lo, 121. CHARACTERISATION OF ASH SAMPLES FROM THE CFB TESTS
Samples were collected from bottom ash, cyclone ash and hot filter dust in some of the CFB test runs carried out with straw. The samples were analysed by light microscopy and electron microscopy. With the first method, an overall inspection was carried out in order to monitor agglomerates and their form. Afterwards, cross-sections were prepared for the electron microscopic analysis (SEM), in which the CCSEM (computer controlled electron microscopy) technique was applied. These analyses were carried out at VTT Manufacturing Technology. In this analysis, sample particles were analysed for their size and chemical composition. In almost all CFB gasification tests, some relatively small (1 - 3 mm) agglomerates were formed in the bed, comprising bed material particles glued by molten ash, the type and form being dependent on the process conditions. These agglomerates were removed from the gasifier together with bottom ash. From the ash samples, cross-sections were prepared and analysed by SEM. In the analysis, particle sizes and their chemical composition were determined. In the analysis of chemical composition, the following elements were measured: Si, Al, Fe, Ca, Mg, K, Na, Ti, P, S, C1, (Mn). The summary results of this SEM analysis are shown in Figure 4. The composition measured for particles was compared with the chemical composition of ash of the feedstock straw (also indicated in Figure 4). According to the results, the most abundant substance in bottom ash was silicon, the amount of which was locally much higher than that of straw ash. The potassium content was quite equal from particle to particle. The dust particles collected from the 20d cyclone contained less silicon than the bottom ash while the potassium and calcium contents were clearly higher. The chlorine content was also higher than in bottom ash. The particle sizes ranged 1.5 - 11 micrometers in the 20d cyclone ash as measured in SEM. In the dust collected from the hot filter, the particle sizes ranged 0.2 - 4 micrometers, respectively (particle sizes were not measured for bottom ash due to their large sizes). In these particles, potassium, calcium and chlorine were clearly concentrated, and there seemed to be no dependence between particle size and composition. When potassium was plotted against chlorine, a linear correlation was found for potassium and chlorine, indicating the presence of potassium chloride in the particles (Figure 5). This observation was in accordance with bulk analysis carried out for potassium and chlorine in the 2nd cyclone and in the hot filter dust (Figure 6 [lo]). The amount of water-insoluble alkaline metals also decreases from bottom ash to the 2nd cyclone ash and hot filter dust, as indicated in Figure 6. In 132
bottom ash 100% is in form of insoluble alkaline metal, in 2"dcyclone ash about 50% and in hot filter dust 30%.
particle
--
s a
E
particle dp. urn
1 Hot filter dust I
1 OD ._
1203 2
80 60 40 20
O O O O O O O O O O O O O O O O O ~ ~ N ~ ~ ~
8 particle dp. urn
$
Fig. 4 The composition measured for particles of bottom, 2nd cyclone ash and hot filter dust samples taken from a CFB test run of wheat straw. For comparison, the chemical composition of ash of feedstock straw is shown as red bars at the right ends of the diagrams (note: for better visualisation, the array of the graphs in bottom ash diagram deviates from that of the other diagrams).
133
D K20
K2O 40
30
u 20 10
0 0
I 10
20
30
40
50
I
0
K2O
Fig. 5 Correlations between chlorine and potassium in the particles of bottom ash, 2nd cyclone ash and hot filter dust (CFB).
134
hot filter feedstock PVEL FEEDER ADDITIVE
bulk:
bulk:
average K content: 2 % average Na content: 0.02 %
average K content: 17 Yo averaie Na content: 0.5 %
average K content: 13 % average Na content: 0.7 %
average K content: 10 Yo average Na content: I%
!Gig.6. Deviation of alkaline metals in CFB gasification of wheat straw (average K, Na content indicates the bulk analysis of ash samples); two CFB tests 40 & 49, the bars indicate the fraction of water insoluble K and Na (i.e. silicates). CONCLUSIONS The following conclusions were drawn fiom the results: 1. Characteristic data on ash behaviour can be obtained by laboratory tests and used for planning, e.g., run conditions for tests with pilot or PDU (process development unit) equipment. 2. With certain fuels (such as alfalfa and willow), ash sintering can be much stronger under pressurised conditions than under atmospheric pressure. This phenomenon seems to be related to the silicate content of ash: when the silicon content was low ash sintering was stronger under pressurised conditions than under atmospheric conditions. 3. The major part of potassium reacted to a water-insoluble form, i.e. potassium silicate, after the char conversion of 50%. 4. The use of kaolin or magnesium oxide additives had a great preventive effect on bed agglomeration and freeboard deposit formation in the bench-scale bubbling fluidised-bed gasification tests carried out with straw. 5 . The differences between Danish wheat straw qualities were relatively small: almost all tested straws had a rather problematic ash sintering behaviour. 6 . The Circulating Fluidised-Bed gasification process was not as sensitive to ash sintering as the small-scale bubbling fluidised-bed reactor. The amount of potassium chloride increased from bottom ash to hot filter dust.
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ACKNOWLEDGEMENT
The support of the following institutions and companies is gratehlly acknowledged The Progas R&D Programme of VTT; the Finnish Liekki 2 Research Programme; Tekes, the National Technology Agency of Finland; Reatech, Denmark; Elkraft, Denmark; and Carbona Oy, Foster Wheeler Energia Oy, Imatran Voima Oy (Fortum), and VTT Energy, Finland. REFERENCES
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Miles P. E. T. R., Miles, T. R. Jr., Baxter, L. L., Bryers, R. W., Jenkins, B. M. & Oden, L. L. (1995). Alkali deposits found in biomass power plants. Summary Report. Golden, CO: National Renewable Energy Laboratory. 82 p. + app. 35 p. Baxter, L. & DeSollar, R. (eds.) (1995). Application of advanced technology to ash-related problems in boilers. New York Plenum Press. Nordin, A., Ohman, M., Skrifvars, B.-J. & Hupa, M. (1995) Agglomeration and de-fluidization in FBC of biomass fuels - mechanisms and measures for prevention. In: L. Baxter and R. De Sollar (eds.). Application of advanced technology to ash-related problems in Boilers. New York: Plenum Press. Bryers, R. W. (1996). Fireside slagging, fouling, and high-temperature corrosion of heat-transfer surface due to impurities in steam-raising fuels. In: Proc. Energy Combust. Sci.,vol.22, pp. 29 - 120. Miles, T. R., Miles, T. R. Jr., Baxter, L. L, Bryers, R. W., Jenkins, B. M. & Oden, L. L. (1996). Boiler deposits from firing biomass fuels. Biomass and Bioenergy, vol.10, no. 2 - 3, pp. 125 - 138. Moilanen, A . ( 1993). Studies of peat properties for fluidised-bed gasification. Espoo: VTT. 34 p. + app. 35 p. (VTT Publications 149). Kurkela, E., Laatikainen-Luntama, J., Sthlberg, P. & Moilanen, A. (1996). Pressurised fluidised-bed gasification experiments with biomass, peat and coal at VTT in 1991-1994. Part 3. Gasification of Danish wheat straw and coal. Espoo; VTT. 41 p. + app. 5 p. (VTT Publications 291). Moilanen, A., Kurkela, E., Laatikainen-Luntama, J. (1999). Ash behaviour in biomass fluidised-bed gasification. In: Gupta et al. (eds.). Impact of mineral impurities in solidfirel combustion. New York Kluwer Academic I Plenum Publishers. Pp. 555 - 567. Ssrensen, L. H., Fjellerup, J., Henriksen, U., Moilanen, A., Kurkela, E. & Winther, E. (2000). An evaluation of char reactivity and ash properties in biomass gasification. Fundamental processes in biomass gasification. Confidential project report. REA11312000, ReaTech. 91 p. + 3 app. Kurkela, E., Moilanen, A. & Nieminen, M. (1999). CFB gasification of biomass residues for co-combustion in large utility boilers - studies on ash control and gas cleaning. In: Sipila, K. & Korhonen, M. (eds.) Powerproductionfiom biomass HI. Espoo: VTT. Pp. 213 288. (VTT Symposium 192). Levin, E. M., Robbins, C. R. & McMurdie, F. (1985). Phase diagrams for ceramists. Vol. I. 5th print. Washington: American Chemical Society. 601 p. International application published under the Patent cooperation^ Treat (PCT). Gaslfication of biomass in a fluidised bed containing anti-agglomerating bed material. Int. Appl. No. PCT/F199/00695. WO 00/1115,2 March 2000 (02.03.00).
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Design of a Biomass Gasification Gas Sampling Systern J.M. Sinchez, E. Ruiz, D. Cillero, J. Otero, A. Cabanillas CIEMAT, Fossil Fuels Department, Avenida Complutense 22, 28040-Madrid (SPAIN)
ABSTRACT This paper describes the main characteristics of a biomass gasification gas sampling system designed and built-up at CIEMAT to characterise the emissions from an atmospheric circulating fluidised bed gasifier when doing research on coal and biomass gasification. This system is primarily used to monitor the performance of the gasifier and to provide data to study the influence of different parameters on gas product. Different types of biomass are going to be gasified. In a first phase they will be for example oak chips, "orujillo" (residue from the olive oil- making industry), etc.
INTRODUCTION Reliable sampling and analysis of products from biomass gasification are essential for the successful process development and economical operation of commercial gasifiers. One of the most difficult analytical tasks is to characterise the emissions from the thermal conversion of biomass. When biomass or other solid fuels are thermally converted into a gaseous fuel to provide energy, in addition to valuable components such as H2, CO and inert gases like C02, H20 and N2, gaseous impurities of different types are also formed. Although the formation of gaseous contaminants in biomass conversion is highly dependent on the nature of the biomass feedstock, the effluents can be divided into the following categories [ 8 ] : (1) Particles: char, ash and gasification additives such as sand, limestone or dolomite resulting in plugging and abrasion of downstream equipment. (2) Low and high molecular weight organic compounds (tars), causing blinding of ceramic filters, pipe clogging and soot formation during gas combustion. (3) Nitrogen containing impurities (NH3, HCN), that convert into NO, during gas combustion. (4) Sulfur containing compounds (H2S, COS) that release SO2 during gas combustion. ( 5 ) Other impurities such as HCl, alkali metals and some vapour-phase metals, causing corrosion of the gas turbine, heat exchangers or poisoning catalysts of the catalytic tar crackers. Within this framework a sampling and analysis system that has been designed to 137
measure and characterise the composition of actual biomass gasification gases from a circulating fluidised bed gasifier is presented here. This system is primarily used to monitor the performance of the gasifier and to provide data to study the influence of different parameters on gas product.
OBJECTIVE The aim of the sampling system presented here is to take samples from a circulating fluidised bed gasifier at pilot plant scale (200 Nm3/h of gases are released). This one has been extensively operated as a combustor and it has recently been revamped to operate as a gasifier. Different feedstocks like oak chips, “orujillo” (a residue from the olive oil making industry), etc. will be gasified in the first studies. Initially, the main target is to gain experience in sampling and analysing actual fuel gas produced in the gasification reactor and provide staff with operating experience. Following that, the next objective is to evaluate the gasifier performance by means of the analysis of the effluent gases. The ultimate objective is to successfully determine the influence of different operating parameters (aidbiomass ratio, temperature, residence time, etc.) on the quality of the fuel gas produced (concentration of valuable components such as H2 and CO, minimisation of tar formation, etc.). Due to the very different nature of biomass feedstocks, with a wide variety of composition and chemical species, it is expected that each kind of biomass will be the source of particular experimental problems, with its own unique characteristics. Hence, it is likely that the composition of the fuel gas and the formation of gaseous pollutants will be highly dependent on the feedstock. Therefore, the sampling system should be in some sense tailor-made and periodic checks of the reliability and accuracy of the sampling system will be made in order to enhance it.
OVERVIEW OF BIOMASS GASIFICATION SAMPLING SYSTEMS IN LITERATURE A wide range of power generation technologies based on the gasification of biomass are available or under development. These technologies have three process steps, the gasification reactor, cleaning of the fuel gas and the power generation unit (gas engine, gas turbine, etc.). The fuel gas from all gasification reactors contains particulates and heavy hydrocarbons which must be removed to avoid damaging the engine or incurring an unacceptable level of maintenance. However, at present, no unified method, in the form of a standard or recommended procedure, exists for monitoring and evaluation of the performance of gasifier systems. In the EU project “Developmentof a standard procedure for gas quality testing in biomass gasifier plantlpower generation systems“ [4], a draft standard procedure for producer gas quality testing was developed. Two Guidelines for producer gas quality testing were evaluated. The methods reviewed are those from UNDP/World Bank, “Guidelines for Field Monitoring of Small Scale Biomass Gasifiers” and the Biomass Research Group of ITT Bombay, India, “Biomass guidelines for field testing and performance evaluation”. From the point of view of determining the biomass gasification system efficiency and monitoring the performance of the gasifier, parameters to be measured are: gas composition (02, C02, H2, C&, C,H,, CO, Nz, tar, particulates, water vapour, etc.), gas temperature, gas flow rate, gas pressure, and calorific value. 138
PERMANENT GASES Permanent gases are either continuously measured on line by means of specific monitors or discontinuously by a gas chromatograph (GC). For example, the on-line measurement of CO, COz. and Cfi content of the product gas is usually based on the absorption of infrared (IR) radiation, the measurement of Hz is accomplished by thermal conductivity detectors (TCD), whereas the 0 2 content is determined by a magnetic susceptibility (MS) device. Both systems, on-line determination and gas chromatography, have advantages and disadvantages. The GC doesn’t need a special sample-preparing unit and gives the possibility to introduce additional columns for tar analyses. An advantage of the IR-, TCD- and MS techniques is the very short analysing time. This enables to follow on line the change of the concentration of the gases. However, all the components cannot be analysed by a single technique. Therefore, a combination of IR and TCD (thermal conductivity detector) and MS (magnetic susceptibility) must be applied when a very short analysing time is required. Fourier Transform Infra Red (FTIR) Spectroscopy is a promising and versatile technique for gas analysis which lately has moved from the laboratory to industrial applications such as emission monitoring of combustion and gasification plants [2]. The single most important advantage of the FTIR is its capability to analyse in-situ virtually all gas species of interest in flue and fuel gas applications. In this way potential sampling artefacts can be avoided. Analysing the gas composition by means of the Orsat analyser is in principle a very good option because of its simplicity and high accuracy, but the analysing time of about one hour is too long for a good evaluation of the producer gas composition. The Orsat analyser is therefore only recommended to be used in remote areas where no electricity is available or the infrastructure is not adequate for installing a gas chromatograph.
PARTICULATES AND TAR The tar problem is conceived as one of the most important technical barriers for the penetration of the biomass gasification technology in the power markets. A number of different sampling and analysis methods have been developed by manufacturers and various institutes working in this field to determine the level of particulates and tar in the raw fuel gas or exiting the gas cleaning systems. Knoef and Koele [7] have carried out a survey of tar measurement protocols as a specific action of the IEA Bioenergy Gasification Task. The detailed information is documented in the Biomass Technology Group web site {http://www.btgworld.com/ }. The most relevant aspects of some of the contributions to the tar protocol survey are summarised in Table 1. The EPA Method 5 for sampling particulates emissions from flue gas is the basis for most gasifier sampling trains. Modifications have been necessary because of the higher tar and particulate loading of gasifier streams. Usually, sampling techniques have been used to simultaneously measure tar and particles to ensure similar process conditions. Isokinetic sampling should prevail during sampling. Common elements to different isokinetic sampling systems are: A heated filter (glass fiber, cellulose, quartz-fiber, ceramic) for trapping the dust particles and a condenser for trapping the tar ([91, [61, [ 141, [ 5 ] ) .
139
P
VlT-hot side tar sampling
TPS method
IVD-dry sampling method. IVD Universitat Sttutgart TOA-Tar Online Analyzer. D/D Universitat Sttutgart SPA tar samplinglanalysis method. Royal Institute of Technology __ KTH, Sweden
am3
GC-FID
Component analysis (GC-FID)
Component analysis (GC-FID)
Gravimetricanalysis. Component analysis (GC-FID) Total condensable hydrocarbons
Offline. Tar Gravimetric analysis, solvent extraction (DCM, evaporation acetone) Component analysis: GC-FID, GC-MS (solvents, EM, acetone) Offline. Tar Gravimetric analysis, solvent extraction (water, evaporation anisole, acetone)
Analysis Applied determination methods Offline. Tar Gravimetric analysis, solvent extraction with evaporation DCM Component analysis: GC-FID
Adsorption (glass fiber filter, Offline. Tar extraction. pore size 1 pm) In situ. 0.5 min. 310-33O"C, 50mg/Nm3- Tar only. Condensation On line. 50 gNm3tar. Isothermal, sampldmain stream flow Adsorption (silica based Offline volume ratio 0.0396,sampling time 10-60 aminophase) s, temperature, 250-3WC, <1 mg/Nm3tar 100 Offline. Tar Isothermal, sampldmain stream flow Absorption (acetone) volume ratio O.OlS, sampling time 15 extraction (solvent: min.,temperature, 10-9WC, 4 mg/Nm3DCM + water) tar ~ so000 g / ~ m Isokinetic, isothermal Absorption (neutralhisidacid Offline. Tar aqueous solutions) extraction.
g ~ r particulates n ~ Isothermal 60 min., 500°C
Protocol name Description Tar separation from gas sample BTG- World bank monitoring Isokinetic, isothermal, sampldmain Filtration (cellulose, cotton method stream flow volume ratio 1% sampling wool) time 1-3 h, temperature, 10-5WC. 1 m ~ m 3 - 1 0g / ~ mtar, ~ I mgmm3-i0 g / ~ m particulates ~ DTI. Dansk Teknologisk Institut Isokinetic, sampldmain stream flow Condensationladsorption (quartz volume ratio 0.5%, sampling time 1-2 h, wool, XAD-2, PUF) temperature, 20-5WC, 20mgNm3-200 g / ~ m tar, ~ 20 m ~ m 3 - 1m~m 3 particulates ETH-Verenum method Isokinetic, isothermal, sampldmah Condensation, adsorption stream flow volume ratio 1%. sampling (cellulose), absorption (anisole) time 2-5 min., tern erature, 20-4WC, 50 m ~ m 3 - g/Nm 1 ~ tar, 10 m ~ m 3 - 1 0
Table 1. Tar measurement protocols.
Stahlberg et al. [ 131 describe the development and rationale of non-isokinetic sampling systems when high flow rates are not possible and when particulate measurement is not to be performed simultaneously to gaseous species determination. There are interesting alternatives to the traditional tar analyses which are based on absorption in a solvent (impinger sampling trains). In this way, a quasi continuous tar quantification method has been developed at the University of Stuttgart [ 113. It is based on the comparison of the total hydrocarbon content of the hot gas and that of the gas with all the tars removed. Hot gas from the gasifier is led directly to the analyser. Hydrocarbons are measured with a flame ionisation detector (FID). It considerably simplifies the experimental procedure compared to other methods described in literature. Sampling and analysis time is just two minutes. Another approach to reduce sampling and sample separation time has been designed by the Royal Institute of Technology (Sweden)[3]. It is based on solid-phase adsorption (s.p.a.) and suitable for intermittent trapping of tar compounds ranging from benzene to coronene. Using eluotropic elution, adsorbates are selectively desorbed into aromatic and phenolic fractions and then determined by gas chromatography with flame-ionisation detection. The sampling step allows collection of one to three samples per minute, compared with one or two samples per hour using conventional cold trapping, and correspondingly more information is obtained. The existing diversity of methods makes the comparison of operating data from different sources very difficult. Some of the differences highlighted are:[ 101 (1) Probe issues: Isokinetichon-isokineticsampling, sampling temperature, material. (2) Tar condensation and collection issues: condenser, traps, solvent nature, condensation, collection temperature. (3) Treatment of collected tars before analysis. (4) Analytical techniques used to characterise tars: gas chromatography, liquid chromatography,total organic carbon. In order to reach consensus on a measuring protocol for determination of tar and dust content in producer gas flow streams, the members of the IEA Bioenergy Gasification Task, the European Commission and the US DOE called a joint meeting in 1998. As a result of the meeting, two measurement protocols have been initiated to meet the described needs. The protocols are meant for distinctive gasifier application ranges: (1) Small scale gasification [l]: This protocol will be used for small-scale, fixed bed, engine based systems. The working group decide that a method similar to that used for sampling PAHs in stack gas would be appropriate. (2) Large scale gasification: This protocol is recommended for use in sampling of large-scale atmospheric and pressurised gasifiers. However, there is also a wide experience in the use of the described methods in sampling of small-scale updraft and downdraft gasifiers. The protocol is based on VTT standard methods for constructing sampling lines, sampling of tars and analysis of tars [ 121. After the two protocols were drafted, it seems to be great similarities between them and it should be examined the possibility to merge them into one tar protocol. Remarks on which consensus exist are the importance of the isokinetic sampling of tar and dust, the possibility of tar polimerisation reactions, the incorporation of
141
gravimetric determination for those high molecular compounds, the sampling lines should be as short as possible, the definitions of tar as hydrocarbons with molecular weight higher than benzene, etc. However there are points in which no consensus exists such as the importance to include particulates determination in the protocol, sequence tar-particulates or reverse, need of an organic solvent during sampling, the definition of the most promising solvent to absorb tar (up to now it has been decided to use dichloromethane (DCM). Another alternatives solventhemperature combinations are: isopropyl glycol at -30"C, acetone at -79°C. Other solvents which have been tested for tar absorption include toluene, anisole and even water), condensation temperature, how to maintain isokinetic conditions, sampling duration and sampling flow, whether there is a relation between the definition and the application of the gasifier, need for GC analyses, sampling duration in relation to the method applied, etc. At present, efforts are being focused on the development of shortcut methods for the measurements of tars. The purpose of these revisions are to simplify the sampling and characterisation procedure (both in time and devices), but still to provide accurate and reproducible information reducing safety hazards and minimising time and cost of operation.
SAMPLING PROCEDURE The sampling strategy proposed is illustrated in Figure 1. The scope of the application is the sampling and characterisation of permanent gases and light and heavy organic compounds from biomass gasification effluents. The potential use, by slightly modifying the system, to trap inorganic pollutants (sulphur gases, nitrogen-containing impurities, chlorine compounds and vapour-phase metals) generated when coal or wastes are gasified may be occasionally evaluated, if needed. The sampling train is mostly based on the system used by VTT in gasification research [13].The reliability of VTT's system has been extensively proved. At present the sampling system presented here is being built up and set up. In the foreseeable future it will be fully operated and then the capabilities of the system will be studied. Since the heterogeneity of gasification feedstocks results in a formation of a wide variety of products of different chemical nature it is envisaged that the sampling system will have to be enhanced and up-dated on a short term basis. Therefore, there is room for improvement in sampling methods in order to meet all the requirements of the different gasification applications. The principle of sampling protocol is described in the following paragraphs. A slipstream of the fuel gas is taken out and diverted to the sampling system by means of an Inconel probe equipped with a sintered stainless steel filter in the tip to remove particles. The probe is placed co-currently to the gas flow and is positioned at a bend in the flue gas duct, after the particle filtering system. Gas temperature at this point is expected to be about 700°C. Downstream of the probe the sampling line is electrically heated to avoid undesirable condensation of organic compounds. Gas temperature is kept at Thus, line blocking is minimised. Heavy polyaromatic temperatures around 300-400°C. compounds (heavy tars, from pyrene to coronene) are condensed at a controlled temperature (in the range from 150-260°C)in a filter trap housed in a forced air circulation oven. The lower the temperature in the filter is, the faster the plugging of the filter is due to a higher deposition of tars. Therefore, the selection of an appropriate temperature in the filter is of paramount importance. It must be chosen as high as
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GAS
EXIT
VOLUMETRIC GAS METER
SAMPLING BAG
FILTER TRAP
Figure 1.Schematicof the gas sampling and analysis system
I
GAS INLET
VACUUI(/PRESSURE
PUMP
COLD TRAP
GAS C H R M T O G R A P H
possible to avoid blocking the filter elements too early, but low enough to prevent thermal decomposition of tars on the filter surface from going on. Hence, temperatures within the range of 150 to 250°C will be used, and usually it will be maintained at 150°C. The filter will plug periodically causing an increase in the pressure drop across the filter housing. The pressure drop is estimated by subtracting the pressure readings taken before and after the filter by means of electronic pressure transducers. These measuring devices are also heated to avoid condensation of tars that could cause misleading pressure readings. When pressure drop reach a pre-set value, the filter element will be replaced. Quantitative analysis of heavy tars is accomplished by gravimetric techniques. After sampling, the filter is cooled down to ambient temperature. The filter element is washed with dichloromethane (CH2ClZ)to extract the organic fraction. Gravimetric analyses are carried out to have a measurement of the total amount of heavy organic compounds. Firstly, the solution is filtered to remove particles and after that the solvent is evaporated. Qualitative characterisation of each chemical species may be discontinuously assessed by coupled gas chromatography-mass spectrometry. The amount of tars obtained will depend on the vapour pressure of individual species at the filtering temperature. Those compounds that have a considerable vapour pressure at the filtering temperature may remain partly in gas phase. Therefore, some improvements should be made in order to obtain reliable and repetitive measurements. Gases leaving the filter may be led directly to a gas chromatograph. In this approach light tars (benzene, phenol, etc.) are kept in vapour phase by accurately adjusting gas temperature and thus continuously analysed on line together with permanent gases (Nz,Hz, CO, COz, C&, etc.). The gas chromatograph is equipped with a thermal conductivity detector (TCD) and a flame ionisation detector (FID). In this arrangement temperature must be kept high (170-300°C) not only along the sampling line but also everywhere inside the chromatograph to avoid tar condensation. Alternatively, gases coming from the filter by-pass the gas chromatograph and head for a cold trap. It comprises an impinger train placed in a cold bath. After sampling the storage of samples until analysis will be carried out according to standards found in literature. If gases are compelled to pass directly through the gas chromatograph, the gases leaving are immediately led through a heated line into the cold trap mentioned in the previous paragraph to collect those compounds that can condense in the line. Thus, heating of the sampling line is no longer required and at the same time damage of upstream pieces of 'equipment is prevented. A vacuudpressure station located downstream of the impinger train is used to ensure a continuous gas flow along the sampling line since actual gases to be analysed are sampled from an atmospheric circulating fluidised bed biomass. Total gas flow rate (from about 5 to 15 NVmin) is measured by means of a volumetric gas test meter. The average concentration of tars will be estimated by the measurement of the amount of tars collected gravimetically (heavy tars) and/or absorbed in the impinger train (light tars), the sampled gas total volume and the sampling time. Downstream of the gas meter there is a three way valve whose purposes are: Firstly, when it is switched to a sampling bag, additional samples of gas are collected. This way, it is possible to analyse, if necessary, inorganic compounds (HzS, NH3, etc.) by specific gas chromatograph detectors (for example flame photometric detector in the case of sulfur compounds). Secondly, it may be turned to feed back permanent gases to 144
the gas chromatograph situated on site. Finally, another three-way valve is used to vent gases into atmosphere. The selection of one way or another will depend on whether gas flow had previously passed or not through the gas chromatograph.
DESCRIPTION OF THE SAMPLING UNIT The sampling system is made up of the following pieces of equipment: (1) Filter trap. (2) Cold trap. (3) Gas metering system and related items. (4) Gas chromatograph. ( 5 ) Control unit. ( 6 ) Gas detection system. A detailed description of every single unit is given in the following.
FILTER TRAP Headline disposable filter elements are used to collect condensed tar at controlled temperature. These elements are manufactured from precise mixtures of borosilicate glass microfibres. The disposable filter elements offer exceptional filtration efficiency (grade designation, 99,9996 removal of 0,l micron particles) at very low pressure drops. The housing is constructed from 316 L Stainless Steel. It has an internal volume of 130 cc. It withstands a pressure of 100 bar (at 200°C) and it can operate up to a temperature of 500°C due to the use of a high temperature gasket. The filter housing is placed inside a Carbolite PF30 forced air circulation oven. A Carbolite type 201 Temperature Controller is supplied as an integral part of the control section of the oven. It is a digital instrument with PID control algorithm. The oven has two ports: One for the tarry inlet gas pipe and the other for the outlet gas stream after condensing of polyaromatic compounds. In addition there are two vents at the back of the oven for the inlet and exhaust forced circulation air.
C O D TRAP The cold trap is the sampling method most often used to collect biomass gasification light tars. The impinger train comprises six bottles connected in series: (1) The first and the sixth bottles are half-filled with glass beads. (2) The second, third and fourth bottles shall contain known quantities of dichloromethane. (3) The fifth bottle is half-filled with glass beads and contains a known quantity of dichloromethane. (4) The first four bottles are placed in an ice bath at 0°C. (5) The two last bottles are placed in an acetone-COz-ice bath at -70°C.
If the aim is to determine other than tar contaminants, such as nitrogen compounds, sulfur compounds,'chlorine pollutants, etc., similar arrangements can be used simply by choosing proper solvents and adjusting carefully sampling temperature.
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Leaks in the impinger train are prevented using leak-free ground glass fittings or any similar leak-free non-contaminating fittings.
GAS METERING SYSTEM AND RELATED ITEMS An Altech vacuum/pressure station (model No. 400-3902) is used to maintain sampling rate and to ensure constant flow from the sampling point through the sampling system at the required pressure. A volumetric gas meter located at the end of the line determines total gas sample volume.
GAS CHROMATOGRAPH A 5890 Series I1 Hewlett-Packard equipped with a thermal conductivity detector (TCD) and a flame ionisation detector (FID) connected in series is used to analyse on line gas composition. The chromatograph has a septum-purged packed inlet. Series 530 pm columns can be fitted to the injector by means of proper adapters. Two columns are connected in series. Depending on the compounds to be analysed it is possible to choose among the following columns: (1) (2) (3) (4) (5)
10 Ft Porapak N, 80/100 GR 5.3 packed column. 6 Ft Molecular Sieve 5%160/80MESH packed column. HP PlotSieve 5A,30mx530pm semi-capillary column. Poraplot Q, 30mx530pm semi-capillarycolumn. HP-5,30mx530pm semi-capillary column.
Two six-port valves are used for gas sampling and column isolation. The column on the second valve can be taken out of the flow line. An adjustable restrictor compensates for the drop in pressure when the valve switches. Typically isolation valves are used when a column can irreversibly absorb some components. Support gases, e.g. carrier gas (helium), reference gas for the TCD (helium), hydrogen and air for the FID, as well as calibrated standard gases (specially prepared calibration samples) are supplied at the required pressure by separate lines. They are placed and stored in a purpose built divided into two separated cubicles (flammable and non flammable gases). The gas chromatograph is commanded by HP Chemstation A.6.03 software. Up to now the preparation of calibration methods to analyse permanent gases has finished. Two different non isothermal methods have been prepared. The first one is meant to be used when gases pass directly through the gas chromatograph. The second one analyses gases collected in sample bags and therefore at a lower pressure. The methods use two packed columns connected in series (10 Ft Porapak N, 80/100 GR 5.3 packed column and 6 Ft Molecular Sieve 5%160/80 MESH packed column). In order to obtain a complete separation of the different compounds, a two level heating and cooling program has been established in the oven zone, coupled with column isolation. Runtime of the chromatographic method is 23.4 minutes. The flame ionisation detector (FID) is used to analyse hydrocarbons (CH, and C3H8).The thermal conductivity detector (TCD) is used to analyse H2, CO, C02, N2 and high concentration of hydrocarbons (CH4 and C3Hg). The polarity and sensitivity of the thermal conductivity detector have been set according to the conductivity response of the 146
different gases to be analysed in order to enhance the integration. Calibrated model gas bottles of known composition have been used to prepare the calibration table. Calibrated peaks and concentration range of each component used to obtain the different calibration curves are summarised in Table 2. The second method has the same characteristics, except for the calibration table. Since the injection pressure when using gas sampling bags is lower, the response decreases and, therefore a new calibration table has to be prepared, mainly due to the non linear nature of the hydrogen calibration curve.
Table 2 Concentration range of the model permanent gases used to prepare the gas chromatographic calibration table. Compound
Concentration range (%) 5-40 1-25 0,02-0,09 37-99 1-5 0,l-5
CONTROL UNIT Gas temperature is precisely controlled by five PID type electronic temperature controllers (TICI, TIC3, TIC4, TICS, TIC6) with phase angle control, so that tar compounds do not condense along the different sampling lines upstream of the impinger train. Each temperature controller monitors the temperature of the sample gas at one point along the sampling arrangement. According to the temperature setpoint the system is heated. Temperature is held in the set value by means of electric heating. Each temperature controller includes an alarm output to avoid gas overheating above alarm set value. In addition to the temperature controllers each heating electric cable is connected to a temperature indicator (TI 1, T13, T14, TI5, TI6). These temperature indicators monitor the temperature at the interface between the heating cable and the tube. If the temperature over-exceeds the set point the electrical heating is turned off. This way, both the heating elements and the tube material are protected against overheating and a longer lifetime of process equipment is guaranteed. Gas temperature at the different points along the sampling system is measured by type-K thermocouples. All the temperature controllers and meters are panel mounted so that they are easily operated from the front of the housing. The oven in which the filter trap is housed has its own temperature controller fitted. It has already been described. Pressure drop across the tar filter trap is measured by means of Kulite HEM-375 Piezoresistive Pressure Transducers. The sensing element is dielectrically isolated "silicon-on-silicon" and exhibits excellent stability and thermal characteristics. They can be operated between -50°C up to 260°C. Pressure readouts at the inlet and the outlet of the filter trap are displayed in two Druck DPI 280 Series Digital Process 147
Indicators. The instruments are designed for panel-mounting.
GAS DETECTION SYSTEM The pilot plant from which sampling gases are taken has its own wall mounting gas detection system which consists of a control instrument installed in a safe area and several flammable sensors and toxic gas sensors mounted remotely from the control unit. Gas sensors are located on the vicinity of those areas considered as the most dangerous (high probability of gas leaks, electrical devices, etc.). Even though, operators of the sampling unit are provided with carbon monoxide portable detectors (Oldham, model OX-TX 12). The instrument triggers an intermittent audible and visual alarm as soon as there is a threshold exceeded (instantaneous, authorised instantaneous maximal value; STEL, short term exposure limit; or TWA, time weighed average). Furthermore, the OX -TX 12 keeps data in memory, the events (alarms, reset) and the measurements for a duration of 120 hours working. So, it is possible to edit statistical charts on a printer using a suitable computer interface.
GENERAL REMARKS AND FUTURE WORK Reliable sampling and analysis of products from biomass gasification are essential for the successful process development and economical operation of commercial gasifiers. One of the’ most difficult analytical tasks is to characterise the emissions from the thermal conversion of biomass. Up to now, although great efforts are being made, standard methods of sampling and analysis have not been defined yet. Within this framework, a gas sampling and analysis system, mostly based on those described in literature, has been designed to measure and characterise the composition of actual biomass gasification gases from a circulating fluidised bed gasifier. At present, the sampling system is being built up and set up. In the future, it will operate fully and then the capabilities of the system will be studied. This system will be primarily used to monitor the performance of the gasifier and to provide data to study the influence of different parameters on gas product. The scope of the application is the sampling and characterisation of permanent gases and light and heavy organic compounds from biomass gasification effluents. The potential use, by slightly modifying the system, to trap inorganic pollutants (sulphur gases, nitrogencontaining impurities, chlorine compounds and vapour-phase metals) generated when coal, wastes or mixtures of them are gasified may be occasionally evaluated, if needed. Since the heterogeneity of gasification feedstocks results in a formation of a wide variety of products of different chemical nature it is envisaged that the sampling system will have to be enhanced and up-dated on a short term basis. Therefore, there is room for improvement in sampling methods in order to meet all the requirements of the different gasification applications.
REFERENCES 1.
Abatzoglou, N.; Barker,N.; Hasler, P.; Knoef H.; “The development of a draft protocol for the sampling and analysis of particulate and organic contaminants in the gas from small biomass gasifiers” Biomass 8c Bioenergy, 18 (2000). pp. 5-17, Jan. 2000. 2. Andersson C.; Coederbom, J. “FTIR analysis of flue gases-combined in-situ and 148
extractive gas sampling”, SVF-590, OSTI, NTIS, DE97722061, Oct 1996. 3. Brage, C.; Yu,Q., Chen, G., Sjtistrtim, K. “Use of Amino phase Adsorbent for Biomass Tar Sampling and Separation”, Department of Chemical Engineering and Technology, Chemical Technology, Royal Institute of Technology, Stockholm, Sweden, Fuel, Vol. 76, No. 2, pp. 137-142, Jan. 1997 4. BTG, Biomass Technology Group BV, The Netherlands; NRI, Natural Resources Institute, United Kingdom; ZrE, Zweckverband Regionale Entwicklung und Energie, Germany; “Development of a standard procedure for gas quality testing in biomass gasifier plant/power generation systems”, JOU2-CT93-0408, Final Report, June 1995. 5. CRE Group Ltd. “Identification and Processing of Biomass Gasification Tars”, Energy Technology Support Unit (ETSU), Department of Trade and Industry, DTI Contract, No. B/T1/00418/00/00, 1997. 6. Esplin, G.J.; Fung, D.P.C.; Hsu, C.C. “Development of Sampling and Analytical Procedures for Biomass Gasifiers”, Can. J. Chem. Eng. 63 (6), pp. 946-953, 1985. 7. Knoef, H.A-M.; Koele H.J.; “Survey of tar measurement protocols” Biomass & Bioenergy, 18 (2000), pp. 55-59, Jan. 2000. 8. Kurkela, E.; “Formation and Removal of Biomass-Derived Contaminants in Fluidized-Bed Gasification Processes”, Espoo, Technical Research Centre of Finland, 1996,47 p. + app. 87 p., VTT Publications 287, ISBN 951-38-4945-7 9. Mc Donald E.C.; Aiken, M.; Esplin, G.; “Development of Analytical Methodology for Biomass Gasification products”, ENFOR Project No. C- 172. Bioenergy Development Program. Alternative Energy Technology Branch. Canmet. Ottawa: Energy, Mines and Resources Canada, 1983. 10. Milne, T.A., Evans, R.J. Abatzoglou, N.; “Biomass Gasifier Tars: Their Nature, Formation and Conversion”, NREL, US DOE Contract No. DE-AC3683CH10093, NREUTP-570-25357, NOV.1998. 11. Moersch, 0.; Spliethoff, H.; Hein, K.R.G.; “Quasi Continuous Tar Quantification with a New On-line Analyzing Method”, 10” European Conference and Technology Exhibition, Biomass for Energy and Industry, Wiirzburg, Germany, 811 June 1998, Proceedings of the International Conference, pp. 1638-1641. 12. Simell, P.; Stahlberg, P.; Kurkela, E.; Albrecht, J.; Deutsch, S.; Sjostrom, K.; “Provisional protocol for the sampling and analysis of tar and particulates in the gas from large-scale biomass gasifiers. Version 1998”, Biomass & Bioenergy, 18 (2000), pp. 19-38, Jan. 2000. 13. Stahlberg, P.; Lappi, M., Kurkela S . ; Simell, P.; Oesch, P.; Nieminen, M: “Sampling of Contaminants from Product Gases of Biomass Gasifiers”, 1998. VTT, Espoo. 49 p. + app. 46 p., VTT Tiedotteita - Meddelanden - Research Notes: 1903, ISBN 95 1-38-5297-0; 95 1-38-5298-9 14. Techwest Enterprise Ltd., “A Workbook for Biomass Gasifier Sampling and Analysis”, ENFOR Project No. C-172, Bioenergy Development Program. Alternative Energy Technology Branch. Canmet. Ottawa: Energy, Mines and Resources Canada, 1983.
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Characterization of Products from Biomass Tar Conversion Ph. MorF, Ph. Haslerz, M. Hugener3,Th. NussbaumerZ ISwiss Federal Institute of Technology, ETH Zurich, CH-8092 Zurich, Switzerland Verenum Research, CH-8006 Zurich, Switzerland 'Swiss Federal Laboratories for Material Testing and Research, CH-8600 DuebendorJ; Switzerland
ABSTRACT: Secondary tar reactions such as homogeneous and heterogeneous tar conversion, partial oxidation and repolymerization are studied in the current project "Fundamentals of Tar Reactions during Biomass Fixed-bed Gasification".The present work describes the results obtained by experiments with homogeneous tar reactions whereas the focus is on the analytical method for tar characterizationand on the modeling of tar conversion. The experiments are conducted in a test stand for tar conversion consisting of three stages (primary pyrolysis, reactors for homogeneous and heterogeneous tar reactions). Homogeneous tar conversion experiments were performed by sweeping a pyrolysis gas through a tubular reactor which was operated at several temperature levels up to 1000 "C. The tar sampling is realized using impingers with an organic solvent. The samples were analyzed by means of Gel-Permeation Chromatography (GPC) with a multiwavelengthUV-detector. Qualitative and quantitative evaluation of the resulting spectro-chromatograms of the tar samples are performed by using chemometric methods. The results show that homogeneous tar conversion tremendously alters the tar composition. The tar depletion by homogeneous reactions can be modeled by a simple one-lump model, assuming plug flow and a single, first-order reaction.
INTRODUCTION
Tars in the producer gas from biomass gasifiers can cause severe operating problems if the gas is used as fuel for internal combustion (IC) engines. Secondary measures such as gas cleaning systems are costly and often do not fulfill the requirements for IC engine &el gas quality reliably. Gasification process design and process parameter optimization regarding the producer gas quality are an alternative strategy for approaching the tar problem. Fundamentals of tar generation and conversion are essential in order to realize improvements that belong to such primary measures. Therefore, the goal of the present
150
work is to provide fundamentals about tar conversion in fixed-bed biomass gasifiers. SECONDARY TAR REACTIONS
Tars are primary products from pyrolysis of biomass or other carbon-containing materials. Their formation cannot be avoided. However, these primary tars are not stable. After having evolved from the solid matter, they undergo secondary tar reactions, altering both in mass and composition. As shown in$g. I, secondary tar reactions
Char
-
Char
Gases -+Gases
Gasas
-9
lntraparticle
w
Y
- Extraparticle
+
Fig. I Schematic of primary pyrolysis and secondary tar reactions pathways within (1) and outside (2) of the wood particle. After Boroson [13 (modified). appear within and outside of the biomass particle. This intra- and extraparticle conversion of tars can be divided in the following classes: (1)
Homogeneous tar reactions in the gas phase, such as tar cracking, partial oxidation, repolymerization
(2)
Heterogeneous tar reactions on the surfaces of (partially) reacted biomass, char particles etc. The same types of reactions as in the homogeneous case (cracking, oxidation, repolymerization)may occur but the surfaces may have a catalytic effect on the reactions.
If the major mechanisms of secondary tar reactions are known, tar mass and composition may be controlled by choosing appropriate reactor conditions. Whereas there exists a substantial amount of studies on primary pyrolysis of biomass, investigations on secondary tar reactions are scarcely found [ 1,2]. The literature on secondary tar reaction mainly consists of studies on homogeneous tar conversion ("homogeneous tar cracking", [3,4,5]). Studies on heterogeneoustar conversion are even more rare [ 11. Therefore, the goal of the present ongoing project is to provide the fundamentals of tar conversion during secondary tar reactions in the fixed-bed of biomass gasifiers. In this paper, the experiments and modeling of homogeneous tar conversion without the presence of oxidants are given.
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EXPERIMENTAL TAR CONVERSION REACTOR SYSTEM
A laboratory reactor system was constructed for the investigation of tar conversion under real fixed-bed conditions. The basic idea behind the reactor design was the spatial separation of primary pyrolysis and secondary tar reactions. A scheme of the system is shown in jig. 2. After the primary pyrolysis stage, the tar containing pyrolysis gas is swept to the conversion reactors where homogeneous and/or heterogeneous tar reactions can be investigated.
Fig. 2 Principle scheme of tar conversion reactor system
The most important features of the system are its continuous operation that ensures a constant and defined pyrolysis gas composition, and the use of large wood chips (10-20 mm) as fuel. The realization of these features led to the design of the wood pyrolysis unit with a fixed-bed of fuel which is continuously conveyed through a heated tubular reactor. Moreover, the choice of this reactor design provides a pyrolysis gas which is generated under reactor conditions similar to real fixed-bed systems. This is important since secondary tar reactions occur as soon as the tar vapors have been evolved from the solid matter and, therefore, "primary" tars generated in a fixed-bed differ considerably from "primary" tars from single particle pyrolysis experiments (different heating rates of wood etc, cf. 161). In the following, the term "primary" tars refers to tars at the outlet of the wood pyrolysis unit (WPU), being aware that these tars have already undergone some degree of secondary reactions. Operatingprocedure for homogeneous tar conversion experiments
A pyrolysis gas with a high content of "primary" tars (approx. 300 g/Nm3) is produced in the WPU by continuously pyrolyzing wood chips (feed rate: 1.6 kgh) at a mean temperature of 380 "C. The evolving gases are swept to the adjacent reactor for homogeneous tar conversion (HOTCR) by means of a N,-carrier gas flow (for these 152
experiments: 0.4 Nus). At the end of the pyrolysis reactor, the produced char is separated from the gas flow and stored in a container. The HOTCR is operated at several temperature levels up to 1000 "C. The carrier gas flow rate and the reactor tube diameter are chosen to ensure a mean gas residence time below 0.5 s. Such low residence times are needed to determine the kinetics of homogeneous tar conversion without oxidants, cf. [l]. Before and after the HOTCR, tar samples are taken and the concentrations of non-condensable gases (CO, C02,CH,, H2) are measured on-line. The gas volume flows before and after the HOTCR are measured by a special-designed Pitot-tube system. TAR SAMPLING AND TAR ANALYSIS
The tars are sampled discontinuously by using impingers with an organic solvent. At the beginning of the sampling train, a liquid-quench system is used. After several experiments with different organic solvents, 1-Methoxy-2-Propanol was found to be an excellent solvent for sampling both primary and converted tars. Approx. 170 NI process gas is sampled from every interesting operation state of the reactor system during an experiment. The samples are analyzed by Gel-Permeation Chromatography (GPC) and subsequent UV-detection. The instrument used is a Waters 150-C GPC-System with a column from Polymer Laboratories (50 A, range 0-2000 MW, length 60 cm). Tetrahydrofuran (THF, Romil high purity, GB) is used as solvent. The temperature of the GPC-system is 30 "C.The parameters of the UV-detector are the following: UV-range: 210-300 nm, sampling rate: s-I, resolution: 1.2 nm. The resulting spectro-chromatograms (SCG) are 3D-representations of the tar matrices with the UV-absorbances as function of the retention time in the gel column and the wavelength of absorption, respectively (jig. 3). Sections of the SCG parallel to the retention time axis at 215 nm UV-absorption ("tar profiles" in the following) enable quick qualitative tar characterization. For the quantitative evaluation of the SCG, chemometric methods such as factor analysis and the classical least squares method are applied. This requires the set-up of a spectral library which contains the SCG of the quantitative important tar compounds.
Fig. 3 Spectro-Chromatogram of a "primary" tar sample from the wood pyrolysis unit
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RESULTS FROM EXPERIMENTS
Evolution of nun-condensablegases during homogeneous tar conversion Fig. 4 shows the determination of the molar fluxes of the non-condensable gas species that are measured on-line during the experiments. The values are determined by multiplying the normalized gas volume flow from the Pitot measurements with the measured gas concentrations.The first three data points (around 420 "C) are measurements of the "primary" pyrolysis gases at the WPU outlet, the other values are determined after the HOTCR.
Temperature ["C] 9
-
x
Temperature ["C]
lo4 CH4
8 .
0
,
5 7 . 0
1 0
3 . 0
0
0
0
2 400
600 800 Temperature ["C]
1000 Temperature PC]
Fig. 4 Molar gas fluxes of non-condensable gases during homogeneous tar conversion at residence times in the HOTCR below 0.5 s. As expected, the gas molar fluxes of all the measured species increase with increasing reactor temperature. The value around 550 "C seems to be an outlier among all species.
This may be due to fluctuations of the gas volume determination which is used to calculate all molar fluxes. Carbon monoxide is the major product of the non-condensable gas species at the highest reactor temperature. In the study of Boroson [ 11, CO was dominant at all reactor temperatures. Therefore, he concluded that CO could be used as an indicator of secondary tar cracking. The present results however, indicate that the hydrogen is even a better indicator for homogeneous tar conversion. Its evolution curve shows the steepest rise, from negligible amounts in the "primary" pyrolysis gas to values almost 30 times higher around 930 "C. Interpreting these results, one should keep in mind that these non-con-
154
densable gas species are not only products from tar cracking (fragments from larger tar molecules) but also may derive from gas-phase reactions of non-condensable products from "primary" pyrolysis or tar cracking, such as the methanation reaction.
Alteration of tar composition during homogeneous tar conversion
GPC analysis is conducted with the tar samples from the experimental series on homogeneous tar conversion with short gas residence times. Up to now, qualitative information about tar composition can be extracted from the comparison of the tar profiles. As outlined before, the quantitative evaluation of the spectro-chromatograms is performed by means of chemometric methods. The resulting tar profiles at 215 nm fiom the present investigations are shown infig. 5.
.c
E l $
-0
a o
$
3
2 1
0 2 1
0
10
15
18
20
25
Retention Time [min] Fig. 5 Tar profiles at 215 nm from the homogeneous tar conversion series (gas residence times below 0.5 s). Indicated areas: a: phenols, b: organic acids, c: hrans, d: PAH
The differently shaded patches indicate the ranges where mainly (but not exclusively) tar substances of similar compound classes appear. As can be seen infig. 5, the compound classes ''phenols'' (a), ''organic acids" (b) and "hrans" (c) all continuously decrease with increasing temperature during the homogeneous thermal treatment of the pyrolysis gas. Tar compounds belonging to these classes have been defined as "primary" and "secondary" tars in the previous literature [7], refering to the extent of secondary reactions undergone. On the other hand, the class "polyaromatic hydrocarbons" (PAH, d) exhibits no clear trend, there seems to occur a redistribution within this class.
155
As mentioned before, the gas residence times in this experimental series were kept low to ensure that a kinetically controlled reaction regime can be investigated. For comparison, fig. 6 shows the tar sample of a similar homogeneous tar conversion experiment, but with high gas residence times (up to 13 s) in the HOTCR [6]. 1
05 0 1
E
0.5
C
m o
1
05 n
70
18
15
20
25
Retention Time [min] Fig. 6 Tar profiles at 2 15 nm from the homogeneous tar conversion series (high gas residence times). Indicated areas: a: phenols, b: organic acids, c: furans, d: PAH
The additional analysis performed in this experimental series indicates that the pyrolysis gas species were in thermodynamic equilibrium (cf. [6]). By comparison of the two tar profile series, the effect of the residence time as limiting parameter with respect to the conversion at a specific temperature level is clearly observeable. In the case of long residence times, a tar sample can be produced which consists almost exclusively of PAH, with some minor quantities of phenolic species. MODELING OF TAR CONVERSION
Mathematical modeling is becoming more important for designing appropriate reactors for thermochemical biomass conversion processes such as pyrolysis, gasification and combustion. As denoted in the literature [2], there are few reliable kinetic models which describe secondary tar reactions. However, for predicting correct tar yields and compositions from pyrolysis and gasification processes such kinetic models are essential [S]. Therefore, modeling of the tar conversion according to the experimental basis is performed.
156
ONE-LUMP MODEL FOR HOMOGENEOUS TAR CONVERSION
As a first approach, the homogeneous tar conversion is described by a simple one-lump model and assuming a single reaction pathway: Tar -+Gases
(1)
Altough this seems to be a very rough simplification of the real situation, similar studies reported in the literature [ 1,5,8] have shown that tar mass depletion by homogeneous reactions can be described in this manner quite accurately. As data base for the tar mass, the gravimetric tar values were determined from the tar samples of the above described experiments on homogeneous tar conversion. Mathematical modeling
The goal of the modeling is to find the best-fit values of the parameters of the Arrhenius constant that, in this case, describes homogeneous tar conversion:
--E k = k,e
RT
The parameters to be found are the frequency factor k, and the activation energy E. For the describtion of the reaction pathway of the one-lump model, a first-order reaction for the tar depletion is assumed: -‘Tar
=
kc,,,
The flow pattern in the tubular reactor (HOTCR) is described as ideal plug flow. The performance equation relates the conversion of the tar XTarand the gas residence time T :
x,, = ‘Tar,
1
(4)
xrdxTar 0
The high concentrations of primary tars in the feed gas result in a significant expansion in the gas volume during the reactions in the HOTCR. Thus, the changing gas volumes due to the gas-phase reactions are considered by the relationship of tar concentration and the tar conversion
-‘Tar - ‘Tar,
-‘Tar +
‘Tar’Tar
where is the fractional change in the gas volume between no conversion and complete conversion of the tar [9]. This value is estimated from the evaluation of a previous experimental series in homogeneous tar conversion with long residence times (cf. remarks in sections above and in [ 6 ] ) .
157
Combining eq. 3,4 and 5 and integrating the resulting expression finally leads to the performance equation for the plug flow reactor with a first-order reaction [9]:
kT = ( 1 - &TUr)ln-
1
- 'Tur
- &TnrXTUr.
The Arrhenius constants are calculated with the experimental data on homogeneous tar conversion using eq. 6. The residence times needed for this procedure are estimated by computing the mean values of the highest and the lowest possible residence time (immediate reaction at the reactor inlet and reaction at reactor outlet only, respectively) at steady-state operation at a temperature level. Finally, the kinetic parameters for homogeneous tar conversion are determined by least squares fitting of the determined values of k. Fig. 7 shows the resulting Arrhenius plot.
1-
.2
0-
Y
c -1
-
-2
-
-3 -
O(85
d.9
O(95
i
l(05
lll
1 / l [l/K]
lll5
lj2
l(25
.3 xl
3
Fig. 7 Arrhenius plot of 1-lump model for homogeneous tar conversion (solid line) and experimental data (circles)
The values found for the parameters of the Arrhenius constant are A. = 223'668 sI and E = 99 kJ/mol. The simple one-lump, single reaction model describes the experimental data quite well @g. 7). Since the error in the experimental determination of the gravimetric tar value is unknown, the specification of an independent "goodness-of-fit" value is not possible [lo]. The absolute error of k is estimated as 1.4 s-l. The findings of this kinetic investigation are compared to previous similar studies. (Table I). The parameters found in in this work lie between the values of the other two studies fig. 9). The Arrhenius constants determined by Boroson [l] and the present study show very good agreement. However, it should be noted that in the literature, different approaches for the reaction rates with respect to the concentration or mass depen-
158
dent term are made. Thus, a common approach in the literature cited (i.e. [ 11) is that the tar conversion rate is modeled as first order in the difference between the ultimate yield of tar and the total amount of tar unconverted at that time. Using this approach, the ultimate yield is another fitting parameter. Table 1: Comparison of kinetic parameters for homogeneous tar conversion fiom different studies Study
k, [s-'I
Liden [ 5 ]
106 49
107.6
Boroson [I]
10498
93.4
Present study
105 35
99.4
E [kJ/mol]
Conversion Temperature [K] Fig.8 Temperature dependency of Arrhenius constants fiom three studies
CONCLUSIONS A lab reactor system for the investigation of secondary tar reactions was constructed. Several experimental series on homogeneous tar conversion were performed. Gel-permeation chromatography (GPC) is successfully used as a tar analysis method. The tar profiles resulting from this technique allow quick qualitative comparison of tar samples
159
produced under different reaction regimes. The tar profiles and the results from the gas analysis show that the present experiments on homogeneous tar conversion are performed in the kinetically controlled reaction regime. Thus, the tar profiles represent tars which are much less converted than the tar samples which are produced with high residence times in the conversion reactor. The latter tar profile series indicates thermodynamic equilibrium. The homogeneous tar conversion is described by an empirical model. The results show that the simple one-lump, single first-order reaction model describes the depletion of the gravimetric tar satisfactory well. The comparison with literature studies shows a good agreement of the determined kinetic parameters.
Future work In the ongoing project on secondary biomass reactions in fixed-bed gasifiers the following issues will be pursued in the near future:
(1)
Experimental investigation of the heterogeneous tar conversion, i.e. the influence of the char fixed-bed on secondary tar reactions
(2)
Development of chemometric methods for full quantification of the biomass tar compounds
(3)
Development of a more detailled model of tar conversion, including several lumps for the description of tar properties
ACKNOWLEDGEMENTS
The authors would like to thank the staff at the Swiss Federal Laboratories for Material Testing and Research for their assistance and interest in the project and for carrying out the GPC analysis. This research project is financed by the Swiss Federal Office of Energy which is gratefully acknowledged by the authors. NOTATION CTar CTar 0
E ;Tar
ko z rT, R T XTar
Tar concentration in pyrolysis gas [s/Nm3] Initial tar concentration in pyrolysis gas [g/Nm3] Activation energy FJ/mol] Expansion factor, fractional volume change on complete conversion of tar [-I Arrhenius or reaction rate constant [s'] Frequency factor [sl] (Mean) gas residence time in conversion reactor [s] rate of reaction of the tar, here [g/Nm3s] Ideal gas law constant, R = 8.314 [J/mol K] Temperature [K or "C] Tar conversion or fraction of tar converted [-I
160
REFERENCES
1. Boroson, M.L. (1987): Secondary Reactions of Tars from Pyrolysis of Sweet Gum Hardwood. Ph. D. Thesis, Massachusetts Institute of Technology. 2. Morf, Ph., Nussbaumer, Th. (1998): Grundlagen zur Teerbildung bei der Holzvergasung. BFE Zwischenbericht. Swiss Federal Office of Energy, Bern. 3. Brandt, P., Hendriksen, U. (1996):Decomposition of tar in pyrolysis gas by partial oxidation and thermal cracking. In: Chartier et al. (Eds.). Biomassfor Energy and the Environment. Elsevier, Oxford. 4. Diebold, J. P. (1985): The Cracking Kinetics of Depolymerized Biomass Vapors in a Continuous, Tubular Reactor. M. S . Thesis, Dept. of Chemical and Petroleum-Refining Engineering, Colorado School of Mines, Golden, Colorado. 5. Liden, A. G. (1985):A Kinetic and Heat Transfer Modelling Stua), of WoodPyrolysis in a Fluidized Bed. MASc Thesis, Dept. of Chemical Engineering, University of Waterloo. 6. Morf, Ph., Hasler, Ph., Nussbaumer, Th. (2000): Mechanisms of Tar Conversion during Fixed-Bed Gasification. Proceedings of the 1st WorldConferenceand Exhibition on Biomassfor Energy and Industry. Sevilla, Spain. 7. Evans, R.J., Milne, T.A. (1987):Molecular Characterization of the Pyrolysis of Biomass. 1. Fundamentals. Energy & Fuels, 1987 (A), l(2). 8. Gr~nli,M. (1996):A Theoretical and Experimental Study of the Thermal Degradation of Biomass. PhD Thesis, Norwegian University of Science and Technology, Trondheim. 9. Levenspiel, 0.(1999):Chemical Reaction Engineering. 3rd Edition. John Wiley & Sons, New York. 10.Press, W. H.,Teukolsky, S. A., Vetterling, W. T., Flannery, B. P. (1986):Numerical Recipes in C. The Art of Scientific Computing. 2nd Edition. Cambridge University Press, Cambridge.
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Guideline for sampling and analysis of tars and particles in biomass producer gases J.P.A. Neeft', H.A.M. Knoef2,G.J. Buffinga2, U. Zielke3,K. Sjostrom4, C. Brage4, P. Hasler', P.A. Smell6, M. Suomalainen6,M.A. Dorrington7, C.Grei18 ': ECN, P.O. Box 1, 1755 ZG Petten, The Netherlands
BTG, P.O. Box 217,7500 AE Enschede, The Netherlands 3: DTI, Kongsvang All6 29, DK-8000 Aarhus, Denmark 4: KTH, Teknikringen 42,10044 Stockholm, Sweden 5 : Verenum, Langmauerstrasse 109,8006 Zurich, Switzerland 6: VTT, P.O. Box 1601,02044 VTT Espoo, Finland 7: CRE Group, Stoke Orchard, GL52 4RZ Cheltenham, UK 8: Lurgi, Lurgiallee 5, D-60295 Frankfurt am Main, Germany 2:
ABSTRACT This paper reports the further development of a Guideline (formerly Protocol) for sampling and analysis of tars from biomass producer gases. This Guideline is being developed in a project in the European Fifth Framework Programme with additional partners from Switzerland and North America. This paper gives the outline and principle of the Guideline. The Guideline is based on isokinetic sampling of particles and tar from the main producer gas duct, particle filtration at high temperature, gas cooling in a liquid quench, tar absorption in a solvent at low temperatures, an optional backup adsorber, and flow measurement and control. The Guideline gives a definition for Gravimetric tar which is the tar number to be determined by the Guideline. Besides, the Guideline gives procedures for compound analysis by GC-MS or GC-FID. Moreover, in this paper the major choices that were made to come to the first version of the Guideline are explained. Finally, at the end of the paper it is described how and on what time scale the development of the Guideline will be completed. This paper does not contain the full text of the Guideline. This full text will be available on the Internet at www.tarweb.net. 1. INTRODUCTION
The main contaminants in the product gases of biomass gasification are dust and soot particles, organic contaminants (often being referred to as tars [ l]), alkali metals, acid 162
gases and alkaline gases. Measuring techniques for these contaminants allow to determine the functioning of the gas cleaning and to assess the quality of the cleaned gas to be used in the engine or turbine. For most contaminants, well-developed measurement techniques exist which are similar to techniques used for related technologies, such as coal combustion and coal gasification. For organic contaminants, however, no well-developed measurement technique exists in related technology fields. As some of the organic contaminants are seen as the major problem-causing contaminants in biomass gasification, manufacturers and other workers in this field have used a number of different sampling and analysis methods to determine the level of organic contaminants. In these methods, also the definition of organic contaminants or tars has been rather diverse. Recently, a project in the European Fifth Framework Programme has started which is aimed at the development of a standard method (Protocol) for the sampling and analysis of organic contaminants (tars) in biomass producer gases. The partners in this project are named in Figure 1. This project has a history dating back to a joint meeting in Brussels (1998) by members of the Gasification Task of the IEA Bioenergy Agreement, US DOE and DGXVII of the European Commission. This meeting led to the development of two draft Protocols for tar measurement, which were presented at the loth European biomass conference in Wurzburg, Germany and were recently published [2-41.The current project is a continuation of this early work.
Group of Contractors
Group of Reviewers
The Netherlands
Enerkern Canada
Lurgi
Sweden
Germany
TUV
Austria
UCL
Belgium
NTUA
Greece
CIRAD France TPS
Sweden
FWEOY Finland
Figure I Project partners (contractors and reviewers) of the EU project with acronym ‘Tar Protocol’
The history, objectives and structure of this EU project were presented in a previous paper [ 5 ] . The current paper will outline the contents and principle of the first version of the Protocol, which is now called Guideline for reasons described later. In the next Chapter 2 the principle of the Guideline will be described. Then, in Chapter 3 attention is paid to choices that were made so far to come to version I of the Guideline. Examples of such choices are the name Guideline instead of Protocol, and
163
the choice for not defining tar whereas Gravimetric tar is defined in the Guideline. In the final Chapter 4 it is shortly described how and on what time scale the development of the Guideline will be completed and which R&D activities are to be performed before completion and field testing of the Guideline.
2. PRINCIPLE OF THE GUIDELINE
INTRODUCTION This chapter contains a concise description of the working principle and the structure of the Guideline for sampling and analysis of tars and particles from biomass producer gases. This description includes the results of many choices, which were made during the development of the Guideline. The rationale behind these choices is not described in this chapter, as this would draw the attention from the working principle and structure of the Guideline. Instead, by using numbers in superscript, reference is made to the next Chapter 3 in which this rationale is explained. One of the most fundamental questions that rises when discussing a Guideline or Protocol or Standard for tar measurement is “what is it’s use?”, in other words “is it really needed?”. This is the first question answered in the next chapter’. The second question will address the word “Guideline” which is now used instead of “Protocol”*.
WORKING PRINCIPLE The measurement principle of the Guideline is based on the discontinuous sampling of a gas stream containing particles and condensable organic compounds. The principle is set-up in such a way that also particles can be measured quantitatively3. Condensable organic compounds will be further referred to as tars although this word tars is not defined in detail4. The measuring principle has been kept as simple as possible because the measuring conditions can vary from very comfortable laboratory rooms to an executive plant gasifier where there is no customised room for measurements or measurement apparatus. Also the weather conditions can be challenging, for instance in northern Europe measurements might have to be performed at temperatures under 0°C. Sampling is performed under isokinetic conditions5. The tar and particle sampling system consists of a heated probe, a heated particle filter, a condenser and a series of impinger bottles containing a tar absorbing solvent. A backup adsorber for volatile organic compounds can optionally be placed after the impinger bottles. The solvent containing bottles are placed in a cold bath so that gradual cooling of the sampled gas from about O°C to the final temperature - 20°C takes place in them. The sampling train is shown schematically in Figure 2. The gas sample is led to flow a sufficiently long time through the sampling line and the filter either with the aid of process pressure or a pump included in the sampling line to the gas impingers. A continuous gas flow and heating is necessary in the sampling systems. The flow rate of the sample gas is monitored and controlled during sampling. The sampling lines including the filter should be heated in order to prevent tar compounds from condensing. However, to avoid thermal decomposition of compounds, the temperature should not be too high. In updraft gasification the temperature should be below 4OO0C, while in downdraft and fluidised-bed gasification it should be below 700°C. According to experience in practice, 200 - 300°C is suitable gas temperature in the sampling lines. 164
Q + Vent
80 Particle filter (heated)
1
Volume flow meter
Backup adsorber
(T = 0 "C)
Gas washing bottles
'Salt and ice
bath (T = - 20°C)
Figure2 Sampling train of the Guideline for sampling and analysis of tars and particles in biomass producer gas
Particles are collected in an externally heated quartz filter6 which is heated to 250°C or 100°C depending on the gasifier type from which producer gases are sampled7.Then, the producer gas is cooled, either on a heat exchanger surface or in an internal circulating liquid quench system in a liquid'. Gas quenching by an externally cooled liquid is done immediately after the particle filter and facilitates sampling from producer gas with high tar loadings or with reactive tar species. The condensate is collected in a glass bottle. After the condensor, the gas passes an impinger filled with solvent before further temperature letdown. In the following series of impingers', tars are absorbed in the solvent" at a temperature of -20°C. A backup residual tar adsorber" can optionally be used as a safety filter after the impinger train. The volume, temperature, pressure, and gas flow rate through the equipment are measured after the impinger bottles. Immediately after sampling the contents of the impinger bottles, including the glass beads, are decanted into a storage bottle. All surfaces (including metal surfaces) becoming in contact with the gas at a lower than the process temperature are washed with the solvent, and the washes are combined with the actual sample. This is quite easy to arrange in atmospheric processes, but very difficult in pressurised systems. The storage bottle is stored tightly closed at cool, c 5"C, temperature for later analysis. The sampling of tars and particles can also be performed separately. For large scale gasifiers where high temperature tar is formed, the generally used sampling strategy is based on separate sampling of particles and gaseous effluents. In case of measuring only tars, isokinetic sampling' is only necessary when the temperature of the gas under study at the sampling site is so low that the organic compounds are condensed and form aerosols andlor drops.
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BASIC CONCEPT OF THE SAMPLING TRAIN The modular sampling train consists of 4 main modules and their submodules. The main modules are gas preconditioning, particle collection, tar collection and volume sampling. The modules are shown in Figure 3 and the purposes of the modules and needed equipment are aggregated in Table I.
I s
Module 4 Submodules
I
T-l-T
Module 3
I
I Submodules I
Figure 3 Basic concept of the sampling train
In the preconditioning module (Module 1) the gas is led to cool or is heated from the process temperature to a constant temperature, 250 "C, in a heated probe. In case of pressurised gasification the pressure is diminished to atmospheric pressure. The equipment needed in the module is a sampling probe including a nozzle and necessary valves, at least one shut-off valve and one control valve, and in pressurised gasification also a pressure relief valve. Also the valves are heated to constant temperature. In the particle collection module (Module 2) a heated thimble filter6 at the probe temperature collects solids from the gas. The tar collection module (Module 3) consists of three submodules. In the submodule 3.1, the gas is cooled and moisture and heavy tars are collected in a condenser' at the temperature O O C . In the simplest version, the condenser is an empty impinger bottle in an external ice water bath. More effective versions use either an additional externally cooled heat exchanger before the condenser or an internal liquid quench system. In the submodule 3.2 tars and volatile organic compounds are absorbed into the solvent at -20°C in a series of impinger bottles9. The optional submodule 3.3 is a backup adsorber' ', which collects residual volatile organic compounds that may have penetrated the impinger train. The backup adsorber is not necessary when enough impinger bottles, appropriate solvents" and collection temperatures are used. In a later
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version of the Guideline, the low temperature impingers may be replaced by an adsorber tower filled with a solid sorbent. The volume-sampling module (Module 4) consists of three submodules. The purpose of submodule 4.1 is the gas suction or pump (not needed in pressurised gasification). In the submodule 4.2 a gas meter measures the volume of gas flow. Additionally a flow meter, a pressure indicator and a temperature indicator are needed.
Table I
General description of modules and submodules with purpose and equipment used
Module
Purpose
Equipment
Module 1 (Gas preconditioning) Gas cooling, pressure letdown Nozzle, valves, sampling lines Module 2 (Particle collection) Collection of solids
Heated filter (high temperature)
Module 3 (tar collection) Submodule 3. I Moisture and tar condensation Condenser at 0°C Submodule 3.2 Submodule 3.3
Tarhydrocarbon collection Backup sampler
Module 4 (Volume sampling) Submodule 4.1 Gas suction Submodule 4.2 Submodule 4.3
Impingers with solvent at -20°C Adsorber tower at Tamb or lower Pump
Gas volume integration Venvexhaust gas handling
Gas meter, P and T indicators Outdoor ventilation
DESIGN OPERATING CONDITIONS The sampling train is designed for sample flow rates ranging from 0,l - 0,6 m3Jh.The volume of the gas drawn to the tar sample is dependent on the tar content of the gas. A suitable sample volume in (pressurised) fluidised bed gasification (at a tar concentration c 20 g/m3,) is about 0,l m3, and in updraft gasification (tar concentration 100 - 200 g/m3,) 0,05 - 0,l m3,. The total content of tar in the solvent should be about 10 g/dm3 for gravimetric determination and about 2 g/dm3 for gas chromatographic analysis. However, the detection limit is significantly lower for single components (0.25 ppm). The temperature in the sampling line of the product gas of the updraft gasifier is kept at 100-250°C (to be examined) and in downdraft and fluidised-bed gasification at 250°C.
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SITE SPECIFIC SAMPLING TRAIN SET-UPS Applying the sampling principle depends additionally on the gasifier type. Most of the gasifiers fall in four categories based on the contact type of the feedstock and on the gasification agent. The four main types are: 1) fixed-bed updraft, 2) fixed-bed downdraft, 3) fluidised-bed and 4) entrained-flow. The yields of tars and particles depend greatly on the gasifier type. The tar concentration in countercurrent (updraft) gasifier is remarkable high (typically 100 - 200 g/m3,) comparing to the other gasification applications (tar content usually below 20 g/m3,). Also the composition of the tars varies depending on the gasifier type. The particle concentration in the product gas is the largest in the fluidised bed gasifiers. The amount of particles and amount and composition of tars determine the specific sample train arrangements. Therefore, this arrangement depends on: 0 Gasifier type (as described above); 0 Gasifier operation condition (specific pre-conditioning steps are foreseen for sampling from pressurised gasifiers or at low temperatures); 0 Gasifier scale (sampling at large scale gasifiers at high temperature might be performed non-isokinetically); 0 Sample location: before or after gas cleaning (when sampling after the gas cleaning, use of a backup tar adsorber might not be necessary). The Guideline will give examples of such site-specific sample trains.
TYPE OF SOLVENT USED The condensate should preferably be collected in a solvent, not in a cold trap in which there is no “diluting surroundings”12.The modular sampling train set-up displayed in this chapter can be used for both water miscible and non-water miscible (non-polar) solvents. For the water miscible solvents, I -methoxy-2-propanol has been selected as’the most promising candidate. In a later stage, only minor modifications will arise in the Guideline when a non-polar solvent will be used. The boiling point (evaporation rate) and the water solubility are the main two parameters, which influence the applicability of the liquid quench system in the sampling train set-up.
POST-SAMPLING A N D ANALYSIS The Guideline will provide procedures for post-sampling and analysis of the tar samples. Basically, tars are analysed in two ways: gra~imetrically’~ and by gas chromatographyi4. The gravimetrical analysis will result in one concentration of gravimetric tars (see note in the next chapter for a definition) whereas the chromatographic analysis will result in concentrations.of individual compounds. At this point the procedures for post-sampling and analysis are not given, as they will to some extent depend on the solvent being chosen. The time frame in which this choice will be made and in which the full Guideline will be available, is given in Chapter 4.
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3. RATIONALE BEHIND CHOICES MADE DURING DEVELOPMENT OF THE GUIDELINE During the development of the Guideline many choices were made. Some of these choices are practical choices to move along with the development of the Guideline. Other choices, however, are fundamental ones that were only made after long discussions in the Group of Contractors. These choices and their background discussions are to some extent explained in this chapter. The choices are numbered: the numbers refer to the number in superscript in the previous chapter “principle of the Guideline”. 1. Why do we need a Protocol or Guideline or Standard for sampling and
analysis of tars? In the past, several tar measurement methods have been developed in the field of biomass gasification. Many of these methods gave and give reliable data on tar concentration, tar being defined as part of the measurement method. Therefore, for tar measurements on single biomass gasification installations the Guideline is not needed and current practice is fine. Current practice becomes less obvious when tar data are to be compared, for instance when comparing the performance of biomass gasifiers or apparatuses for gas cleaning, or when setting or working with tolerances (maximum allowable concentrations) for tar in gas cleaning devices or prime movers. The value of the comparison will depend on definition of tar and on the ability to measure tar according to the definition. A recent report on parallel measurements with different tar measurement methods at the same location [6] shows that comparison of tar data can be a difficult task. So the expected value of the Guideline is in the field of comparison. Two types of comparison can be distinguished: 1. Comparison of tar concentrations from different locations. To do so, the Guideline has to be used at the two sites. 2. Comparison of existing methods for tar measurement. A requisite for reliable comparison of tar concentrations with different measurement methods is that these methods give good results in the ranges of conditions applied (temperatures, pressures, and concentrations and compositions of tars). The Guideline, which will cover large ranges of conditions, can be used to check the applicability of these methods. The comparison mentioned under 2. will result in knowledge on which method can reliably be used under which conditions. As other methods might be simpler (in terms of man-hours and equipment needed to perform the measurement) compared to the Guideline or produce online results on-site, we have named them “shortcut method^"'^. Examples of these methods are a) the solid phase adsorption (SPA) method developed by KTH [7], b) a number of solvent-free tar collection systems used by BTG, BEF, IGT [8-101 and c) the FID online tar analysing method under development at the University of Stuttgart [ I 1,121. Currently it is unknown under which conditions these methods give reliable results, for example it is unknown whether the SPA method can be used for updraft gasifier tars and at which conditions the solvent-free methods fail to collect all tars, for instance as a result of aerosol formation. In conclusion it can be said that the Guideline will not stop other methods from being used, which is also not the intention of the Guideline as it will be a method that is elaborate in comparison with many existing methods. The Guideline
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2.
3.
4.
5.
will provide a common method -with which other methods will be compared- that applies for the whole range of measurement conditions. Protocol versus Guideline. It was decided to not longer use the word ‘Protocol’ as it is rather formal and has a broad meaning. Instead, the word ‘Guideline’ will be used. This word better represents the current status of the method for sampling and analysis than the word ‘Protocol’. It should be kept in mind that the Guideline has a temporary status; the aim is to have a Standard. To reach this goal a Standardisation trajectory at CEN will be started during the execution of the EU project. The full title of the Guideline is ‘Guideline for sampling and analysis of tars and particles in biomass producer gases’. Measurement of tar and particles. Among the members of the Group of Contractors in the EU project ‘Tar Protocol’, it is recognised that the interaction between carbon-rich particles and tar can strongly influence measured tar concentrations. For example, measuring tars in a producer gas with relatively high carbon-rich particle concentrations and relatively low tar concentrations can be a difficult task because of these interactions. In the Guideline, the post-sampling procedures of the particle filter will receive thorough attention in order to produce reproducible results when measuring the concentration of Gravimetric tar (see the next point 4).Likewise, the concentration of particles is also influenced by these post-sampling procedures (e.g. during parallel measurements it was observed that the low particle concentrations in updraft gasifiers are difficult to measure as a result of tar condensation on the particle filter). Because of the attention paid to sampling and post-sampling procedures for particles, the Guideline will also give a number for particle concentrations and, therefore, is a Guideline for tar and particles. Definition of tars. The word tar is not defined in the Guideline because a definition (on top of the many definitions that have already been given, see [13]) will likely cause (further) confusion. The authors of this paper feel that the word tar is an ambiguous term that has no defined meaning and is used in several contexts with at least slightly different meanings. Even the definitions “compounds that condense” or “compounds that contaminate linings or compressors or moving parts in the prime mover” are not unambiguous as condensation or contamination varies with temperature of, and pressure in, the linings/compressors/moving parts as well as with concentration of tar itself. Besides, the Guideline is not developed for determination of a tar concentration. With the Guideline, two concentrations can be determined: - The concentration of Gravimetric tar; - The concentration of individual organic compounds; Gravimetric tar is defined as the evaporation residue at given standard conditions (temperature, pressure, duration). These conditions will be set later when the solvent has been chosen. In the Guideline, a compound list will be added with individual organic compounds including chemical abstract service (CAS) registry numbers. This list is compiled from experimental data on the presence of compounds found in biomass producer gases from updraft, downdraft and fluidised bed gasification. Isokinetic conditions. Because particles are sampled and analysed as part of the Guideline, the sampling is performed isokinetically. Instructions of isokinetic sampling of flue gases are given in the standards I S 0 9096 or VDI 2066. Isokinetic sampling with realistic probe dimensions requires, even at ambient pressures, high gas velocities up to 20 m / s with corresponding gas flow rates between
170
0.1 and 0.6 m3dh . Especially high flow rates are needed in pressurised systems to maintain isokinetic conditions and this is unpractical for impinger sampling using liquids. Therefore, a non-isokinetic sampling arrangement is practical for condensable organic compounds from pressurised systems at temperatures above their condensation point. Non-isokinetic tar sampling is also practical in large-scale atmospheric gasification applications where the pipe diameter is large causing the isokinetic sampling according to the standard to be very difficult and time consuming. The gas temperature above 350°C at the sampling point assures good accuracy in measuring tars also non-isokinetically. In non-isokinetic sampling the alignment of the probe in relation to the gas flow as well as the shape of the probe nozzle can be designed more freely to prevent the nozzle from blocking. This is important especially in pressurised gasification since the probe cannot be removed from the gas line during operation. 6. Particle filters. Particle filters to be used in the Guideline are thimble filters because of the higher capacity and lower pressure drop of thimble filter compared to flat filters. Only for low particle concentrations (< 20 mg/m3,) flat filters will be used as these have a lower weight so that the accuracy of particle weighting increases. The filter material is quartz. 7. Temperatures of particle filters. The temperature of the particle filter varies with the gasifier type. When sampling high temperature tars (downdraft and fluidised bed gasifiers) the preferable filter temperature is 250°C or higher in order to avoid condensation of tars on carbon rich particles. When sampling low temperature tars (updraft gasifiers) the filter temperature is 100°C to avoid polymerisation of the tars. These are draft temperatures to be evaluated. 8. Cooling and/or quenching of the producer gas before the impinger. After the particle filter, the gas must be cooled to condense the moisture and collect the tar in the impingers. Cooling can be done either by using conventional heat exchangers or by using a liquid quench system. In both cases, the condensate is collected in a condenser bottle. A liquid quench system is optional but facilitates post-sampling cleaning procedures because sampling lines are kept clean and polymerisation of reactive tar species is minimised. The use of a liquid quench strongly depends on the tar impingement solvent chosen. With a water miscible solvent exhibiting a medium volatility (like methoxypropanol) a liquid quench can be used using the same solvent. With water miscible solvents exhibiting a high volatility (like acetone) a liquid quench with the solvent cannot be used. Instead conventional cooling (with heat exchangers) or a liquid quench with water can be used. With non-polar solvents (like iso-octane) a liquid quench with the solvent cannot be used as it would lead to two-phase liquids or even formation of emulsions. Instead conventional cooling (with heat exchangers) or a liquid quench with water can be used. When a liquid quench with a water miscible solvent is used, obstruction of the gas flow through ice formation can be safely avoided. 9. Impinger bottles. One of contractors has very good experience with custom-made impinger bottles. During the development of the Guideline, the performance of standard impinger bottles (which are commonly available) will be compared with the use of modified impinger bottles. 10.Solvent. At the time of writing of this paper, a final choice for a solvent has not been made. In the Group of Contractors it is clear that dichloromethane (DCM) cannot longer be used for various reasons (toxicity, volatility, safety). During the first year of the development of the Guideline, a number of alternative solvents will be evaluated. Criteria for this evaluation are set (see below in Chapter 4). The
171
choice of the solvent will determine the sampling temperature. A temperature of -20°C is used as a start. 1 1. Backup adsorber. First tests with tar measurement following the Guideline will be performed with a backup adsorber as a safety device and to learn which volatile organic compounds andor tar compounds pass the impinger bottles. The backup adsorbed is currently an optional submodule in the Guideline. In the parallel R&D program (see Chapter 4) a number of potential adsorbing materials will be evaluated, such as activated carbon, XAD, and SPA material. In a later version of the Guideline, the solvent impingers can possibly be replaced by an adsorber tower filled with an appropriate sorbent. 12.Solvent or no solvent. For tar sampling from updraft gasifier producer gases, a solvent is needed because direct condensation of the tars without diluting media can result in further reactions (polymerisation). The reactivity is suppressed by using a diluent like an absorbing liquid or an adsorbent resin. For tars sampled from downdraft and fluidised bed gasifiers, it has to be shown whether “no solvent” can be as effective as sampling with a solvent. For the first version of the Guideline, the Group of Contractors has chosen to use a solvent. This choice will be evaluated during the running EU project. 13. Gravimetric analysis. The weight of tar is determined gravimetrically by means of solvent distillation and evaporation according to, a defined procedure (temperature, pressure and duration of distillation and evaporation). The resulting number is the amount of Gravimetric tar. 14. Compund analysis by GC. Individual compounds are determined by gas chromatography (GC). GC in combination with a Mass-Spectrometer (MS) is preferred, GC in combination with a Flame Ionisation Detector (FID) with internal standard calibration is an alternative. 15. Shortcut methods are not described in the Guideline. However, reference is made to a web page where these methods will be described and where reference to open literature is made.
4. FUTHER DEVELOPMENT OF THE GUIDELINE
The Guideline is being developed in a two year EU project in which also two partners from North America participate. The participants were named in Figure 1. The time frame of further development of the Guideline and the anticipated R&D work are described in this chapter. More information, including the inventory of required R&D and the most recent version of the Guideline, can be found on a web page (www.tarweb.net).
TIME FRAME OF FURTHER DEVELOPMENT Figure 4 gives a rough time frame of the main activities performed, and results obtained, in this project. By the time this paper appears, a first version of the Guideline will be ready. This version will be updated by the results of critical and short-term R&D needs, amongst which is the selection of the solvent. The resulting second version of the Guideline will be reviewed. The input from the Reviewers will be used together with the results of ongoing R&D activities (for example towards the interaction between particles and tar) to prepare a final version of the Guideline. Early
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2002 a workshop will then be organised on tar sampling and analysis in general and the new Guideline in particular.
R h D ACTIVITIES TO BE PERFORMED Because the EU project “Tar Protocol” is a Concerted Action project, no R&D will be performed in the project. However, it is recognised that some technical details have to be found out before the Guideline will perform satisfactory. Therefore, included in the project is the co-ordination of R&D projects executed under national funds. One of the major R&D questions to be answered is “which is the preferable solvent to be used in the Guideline?” as DCM is no longer an acceptable option. The Group of contractors has proposed a number of candidate solvents: methanol, isobutyl methyl ketone and 1-methoxy-Zpropanol. In relatively short R&D projects these candidates will be evaluated using a number of selection criteria that were proposed by the group of contractors. These selection criteria are:
Selection criteria by exclusion 0 Solvent is not toxic or carcinogenic 0 Solvent is non halogenated 0 Solvent or impurities in solvent are not present in relevant concentrations in the tar matrix 0 No additional solvent is needed for equipment cleaning Quantitative solvent selection criteria 0 t ar solubility (e.g. oxygenated organics, PAH, Gravimetric tar) 0 Collection efficiency for individual compounds Suitability for standard analytical procedures 0 Evaporation rate at 0°C and -25°C 0 Ice formation tendency 0 Freezing point of solvent Occupational exposure risk factor (e.g. as lOhr TWNvapor press.) 0 Time requirement for sampling and post-sampling procedures Availability / cost Other R&D questions that are recognised from within the project “Tar Protocol” are amongst others the effectiveness of tar adsorbents, whether tar polymerisation occurs at the proposed sampling and post-sampling conditions, and the interaction between particles and tar during sampling and post-sampling. The complete R&D inventory can be found on the web site (www.tarweb.net). Any institution or company having knowledge, performing R&D or planning to perform R&D is invited to contact the authors of this paper to exchange knowledge andlor to co-ordinate or perform R&D activities.
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I
REVIEWERS
CONTRACTORS
OUTSIDE EU PROJECT
2000 April August
2001 February
, \
lCoordination of RStDk---
.i”-----.1 R&D projects
.L
Guidelioe version 3
December 2002
I
February
*
Workshop on Samptig and Analysis of Tars and Particles Y Final Guidcllde
I
I
Figure 4 Time frame of main activities of the EU project ‘Tar Protocol’
5. REFERENCES 1. T.Milne, N.Abatzoglou and R.J.Evans: Biomass gasifier “tars“: their nature,
2. 3. 4.
5.
6. 7.
formation and conversion, Golden, CO (USA), NREL, NREL/TP-570-25357.68 p. (1998). K.Maniatis and A.A.C.M.Beenackers: Introduction. Tar Protocols. The IEA Bioenergy gasijkation task. Biomass and Bioenergy 18 (1) 1-4 (2000). N.Abatzoglou, N.Barker, P.Hasler and H.Knoef: The development of a dra@ protocol for the sampling and analysis of particulate and organic contaminants in the gasfrom small biomass gasqers. Biomass and Bioenergy 18 (1) 5-17 (2000). P.Simel1, P.Stahlberg, E.Kurkela, J.Albrecht, S.Deutsch and K.Sjostrom: Provisional protocol for the sampling and anlaysis of tar and particulates in the gas from large-scale biomass gasifiers. Version 1998. Biomass and Bioenergy 18 (1) 19-38 (2000). J.P.A.Neeft, H.A.M.Knoef, U.Zielke, K.Sjostr6m, P.Hasler, P.A.Simell, M.A.Dorrington and C.Greil: Tar protocol. Development of a standard method for the measurement of organic contaminants (“tar”) in biomass producer gases. In: Biomass for energy and industry. Proc. 1st World conference and exhibition, held in Sevilla (Spain), 5-9 June 2000 (Eds. S.Kyritsis et al.), EnergiaTA, Florence (Italy) (2000). U.Zielke, P.D.Kellberg, H.Knoef, P.Hasler and P.Simell: Parallel measurements of tar and particulates, Aarhus (Denmark), Danish Technological Institute, 46 p. + Appendices ( I 999). C.Brage, Qizhuang Yu,Guanxing Chen and K.Sjostrom: Use of amino phase adsorbent for biomass tar sampling and separation. Fuel 76 (2) 137-142 ( 1997). 174
8. H.A.M.Knoef: The UNDPIWorld Bank monitoring program on small scale biomass gasifiers (BTG's experience on tar measurements). Biomass and Bioenergy 18 (1) 39-54 (2000). 9. A.Das: Contaminant testing method for biomass gasifier engine systems, Golden (USA), Colorado School of Mines ( I 985). 10. R.A.Knight: Experience with raw gas analysisfrom pressurized gasification of biomass. Biomass and Bioenergy 18 ( I ) 67-77 (2000). 1 I. P.Hasler and T.Nussbaumer: Kontinuierliche Teermessung mittels FID bei einem Holzvergaser, Zurich (Switerland), Verenum, 63 p. (2000). 12. O.Moersch, H.Spliethoff and K.R.G.Hein: Tar quantification with a new online analyzing method. Biomass and Bioenergy 18 ( I ) 79-86 (2000). 13. N.Abatzoglou: Raw gas contaminants: an overview. Lecture held at the IEA Thermal Gasification on Biomass Task, EC, and USDOE Meeting on Tar Measurement Protocol, Brussels, March 18-20 (1998). 6. ACKNOWLEDGEMENTS The European Commission, The Netherlands Agency for Energy and the Environment (NOVEM) the Swiss Federal Office of Education and Science, the US Department of Energy and National Resources Canada are greatly acknowledged for their (financial) contribution to the work presented in this paper.
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Modelling the Characteristics of the Endothermic Reaction Potential of Tar for Flue Gas Clean-up in Advanced Thermochemical Conversion Processes Georgios Taralas Energy Services Consulting and Research (ESCOR), 9 K. KIystalli Str, 151 26, Maroussi, Athens, Greece.
ABSTRACT As a step in the application of the cracking of tar in fuel gas amelioration the characteristics of the endothermic reaction potential of tar was studied e.xperimentally and theoretically. In this context, however. due to the structural complexity of tar andor tarry constituents in fuel gas well defined hydrocarbons as tar model compounds were applied with inexpensive and readily available materials (dolomites, dolomitic magnesium oxide quicklime [CaO]). The effects of operation condition on extent of hydrocarbon conversion. gas product composition. and corresponding endotherm of the reaction potential have been e.qlored. The results obtained in this work provide a basis for future considerations of catalflc tar cracking.
[Me].
INTRODUCTION As advanced thermochemical conversion processes (i.e., gasification, pyrolysis, incineration, etc) into the twenty-first century, increasingly severe demands are being placed on tar minimization, structural and conversion capabilities of the device design, dried biofuel economy, and pollutant-emission controls. Therefore. the designed processes should provide a fuel gas behaved as the replacing fuel viz., high-value fossil fuels such as natural gas, distillate oil, in terms of heating value, apparent thermal efficiency [1.2], and other physical and chemical gas properties [2]. The product gas with a severe heating value is an important fuel for gas turbines and in designing lugh velocity burners demanding a gas of rapid burning rates, short ignition delay times. good flame stability as well as an increased flame adiabatic temperature and high flame speed. A iugh flame speed will. in general improves flame stability and creates a more intense and shorter flame. Biofuels being feedstocks result in a high yield of permanent gas (i.e., H2, CO, CO?, CH,, etc), volatiles, char, and condensible hydrocarbons i.e., tar, due to their higher hydrogedcarbon. and oxygerdcarbon ratios than older fossil fuels. However, if the tar is allowed to condense or polymerise when the produced he1 gas is cooled down considerable problems with equipment contamination can result. Power generation requires high levels of gas clean up especially in gas turbines and fuel cell systems. Also associated with the Laboraton, of Food Chemistn, and Techriologv, Department of ChemistT, University of Ioannirra, GR-45 I I0 loannina, Greece
176
However. volatile species and tar offer potential for increased fuel gas production by undergoing chemical decomposition (cracking) on low cost mineral materials such as calcined dolomite [CaMg(O)-]. magnesites [MgO], calcined limestone. quicklime [CaO]. etc [3,1.5.6]. Although mineral materials might have some drawbacks, such as thermal instability, loss of surface area by sintering and Severe elutriation rates they are very cost effectiveas catalysts and their d~sposalis not considered to.xic waste [7]. In this study. the endothermic reaction potential of n-heptane was studied. Cyclic hydrocarbon (cycloalkane, c-C,,H2,,.n=6) and saturated n-heptane (alkane, C,,H2,+:,, n=7) were used as tar surrogate molecules (without n-electrons). However. the work does refer to Wiirzburg provisional definitions of condensable (at reactors outlet conditions) tars [8.9] and the activities undertaken under the EU project ‘Tar Protocol’ [lo] and only divide the whole tar material in substances (explicit behaviour of flue gas)consisting of saturated and unsaturated molecules [3]. In order to have confidence that a catalytic tar cracking process will achieve the desirable level of performance. trial model substance cracking may be required. These trials may be conducted using limited quantities of condensable tar, specdic hydrocarbons, or by using surrogate materials. which wlule possessing less hazard materials, place similar demands on cracking parameters as does homogeneous tar that would be used as a feedstock. ENDOTHERMIC CRACKING OF COMBUSTIBLE VOLITILE MATTER AND TAR In general. thermal and catalytic steam cracking of normal and cyclo-alkanes used as tar model compounds may follow several different reaction paths and produced hydrogen and a spectrum of saturated and unsaturated hydrocarbons products [ 1 1,121. Catalytic steam cracking of normal and cyclo-alkanes on these mineral materials to form olefins and hydrogen is strongly endothermic and can be performed at temperatures that are compatible with existing materials [3,12]. Unlike the saturated ring (naphthenic) endothermic hydrocarbons. which react to form aromatics, catalytic craclung results in the formation of olefinic and low-molecular-weight parattinic species [l3]. Consequently, the aim of t h s investigation was to e.uploit the cracking reactions concerning the saturated hydrocarbons i.e.. n-heptane and cyclohexane. Their product composition may be affected by the reaction temperature, which corresponds to Merent amounts of heat being absorbed. As a result a sink of heat can occur wluch comes both from the physical heating of the hydrocarbodtar (raising its temperature) and from a heat absorbing (endothermic reforming) chemical reaction. Furthermore, a measure of the endotlierm can be determined through thermochemical calculations based on product analysis and measurements of the extent of conversion [ l l ] . The energy absorbed as a result of the chemical reactions is the difference between in enthalpies prior to the chemical kmetics, calculated over the chemical time step at constant pressure where To is the temperature at the start of the time step. THERMAL CRACKING OF CYCLOHEXANE AT TIME ZERO
The product distribution of cyclohesane pyrolysis under given reaction conditions may depend on the temperature and convcrsion. Taralas [ 3 I concluded that the selectivities
177
of gaseous products increased with increased conversion. Earlier works [ 14,151 agreed that the product distribution of cyclohexene decomposition consists of a considerable amount of ethylene, butadiene, propylene, methane, and hydrogen. Small quantities of ethane, propane, and butenes are also produced. Traces of benzene, cyclohexene, and toluene indicate that some dehydrogenation reactions are occurring to a slight extent. Many hypotheses were made concerning the mechanism of the decomposition of cyclohexane. Moreover, [16,17] described the elementary steps of the thermal decomposition of cyclohexane. In their elementary radical steps the formation of methane was included. The e.qerimental results obtained by Taralas [3] at approximately 700 "C permits to propose in this communication the primary stoichiometric equations at time zero:
Cycloheme is fairly easily dehydrogenated into benzene, and even at very low extents of reaction, stoichiometry reaction (6) can be replaced by the secondary stoichiometry reaction (12). For cyclohexane, the constituents are (apart from cyclohexane) hydrogen. methane, ethane, ethylene, acetylene, propene, 1-butene, 1,3-butadiene, cyclohexene, and benzene [3]. However, one can check that the equations written are independent, using the Jouguet [I81 criterion, (i.e., yl = n-co). In this criterion, the number of the independent constituents, yl, for a chemical system is equal to the required constituents, n, fi.e., H2, Cl, C2, c2H6. C3, C,, C,. c'c6, Cd, subtracting the number of independent stoichiometric equations. o. Moreover, by combining the Brinkley's [ 191 criterion, o = n-yl, one can calculate the number of independent stoichiometric equations. In this criterion the number w of independent constituents of a chemical system is equal to the rank of the matrix of the indexes of the elements in the formula of the constituents, hence o = n- y = 9-2 = 7. Kinetic considerations have led to write 11 stoichiometric equations. It must be checked by Jouguet's criterion [18] that these 11 stoichiometric equations are independent. In the case of cycohexane the stoichometric equations are 2, 3, 4, 8, 9, 11, and 12. Thus these seven stoichiometric equations are independent and describe the decompositionof cycloheme. n-HEPTANE
Depending on the temperature at which the reaction begrns and the products formed. the endothermicity of the decomposition can vary substantially. As shown for nheptane in Figures 1 and 2 increased unsaturation of the product results in increased heat absorption, which require hgher reaction temperatures for comparable conversions [ 111. 178
104
i
1
0 &
103
&
0
/ &
300
400
500
600
700
800
900
lo00
1100
Temperature, "C
Fig. 1 Endothermic cracking of n-heptane.
be
-
/ _ - - -
Fig. 2 Endothermic cracking of n-heptane.
shown in Fig. I , and aromatics n-C7HI6 + H2 + c&6
f
C7H8
shown in Fig. 2 tends to decrease the endotherm, but increase the equilibrium fractional conversion at a gwen temperature [ 1 1,3]. Production of alkenes such as ethylene n-C7Hi6
-b
H2
f
c2H4
(15)
179
depicted in Fig. I leads to a substantial heat absorption, whereas formation of acetylene n-C7HIb + H2 + C2H2 (16) shown in Fig. 2 leads to even greater heat absorption but requires reaction temperatures above 800 “C. In addition, a multiplicity of low-molecular-weight products is desirable. The formation of a large number of products may increase the entropy change during the reaction, which in turn makes the reaction more favorable thermodynamically [ 1I]. In addition to providing high heat absorption, the low-molecular-weight olefins would result in a fuel gas with excellent combustion characteristics. Severe amounts of alkanes (methane) and aromatics are acceptable to enhance the thermodynamic driving force for high conversion levels at a given temperature.
THE CATALYSTS USED IN THIS WORK Different media arise in such diverse contexts as catalyuc beds and sedimentary rocks [4]. However, in this case the mineral rocks, dolomitic magnesium oxide [MgO], quicklime [CaO] and dolomite [CaMg(CO3)2] are commercial quarried products and were used as the catalysts [ 111. The geological evolution of carbonate rocks like the crystalline dolomite begins with high porosity packing of CaC03 grains, with initial grain sizes in the 1-10 pm m g e . The process of dolomitization begins with the nucleation, at random centers (most often at the surface of the CaC03 grains) of CaMg(C03)2rhombohedral crystals which, over millions of years, grow into grains whose size is of order tens of micrometers. The Mg ions originate in the brine that saturates the pore space. The replacement of CaCO, by CaMg(CO& involves a volume reduction that introduces small-scale (intercrystalline)porosity at the same time that the original, larger scale porosity is being reduced by compaction and cementation. When the evolution is completed the solid phase is entirely constructed of random CaMg(CO& “rhombs” and the spectrum of pore sizes is quite broad [20]. However, during calcination of dolomite CaO crystallized independently from MgO, and form grains attached to the surface of much lager MgO grains [21]. BENCH-SCALE TESTING CATALYTIC CRACKING OF CYCLOHEX4NE ON CALCINED DOLOMITE
In catalyst systems considered for tar cracking and/or minimization for thermal gasification applications, two methods for supporting the catalyst have been applied. The first is a packed bed (catalytic reactor vessel) in which the catalyst is formed into crystallized granular (i.e., dolomiteparticles, etc), and the crystallizedgranularmaterial is placed in the fuel gas production passage. The second technique involves placing of the mineral material as the catalyst directly in to the bed of the gasifier [7]. In this work, however, tests were designed to allow screening of hydrocarbodcatalyst combinations through measurements of endotherm of reaction, extent of conversion, and identification of reaction products. However, the atmospheric-pressure bench scale test apparatus has been described elsewhere [ 111. Parametric studies of the catalytic decomposition of n-heptane and cyclohexane were conducted to provide their stoichiometric analysis. The proposed independent stoichiometricequations describing the catalytic decompositionof cyclohexane are: 180
CATALYTIC CRACKING OF HEPTANE ON CALCINEDDOLOMITE
The packed bed configuration has been tested in the apparatus, differing in the mineral materials. The heat input to the reactor have been used, viz., resistive heating. Resistive heating, that is, heating by an electric current, has the characteristic that the heat transfer rate is nearly uniform for a given tube cross-sectional area along the length of the reactor and the wall temperatures vary fiom inlet to exit. Endotherm has been computed from the extent of the conversion and the product composition [131 with the proposed in this work, stoichiometricequations at time zero:
The product chemical composition analysis allows assessment of the dolomite catalyst selectivity to be made and is used to compute the heat of reaction and thereby provide a measure of the endotherm. The sink of heat of n-heptane is associated with hydrocarbon conversion to the equilibrium composition of the products where any variance could also be attributed to kinetics and differences in reaction products [13].
TEST RESULTS OPERATIONAL CONDITIONS
Tests have been conducted primarily with calcined dolomite [CaMg(Ot)] according to previous studies of the subject [11,12,21,22], quicklime [CaO] and dolomitic magnesium oxide WgO] as the catalysts, however, a limited number of test have been performed without catalysts (exploring pure thermal cracking). In this article it is reported new experimental data and it is also reexamined previously obtained results for cracking of highly saturated hydrocarbons. The tests were conducted at Gaseous Hourly Space Velocities (GHSV)of up to 4000 h' and atmospheric pressure. The Table 1 represents the gas analysis results showing the attractive product distribution for n-heptane conversion with calcined dolomitic magnesium oxide WgO] and quicklime [CaO] catalyst compared to that without a catalyst (thermal cracking).
181
Table I. Operating Conditions and Comparison of Main Product Selectivitiesand Conversion of the Thermal (empty reactor) and Catalytic Cracking of n-heptane.
Dolomitic Empty Reactor MgO --973 ---973-Temperam, 6) Conversion Xc,(mol.%) 20.8 33.3 27.2 40.7 Elapsed Cracking Time, (min) 229 162 183 164 Coke on Catalyst, (wt%) 1.02 M Product Selectivity 63.9 100.1 39.0 104.8 48.2 74.9 52.6 59.1 152.3 193.7 161.1 168.9 6.37 6.15 8.27 8.56 44.1 44.9 40.7 42.5 24.9 15.5 20.2 24.4 0.0 1.87 1.84 0.45 10.1 4.49 6.74 0.25 0.0 3.74 0.39 0.49 3.74 0.15 2.63 1.53 0.0 0.75 0.0 0.51 0.0 0.0 3.90 0.0 0.0 0.0 0.06 0.03 0.0 0.0 0.0 0.0 co2 0.0 0.0 1.74 12.36 Total cracked gas 353.6 439.1 339.0 431.1 na: not analy&d
--
Qucklime CaO 973--20.7 3 1.8 I -
112 0.65
165
50.5 89.9 147.5 6.06 41.5 21.6 1.38 9.19 1.42 3.46 0.11 0.0 0.05 0.0 7.65 380.2
108.0 65.7 163.1 9.02 41.7 22.1 1.23 2.97 0.29 0.97 0.0 5.00 0.03 0.0 19.32 439.4
M
Thermal cracking: Equivalent space time, ~,=0.28-0.55 s; Partial pressure of nheptane, P,hP,,,,,=3.0-4.7 kPa; Partial pressure of steam, PHZ0=18.7-30.9kPa; Ratio, H~O/n-heptune=O.89-0.95mol/mol. The definition of the equivalent space time rq has been described elsewhere [131. Catalytic cracking: Time factor, ~ 0 . 0 5 - 0 1 3kgh/m'; Partial pressure of n-heptane, P,,. hprupronr=2.9-3.9 kPa; Partial pressure of steam, P~20=17.8-31.5 Wa; Ratio, HZO/nheptune=0.89-1.20 moYmol. The time factor t and conversion Xc have been defined previously [13 J. In the absence of the catalysts, the products represent the result of thermal cracking, yielding high-molecular-weight species that are primarily saturated, and may corresponded to a low endotherm. With the dolomitic MgO and quicklime CaO catalyst, the products are substantially lower in molecular weight and are primarily unsaturated, making them more desirable from the point of view of both high endothem and with favorable combustion characteristics. Cracking of hydrocarbons is known to occur at relative high temperatures even in the absence of any catalyst. Moreover, thermal cracking of hydrocarbons for a-olefin production is carried out in different reactors (12,131. The reactor design can be optimised in order to obtain a favowable product distribution. However, the transfer of heat into the hydrocarbon may be accomplished in a catalytic reactor with constant 182
temperature, Within a reactor flow passage, the mathematical model of the rate at which the tar-surrogated hydrocarbon is cracked may depend on the balance among the heat transfer throughout the wall, the mass transfer of reactant to the catalyst coupled with the chemical kinetics associated with the cracking reaction:
dT a2T iaT c (l-&)-P=A (++-')+hp~p(Tg-Tp)+U
p P
p
a6
(33)
rdr
The T, and T, are temperatures of the packed particles and the fluid, reswvely, U is the rate of such an individual reaction. C, is the specific heat, p, is the density, E is the void fraction in bed, 8 is the time, a is the specific area of a particle, Uo is a coefficient which is dependent on the temperature, q is the energy of activation per mole, R * is the gas constant, and T is the absolute temperature. The initial and boundaq conditions expressed in polar coordinate system (shape of reactor) are
0=0, O I r S R , OIzIZ,; T p = T B + O , r=O, OIZSZ,; 8Tplar=0 0 >- 0, r = R, 0 Iz IZ,;
- Ap(aTp/a) = hw(Tp- T,)
where R is the radius of the packed bed, 2, is the height of the packed bed, h, is the apparent heat transfer coefficient at the wall, and T, is the temperature of the wall of the inner tube. The effectiveness of the calcined dolomite in the conversion of the hydrocarbon to desirable products is evident in the product analysis shown in Figures 3, 4and5.
= c b
N-
3
= 8
Fig.3 n-C7H16 selectivity product distribution at 750 "C.
183
Gaseous Product
These Figures represent the selectivity (moles of produd100 moles of hydrocarbon decomposed) results showing the gaseous product for n-heptane and cyclohesane with dolomite catalyst compared to that without a catalyst (bare tube). Using dolomite catalysts, the products are substantially lower in molecular weight and are primarily unsaturated, making them more desirablefrom the point of view of both endotherm and favorable combustion characteristics: The product gas arises, such as H2, CH+ C2H4, etc, is an important fuel for boilers, gas turbines and internal combustion engines. The results indicated a conversion approaching 90 percent at 800 "C [11,3], with a corresponding maximum endotherm (i.e., the heat absorbed in the cracking reaction of n-heptane equilibrium calculation: n-C7H16 + H2+CH,+C~H2+CtH,+C2H6+ cfl&cfi8+cfi8) of apprnximately 680 kJ/kg.
8
Gaseous Product
Fig. 4 n-C7HI6selectivityproduct distribution at 800 "C.
CONCLUSIONS Endothenns and attractive reaction products have been achieved through the endothermic cracking of tar model compounds (without x-electrons) using inexpensive calcined mineral materials as the catalysts. These results provide the basis for the continued development of tar cracking systems, which enable thermochemical processes operation with practical catalysts. The following specifk conclusions can be drawn from the results of this work. 0
Sink of heat canbe realized by catalytic craclung of saturated hydrocarbons as tar model compounds. Endothermic reactions can be achieved using inexpensive dolomite stones as catalysts. Catalytic cracking can be carried out with good product selectivity. 184
Any practical system for reducing the condensable hydroahon constituents (tar) yield will benefit from a catalytic surface to achieve adequate endotherm and produce desirable reaction products.
400
O
-
n
Fig. 5 Cyclohexane selectivityproduct distribution at 700 "C.
ACKNOWLEDGEMENTS
This work was carried out under the JOULE III-RES Program of the EC DG-XI, Project No. JON-CT98-0306.
REFERENCES 1. Gil, J., Aznar,M. P., Caballero, M. A., Frances, E., Corella, J. (1997) Biomass Gasification in Fluidized Bed at Pilot Scale with Steam-Oxygen Mixtures. Product Distribution for the very Different Operating Conditions. Energy and Fuels, 11 (6), 1109-1118. 2. Gil, J., Corella, J., Aznar,M. P., Caballero, M. A. (1999a) Biomass Gasification in Atmospheric and Bubbling Fluidized Bed: Effect of the Type of Gaslfylng Agent on the Product Distribution. Biomass & Bioenergy, 17, 389-403. 3. Taralas, G. (1996a) Cyclohexane-steam Cracking Catalysed by Calcined Dolomite, In: Developments in Thermochemical Biomass Conversion, (ed. A. V. Bridgwater and D. G. Boocock), pp. 1086-1100. Blackie A & P, London, UK, VOl. 111,
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4. Namhez, I., Corella, J., Orio, A. (1997) Fresh Tar (froma biomass gasifier) Elimination over a Commercial Steam Reforming Catalyst. Kinetics and Effect of Different Variables of Operation. Ind. Eng. Chem. Research, 36 (2). 3 17-327. 5 . Delgado, J., Aznar, M. P., Corella, J. (1996) Calcined Dolomite Magnesite and Calcite for Cleaning Hot Gas from a Fluidized Bed Biomass Gasifier with Steam: Life and Usefulness. Znd. Eng. Chem. Research, 35, 3637-3643. 6. Simell, P. A,, Lepprilahti, J. K., Kurkela, E. A. (1995) Tar-decomposing Activity of Carbonate Rocks under high COz Partial Pressure. Fuel, 74, No. 6, 938-945. 7. Gil, J., Martin, J. A., Aznar,M. P., Caballero, M. A., Corella, J. (1999b) Biomass Gasification with Air in Fluidized Bed: Effect of the In-bed Use of Dolomite under Different Operation Conditions. Ind. Eng. Chem. Research, in press. 8. Abatzoglou, N. Barker, N., Hasler, Ph., Knoef, H. (1998) The Development of a Draft Protocol for the Sampling and Analysis of Particulate and Organic Contaminants in the Gas from Small Biomass Gasifiers, Provisional Protocol, version 98. Prepared of the IEA Bioenergy Agreement. Version 1.I for discussion and comment. 9. Albrecht, J., Deutsch, S.,Kurkela, E., Simell, P., Sjiistrom, K. (1998) The Development of a Draft Protocol for the Sampling and Analysis of Particulate and Organic Contaminants in the Gas from Small Biomass Gasifiers Prepared of the I . Bioeneray Agreement. Version 2 for discussion and comment. 10. Neeft, J., Hasler, Ph. (2000) Draft Minutes of the Project Tar Protocol Contract ERK6-CT-1999-20002. First World Conference and Exhibition of Biomassfor Energy and Industry, Sevilla, 6 June 2000, Spin. 11. Taralas, G., Vassilatos, V., Delgado, J., Sjostriim, K. (1991) Thermal and Catalytic Cracking of n-Heptane in Presence of CaO, MgO, and Calcined Dolomites. Can. J. Chem. Eng., 69, 1413-1419. 12. Taralas,G. (1996b) Catalytic Steam Cracking of n-Heptane with Special Reference to the Effect of Calcined Dolomite. Ind. Eng. Chem. Research, 35, 2 121-2 126. 13. Taralas, G. (1999) Modelling the Influence of Mineral Rocks, Active in Different Hot Gas Conditioning Systems and Technologies, on the Production of Light a-Olefins. Can. J. Chem. Eng., 77, 1205-1214. 14. Levush, S.S.,Abadzhev, S. S.,Shevchuk, V.U. (1969) Pyrolysis of Cyclohexane. Nefekimiya, 9,716-72 1. 15. Aribike, D. S.,Susu, A. A., Ogunye, A. F.(1981) Mechanistic and Mathematical Modeling of the Thermal Decomposition of Cyclohexane. Thennochim.Acta, 51, 113-127. 16. Brown, T. C., King, K. D., Nguyen, T. T. (1986) Kinetics of Primary Processes in the Pyrolysis of Cyclopentanes and Cyclohexane. J. Phys. Chem. 90,419-424. 17. Billaud, F., Chaverot, P., Berthelin, M., Freund, E. (1988) Thermal Decomposition of Cyclohexane at Approximately 810 "C. Znd. Eng. Chem. Research, 27,759-764. 18. Jouguet, E. (192 1) Thermodynamique Chimique. J. Ec. Polyrech. (Paris), 21, 61. 19. Brinkley, S.R (1946) Critere de Brinkley. J,Chem. Phys. 14,563-586. 186
20. Crossley, P. A., Schwartz, L. M., Banavar, J. R. (1991) Image-based models of porous media: Application to Vycor glass and carbonate rocks. Appl. Phys. Lett, 59 (27), 3553-3555. 2 1. Taralas, G. (1998) CatalyUc Steam Pyrolysis of a Selected Saturated Hydrocarbon on Calcined Mineral Particles. Can. J. Chem. Eng., 76, 10931101. 22. Taralas, G.,Corella, J., Kakatsios, X.(1999) Probing the Gaseous Product Yields from the Py-rolysis of Tar Model Compound for Biomass, Sludge and Waste Advanced Thermal Gasification,In International Conference of Disposal and Utilisationof Sewage Sludge TreatmentMethods and Application ModalitiesJAWQ, pp. 543-550, Athens, Greece.
187
Fundamental Fluid-Dynamic Investigations in a Scaled Cold Model for Biomass Steam-Gasification R. Kehlenbeck*, J.G. Yates* and R. Di Felice’ *Department of Chemical Engineering, University College London, Torrington Place, London WClE 7JE, UK ‘Istituto Di Ingegneria Chimica E Di Processo, Universita Degli Studi Di Genova, Via Opera Pia 15, I6145 Genova, Italy
ABSTRACT: The hydrodynamic behaviour of a one-fifth scale cold model of a 500 kWmpilot plant has been studied. The plant is based on the circulating fluidized bed (CFB) principle and consists of a bubbling bed steam gasifier and a riser combustor. It is designed to produce a hydrogen-rich gas to be fed to a he1 cell for electricity generation. The cold model was constructed on the basis of established fluid-bed scaling laws and was built of perspex to enable the bed behaviour to be observed directly. Solids circulation rates and gas cross-flow between combustor and riser were investigated. A novel scaling parameter believed to be generally applicable to flow in risers was developed on the basis of these experiments. INTRODUCTION This work forms part of a project to develop a pilot-scale biomass gasifier for the production of a hydrogen-rich gas to be fed to a phosphoric acid fuel cell for electricity generation on a local scale. The plant which is currently under construction in southern Italy employs a bubbling fluidized bed steam-gasifier operating at around 8OOOC and a riser combustor at 900°C. The bed material is sand which circulates between the two units and transfers the heat necessary for the endothermic gasification reaction. The process is a development of two previous studies reported in references (1) and (2). The work at UCL was directed at understanding various aspects of the hydrodynamics of the plant based on studies of a scaled model operating at ambient conditions. Wellestablished scaling rules were used to achieve hydrodynamic similarity between the two systems. Several attempts to develop scaling parameters for fluidized beds have been reported but probably the best known and most widely used are those derived by Glicksman et al., extensively reported in (3). On the basis of the governing equations of conservation of mass and momentum of fluid and particles, they derived a set of nondimensional parameters which must be matched in order to obtain hydrodynamic similarity between a model and a full-scale plant. Their so-called “full” set of scaling parameters is:
188
containing the Archimedes number, the solid to gas density ratio, the Froude number and the Reynolds number based on the bed diameter, D. C, is the dimensionless solid circulation rate calculated by:
c, =-,
G S
P s ‘UO
The bed geometry defrnes the scale factor and I$ is the sphericity of the particles. By matching this set of parameters hydrodynamic similarity is almost obtained as has been confirmed by several experimental investigations (4). In addition, Glicksman et al. (5) derived a “simplified” set of scaling parameters as in (1) but omitting the Archimedes and Reynolds numbers and adding the ratio between the superficial and minimum fluidization velocities, U”:
This simplified set of parameters provides greater flexibility in choosing the operating conditions in the model compared to the full set of parameters. In the work to be described here, a one-fifth scale cold model of the circulating fluidized bed (CFB) pilot plant was built on the basis of the full set of scaling parameters. Using the cold model, the influence of the operating parameters on the hydrodynamics of the system was investigated and we present in this paper the results obtained for the solid circulation rate (SCR), which led to a novel scaling parameter for circulating fluidized beds, and the gas cross-flow between the riser and the bubbling bed. EXPERIMENTALEQUIPMENT The basic design of the cold laboratory model is shown in Figure 1. The main parts were made of perspex in order to observe the fluid dynamics of the system visually. It can be divided into three main sections: a bubbling-bed gasifier (DW,= 160 mm), a = 54 nun) and a downcomer (Ddc= 24 mm) which is separated riser combustor (Dcomb from the gasifier by a siphon. The basic idea for the pilot plant is that the char resulting fi-om the steam gasification of biomass is transported together with the circulating bed material (olivine) into the combustor (fluidized with air) where the char is partially burned. The exothermic combustion heats the olivine particles which re-circulate via the downcomer into the gasifier. The hot bed material provides the heat needed for the endothermic gasification. The bed inventory contains a quantity of a novel nickel-onolivine catalyst which lowers the temperature of the tar-forming reactions as well as partly reforming the methane product to hydrogen. The volumetric flowrate into the windbox of the gasifier is indicated with V*w,; the combustor has two air inlets, a primary, indicated with V*pm and a secondary, indicated with V*=. The flowrate into the primary inlet should be as low as possible in order to avoid gas cross-flow into the gasifier which would decrease the quality of the product gas by increasing its nitrogen content. Thus, with V*P,,,,,particles are lifted up to the secondary air inlet where the 189
major air is injected. The ratio of V*pnmN*sdc is of the order of 15. In order to make the particles circulate fiom the downcomer into the gasifier the syphon is fluidized with the ’ ~ ~ ~ ~ flowrate v
Figure I Basic design of the cold model For the operation of the cold model ambient air was chosen as fluidising gas for the riser, and in order to match the required density ratio the gasifier was fluidized with a gas mixture of 55:45 % helium to air. Spherical bronze particles with a density of 8900 kg/m3and a mean particle diameter of 180 p were chosen as bed material. The geometric scale factor between the two units was five and the normalised particle size distributions of the bronze particles and the sand used on the plant were comparable. The one scaling parameter it was not found possible to match was the particle shape factor 4. In the cold model the particles were spherical (el) whereas sand particles are normally considered to be “broken solids” for which 9 0 . 6 3 (6); in operation in the CFB however attrition will be likely to increase this to nearer unity. EXPERIMENTAL RESULTS
190
The two main hydrodynamic features studied were the solid circulation rate and its variation with the operating parameters, and the cross flow of gas between the risercombustor and the bubbling-bed gasifier. It is of course essential to minimise the extent of cross flow in order to avoid as far as possible the introduction of air into the gasifier product gases. SOLID CIRCULATION RA TE The solid circulation rate is one of the basic operating parameters of a circulating fluidized bed since it influences not only the temperature difference between the different sections of the reactor but also the residence times of particles and therefore reaction and conversion times. For the gasification of biomass for example high temperatures are required in the gasifier for the tar pyrolysis as well as for the efficiency of the catalysts used in the pilot plant. In the cold model the solid circulation rate (SCR) could be determined by switching off the gas flow in the syphon of the downcomer thereby preventing particles fiom flowing over into the gasifier; a fixed particle bed with a clearly defined surface then gradually filled the downcomer. Measuring the time, t, required for the particle bed to rise with respect to the bed height, h, the solid mass flow, m*, was determined as:
and the mass flux,G ,as:
where DQ,is the diameter of the downcomer and D- the diameter of the riser (combustor). In addition to bronze, sand, salt and amberlite resin particles were used. The measured particle properties are listed in Table 1.
Table I Measured Particle Properties for all Particles Used
II
4 twl Ps[kg/m31 Eb
[-I
u, W S l
Bronze 180 8900 0.4 3.4
I
Sand
I
170 2622 0.44 1.4
Salt 348 2150 0.42 2.3
1
Amberlite Resin 230 I 860 1480 0.40 I 0.44 1.2 14.37
Three parameters of possible influence on the circulation rate were investigated: (1) the volumetric flowrate into the gasifier, VWi. (2) the volumetric flowrate into the riser, Voph and V *,. (3) the total mass load in the unit, m.
191
It was found that the SCR was independent of the gas flowrate into the gasifier over the whole range of interest (1 80 - 320 Ymin) and that the dominant influence was the gas flowrate into the riser. The results obtained for G, are shown in Figure 2 as a function of the superficial velocity, uo, in the riser; no difference in circulation rate was observed by varying both the primary and secondary gas flowrates into the combustor as long as the superficial velocity, u,,, calculated as:
was kept constant.
1201 Bronze. total mass load: 5kg 0 7kg 0 9kg 0 11 kg A 15kg
v
z
H
Ic
r
i
20
0 2.0
-re
I
2.5
3.0
3.5
4.0
4.5
5.0
superficial velocity u,, [ m / ] Figure 2 Solid mass flux G, as a function of the superficial velocity u,, in the riser
Horio (7) has stated that in a CFB combustor the solid flux should be of the order of 10 to 30 kg/m2s and it is evident fiom Figure 2 that this value is reached easily with the model investigated here. Furthermore, it can be seen that the solid mass flux increases strongly with increasing superficial velocity over the range investigated and also that at a given value of uo, G, increases with increasing total mass load in the system, m. A NOVEL CORRELATION FOR PREDICTING SOLID CIRCULATION RA TE A graph of similar form to that shown in Figure 2 was obtained when the dimensionless solid circulation rate (defined in equation 2) was plotted as a fimction of the
192
dimensionless velocity ratio U defined as the ratio of superficial velocity in the riser, u,, to the terminal fall velocity of the bed particles, u, (Figure 3).
V
M417 0 M=Q2 0 M430 0 M437
. . .
I
I
A M4.51
0.7
0.8
0.9
1.0
1.1
12
1.3
1.4
1.5
dimensionlessvelocity ratio u= u& [-] Figure 3 Dimensionless solid circulation rate of bronze particles as a finction of a dimensionless velocity The dimensionless parameter M" in Figure 3 is the ratio of the total mass of particles in the unit as a whole to the mass contained in the riser when filled to its maximum value:
When the dimensionless solid mass flux was plotted against the total mass load in the bed the relationship was found to be linear (Figure 4)and from this and on the basis of the data plotted in Figures 2 and 3 a novel scaling parameter, the dimensionless mass turnover, M, was defined as follows:
193
Plotting this group against the dimensionless velocity, U, all data points obtained for bronze at different mass loadings fall onto one curve and these data together with those obtained for sand, salt and the resin are plotted in Figure 5. The dimensionlessmass turnover can be expressed as a second order function of U as follows:
M = a o U 2 -1.6a0U+0.64a0
(8)
with ao = 0.0015 = const.
(9)
9080-
70:
B so-
&
*
50-
total mass load in the bed m pg] Figure 4 Solids mass flux in the riser as a function of the total mass load in the system All the experimental data can now be approximated by a very simple equation (equation 8) containing a single parameter, %, and this is shown as a continuous line in Figure 5 from which it can be seen that with a small number of exceptions the data points are well approximated. The reason for the deviation of the results for the sand bed shown in Figure 5 is believed to be due to differences in the particle size distribution of the material. Size analysis showed that the sand contained a higher proportion of both large and small particles than the other powders studied and thus it 194
is to be expected that solid circulation will start at a lower velocity and proceed to higher velocities than for materials with a more restricted PSD. The results give a clear indication of the importance of PSD in the scaling relationships referred to above. With the circulating fluidized bed used here, solid circulation measurements were carried out up to a maximum solid flux of about 120 kg/m2s; exceeding this value the downcomer gradually filled with particles and thus setting a limit to the measurements.
0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 1.6 1.8 2.0
dmensionlessvelWU[I] Figure 5 Dimensionless mass turnover M as a function of dimensionless velocity U
GAS CROSS-FLOW
This was determined by injecting a tracer gas (COz) into the two air inlets to the riser and measuring by means of an infia-red analyser the CQ content of the outlet gas stream fiom the bubbling bed gasifier. Preliminary experiments showed there was no cross flow, V*cross, fiom the secondary air inlet into the gasifier, V*=, and no gas carried over fiom the downcomer. The gas cross flow was thus calculated solely on the basis of the flow into the gasifier, V*@i, the flow into the primary air inlet, V*- the volumetric flow of the tracer, V*coz, and the tracer concentration in the gasifier outlet, C&:
195
Figure 6 shows the strong effect of increasing the primary air flow while keeping the total flows into the riser and gasifier constant. Gas cross-flow is also influenced by the flow into the gasifier itself since at higher flowrates the voidage immediately above the solids transfer region is decreased thereby decreasing the resistance to cross-flow. Figure 7 illustrates this but also shows a decreasing fraction of tracer in the gasifier outflow. We conclude that with proper division of flow between primary and secondary air to the riser and total flow into the gasifier the air content of the product gas can be maintained at around the 3% level.
a 0
0
10
20
30 40 50
60
flowrate primary inlet v",
70 80
90100
[vmin]
Figure 6 Gas cross-flow into gasifier as fimction of primary air flowate
CONCLUSIONS
Measurements of the solids circulation rate in the scaled biomass gasifier have revealed a novel scaling parameter, the dimensionless mass turnover M, which is strongly related to the dimensionless superficial gas velacity in the riser part of the circulating system U. M and U are correlated in a second order algebraic equation (equation 8) with a single parameter that remains constant for particles of different density but with the same particle size distribution. M would appear to be a more generally applicable 196
scaling parameter than the dimensionless solid circulation rate suggested by Glicksman et al. (3) but its verification remains to be tested in the pilot plant currently under construction results from which will be reported in due course. The cross flow of gas between the riser-combustor and the bubbling-bed gasifier is strongly influenced by the flow through the primary air inlet but with appropriate division of gas flow between primary and secondary inlets and the total flow into the gasifier the air content of the product gas can be controlled at an acceptable level. The results of these and other tests with the cold model will now be compared with measurements that are shortly to be made on the hlly operational pilot plant so as to test the validity of the scaling relationships that formed the basis of the present work.
flowrate gasifier VSasi [Vmn] Figure 7 Gas leakage into the gasifier as a function of the gasifier flowrate
REFERENCES 1. Foscolo P. U. (1998) Production of Hydrogen-Rich Gas by Biomass Gasification: Application to Small-Scale Fuel Cell Electricity Generation in Rural Areas, EU Final Report, Contract JOR3-CT9.5-003 7 2. Fleck T. (1997) Gaserzeugung aus Biomasse in Schnell Intern Zirkulierender Wirbelschicht - Integration des FICFB Prozesses, Doctoral Thesis, Technical University of Vienna 3. Glicksman L. R, Hyre M. R and Farrell P. A. (1994) Dynamic Similarity in Fluidization, Int JMultuiphase Flow, 20,33 1-386 197
4. Nicastro, M. T. and Glicksman, L. R. (1984) Experimental Verification of Scaling Relationships for Fluidized Beds. Chem Engng Sci, 39,1381-1391 5 . Glicksman L. R, Hyre M. R and Woloshun K. (1993) Simplified Scaling Relationships for Fluidized Beds. Powder TechnoZogv, 77, 177- 199 6. Kunii D. and Levenspiel 0. (1991) Fluidizution Engineering, 2d edn. ButterworthHeinemann, London 7. Horio M. (1997) Hydrodynamics. Circulating Fluidized Beds, Ed: Grace J. R, Avidan A. A. and Knowlton T. M., Blackie, London, 21-85 NOTATION A
a' C
CS
D GS
h M
w m
m* PSD t U
u"
uo u, V*
cross-sectional area constant in equation 9 mole fi-action solid circulation rate diameter solids mass flux height dimensionless mass turnover dimensionless mass load total solids mass load solids mass flow particle size distribution time velocity ratio du, velocity ratio d b superficial velocity terminal fall velocity volumetric flowrate
m2
m kg m*2d' m kg kg s-' S
m s-' m s-' m3 s-'
Greek &b
cb v Pf
Ps
voidage particle sphericity gas viscocity gas density particle density
Pa s kg m-3 kg m-3
Subscripts comb combustor cross cross flow dc downcomer gasi gasifier Prim Primary sec secondary
ACKNOWLEDGEMENT The work was carried out in the flamework of the European Non-Nuclear Energy Programme Joule 111 Project No. JOR3-CT97-0196.
198
Stoichiometric Water Consumption of Steam Gasification by the FICFB-Gasification Process H.Hofbauer, R.Rauch, Institute of Chemical Engineering, Fuel and Environmental Technology Renet Austria Getreidemarkt 9459, A-1060 Vienna
ABSTRACT: The FICFB (Fast Internally Circulating Fluidized Bed) gasification process [ l , 2, 3, 41 is an innovative process to produce a high grade synthesis gas from solid fuels. The basic idea of the FICFB concept is to divide the fluidized bed into two zones, a gasification zone and a combustion zone. Between these two zones a circulation loop of bed material is created but the gases should remain separated. The circulating bed material acts as heat carrier from the combustion to the gasification zone. The design of the gasifier and the use of steam as gasification agent give this process a small heat loss and a nearly nitrogen free product gas with a high calorific value of 13 MJ/Nm3dry gas. This process has been studied over five years and a lot of experiments were carried out in a lOOkWth pilot plant. By using a natural catalyst and gasification temperatures above 800°C the tar content could be reduced below 3 g/Nm3. In former work the general behaviour of this gasification system was studied. In this paper the influence of water on the product gas composition was investigated. As known from literature the steam-fuel ratio has a high influence on the gas composition and the tar content of the product gas. INTRODUCTION In Austria the most common utilisation of biomass for energy is the combustion for heating applications. Gasification could become a second important route especially for power production [5]. Usually, biomass gasification is carried out using fixed or fluidised beds. As the overall gasification reactions are endothermic, the gasification process must be supplied with heat. The easiest way is to use air as gasification agent and to burn biomass partially within the gasification reactor. In this case the product gas has a low calorific value (around 4-6 MJ/Nm3 dry gas) and a high nitrogen content of 45-55%. A gas with a low nitrogen content and a higher calorific value (about 12 MJ/Nm3dry gas) can be produced with pure oxygen as gasification agent but the costs for the oxygen production are high. Another possibility is to supply heat with heat exchangers but here material problems due to the high temperature level will arise. The dilution of the product gas by nitrogen can also be avoided by using a dual fluidised bed system, realised in the FICFB-process. In this case no oxygen generator is necessary and also no serious material problems due to high temperatures will appear. A good overview about such systems is given by Bridgewater [6].
199
CONCEPT BASIC CONCEPT
The basic idea of the FICFB concept is to divide the fluidized bed into two zones, a gasification zone and a combustion zone. Between these two zones a circulation loop of bed material is created but the gases should remain separated. The circulating bed material acts as heat carrier fiom the combustion to the gasification zone (see figure 1). The fuel is fed into the flue product gasification zone and gas gasified with steam. The gas produced in this zone is bed materiai with therefore almost free of thermal energy nitrogen. The bed material, together with some char coal, circulates to the combustion zone. This zone c is fluidized with air and the ._ v b e d material w i t h w is burned. The charcoal 3 ungasified f u e l P exothermic reaction in the combustion zone provides the energy for the endothermic gasification with steam. Therefore the bed material at the exit of the combustion zone has a gasification air agent higher temperature than at the entrance. The flue gas is figure 1: basic concept of the FICFB-gasifier removed without coming in contact with the product gas. With this concept it is possible to get a high-grade product gas without the use of pure oxygen. The FICFB-process has the following advantages against other existing dual fluidized bed systems: 0 simple reactor design 0 low investment cost because of compact design 0 reduced energy losses because of efficient thermal household 0 stable operation conditions, because circulation rate depends only on velocity in the riser
0 L!
The gasifier produces a product gas without nitrogen, therefore the gas can be used in the following applications: 0 combustion of the gas in an engine or turbine to produce electric power 0 use of the gas in a fuel cell to produce electric power 0 the production of synthetic natural gas 0 the production of methanol 0 use of the gas for direct reduction of iron ore in the steel industry 0 use of the gas in the synthetic chemical industry
200
PILOT PLANT
The experience with a cold flow model and the 10 kWth test rig [7] was used to design the 100 kWth pilot plant. The FICFB-gasification process consists of two zones. The gasification zone is fluidised with steam and the combustion zone (riser) is fluidised with air. To avoid large amounts of gas mixing a siphon was introduced in the line from the combustion zone to the gasification zone. The bed material is splitted from the riser gas stream using a separator. The product gas and the flue gas have separated exits out of the reactor. The fuel feeding system consists of two screws. The first is controlled by a frequency converter to adjust the amount of fuel, the second one is a fast rotating screw direct into the fluidised bed. Air is supplied by blowers into the riser and during the start up period also into the gasification zone. Steam is produced by an electrical steam generator and superheated by an electrical heater. The product gas is cooled by a three step heat exchanger and is afterwards cleaned from dust and partly from tar by a bag filter or a sand bed filter. Additives can be fed into the product gas stream in front of the bag filter, to improve the dust and tar separation efficiency. After the particulates are removed the tar is separated by a scrubber. In the scrubber the gas is also cooled down to a temperature of 25-55°C. After these cleaning steps the product gas is burned in a cyclone burner (see figure 2).
figure 2: Flow sheet of the FICFB-process Table 1 contains characteristic data and dimensions of the pilot plant. The reactor is manufactured with stainless steel and is insulated. The warm up is carried out with electrical preheating of all air streams and by combustion in both zones. The whole warm up lasts about 4 hours. An oil feeder was installed into the riser, which gives the possibility to change the temperature level of the system without varying other operation parameters. With this installation parameter studies can be carried out very easily.
20 I
Table 1: Characteristic data of the pilot plant
thermal output fuel reactor diameter riser diameter riser height bed material bed mass mean diameter
100 kW wood pellets, wood chips, biomass residues 300 mm 100 mm 4250 mm quartz, natural catalyst, catalyst 70 kg 0.5 mm
Temperatures, pressures and CO, C02 and H2 content of the product gas are measured and recorded continuously. Gas samples are taken and analysed by gas chromatography to measure the concentrations of H2. CO, C02, hydrocarbons, N2 and O2 in the product gas. CO, C02, NO, NO2 and O2 from the flue gas are measured and recorded continuously. Dust, tar, H2S and NH3 content of the product gas is measured in front and after the heat exchanger, after the filter and after the scrubber. So all separation efficiencies can be determined. Particulates and tar are measured with a method similar to the tar protocol, which was developed by the IEA-gasification group. Gas is sampled isokinetically and filtered in a glass fibre filter at 150°C to remove particulates and heavy tars. Afterwards the gas is cooled in gas washing bottles, filled with toluene at O"C, to separate the light tars. The glass fibre filter is extracted with dichlormethane in a Soxhlet-Extractor, and then the burnable particles are combusted. So the amount of heavy tars, fly coke and unburnable part can be determined. The condensables of the gas washing bottles are extracted, the amount of water is measured to calculate the steam content of the product gas and the amount of light tars is determined gravimetically. For this paper the tar content was measured after the heat exchanger and presents the s u m of heavy and light tars. RESULTS EXPERIMENTAL RESULTS
In 1999 a new gasifier was installed at the Institute. This gasifier has the same basic design as mentioned above, but an improved operation performance as the previous one. The experiments, shown in this paper, were carried out in the old pilot plant as well as the new one. As fuel for the experiments wood pellets were used. These have the advantage, that the water contents is very low (40%) and it is a standardised fuel with nearly the same composition over the whole year. The steam-fuel ratio was varied from 0.18-0.8 kg steam per kg dry fuel. The steam-fuel ratio is calculated in the following way: steam -fuel ratio =
sum water input [kg/h] dry fuel input [kglh]
Sum water input is the sum of steam, which is used in the gasifier (for fluidisation of the gasification zone and siphon) and the water content of the fuel. As bed material a natural catalyst was used and the gasification temperature during the experiments was varied from 750 to 900°C.
202
Gas Composition
As known from literature, the gas composition depends mainly on the used fuel, on the temperature and on the steam-fuel ratio [4,8,9]. The gas composition depends also on the residence time, but in all experiments the residence time was kept as constant as possible. Therefore in the following diagrams the dependency of the dry product gas composition to these parameters is shown. The nitrogen content for all experiments was below 5 vol% and is not shown in the diagrams. The rest to 100 % is nitrogen and higher hydrocarbons. From the gaschromatographic analysis it can be seen, that the main component of the these hydrocarbons is ethene. product gas composition dependency of temperature
-
4
45 40 -
3530 0 ‘= 25 ! 20 0
>
*
**
-
.
-2
A A
-
M
*
:
I
A A
*
I
.
*
*
.*
.
’
HZ
co
A.
4 r A
A
8
E 10 15-
coz x *
x
x
*x
..x
1
CH4
X
501
760
780
800
820
840
860
880
900
920
temperature [“C] From the diagram “product gas composition dependency of temperature” it can be seen, that with increasing temperature the hydrogen and carbon monoxide concentrations are increasing and the carbon dioxide and methane concentrations are decreasing with increasing temperature. The reasons for these dependencies are, that the reactions at higher temperatures are faster and the gas composition is nearer to equilibrium. In the next two diagrams the gas composition in dependency on the steam-fuel ratio is shown. For temperatures between 800 and 850°C the hydrogen and carbon dioxide contents are increasing with a higher steam-fuel ratio. Carbon monoxide is decreasing and methane is almost constant. In the temperature area between 850 and 900°C the same tendencies can be seen, the only difference is, that the gradient of hydrogen is higher and methane is decreasing with higher steam-fuel ratio. With the results of these measurements the gas composition of the product gas can be calculated for different steam-fuel ratios and temperatures. The next step will be to improve the model of the gasifier on basis of this measurements.
203
product gas composition dependency on steam-fuel ratio (800-850°C)
50 45
-
-40 g 35 -
-g
30 .c .-
H2 b
1
25
-
E20
t?
i15El0
f
A
-
-
co
8
-I
c02
x
x
M
CH4
5 -
01 0,OO
0,lO
0,20
0,40 0,50 steam-fuel ratio [kglkg]
0,30
0,60
0,70
0,80
product gas composition dependency on steam-fuel ratio (850-900°C)
50 1 H2
45
-
g 40
g 35 5 30 .c .g 25 I
E20 8
15 El0 u)
5
0 0,OO
0,lO
0,20
0,30 0,40 0,50 steam-fuel ratio [kglkg]
0,60
0,70
0,80
Tar Content The values shown in this paper are only fiom the new gasifier, because the tar measurement method was changed, and the results fiom the old gasifier cannot be compared with the results of the new gasifier. As trend line in the diagrams the same exponential trend was used as it was found in the results of the old gasifier. As known fiom previous experiments the tar content depends strongly on the gasification temperature. Here the dependency of the tar content on the steam-fuel ratio was studied, In the first diagram the dependency of the tar content on the steam-
204
fuel ratio is shown. It can be seen, that with increasing steam-fuel ratio the tar content is decreasing. tar content dependent on steam-fuel ratio (850°C)
.
8 7 r 6 E
t
0) c
5-
Y
E4-
c
i 3 L
5 2 1 -
0 0,oo
0,20
0,lO
0,40
0,30
0,50
0,60
steam-fuel ratio [kglkg]
In the next diagram the tar dependency on the temperature is shown at different steam-fuel ratios. For this diagram two steam-fuel ratios were used to show the dependency of the tar content on the temperature. tar content depending on temperature for different steam-fuel ratios
10
9-
al
2 48 3F? 2 1 -
01 800
820
840
860
880
900
temperature [“C]
It can be seen from the diagram, that the dependency on the temperature is higher at a low steam-fuel ratio. It is also shown in this diagram, that a higher steam-fuel ratio causes a lower tar content in the product gas. From these results it was realised, that the optimal steam-fuel ratio for a low tar content is higher than 0.5 kg steam per kg dry fuel. The optimal temperature for a low tar content is higher than 850°C.
205
Water Consumption In the next diagram the dependency of the steam content in the product gas on the steam-fuel ratio is shown. It can be seen, that the steam content in the product gas increases with increasing steam-fuel ratio. It can be also seen, that there is no linear dependency. At a steam-fuel ratio of about 0.1 kgkg all steam would be used by the gasification reactions. This steam-fuel ratio could not be investigated by experiments, because the lowest steam-fuel ratio, which is possible at the lOOkW,,, gasifier is 0.15 kgkg. steam content in product gas depending on steam-fuel ratio
40
-
g 35
-
0
b 30 CI
C
3 25 -
5 20-
5
1510-
5
0 0,OO
0,lO
0,20 0,30
0,40
0,50
0,60 0,70
0,80 0,90
steam-fuel ratio [kgkg]
In the next diagram the water conversion is shown in dependency on the steam-fuel ratio. The water conversion is an important figure to estimate the efficiency of the steam gasification system. The water conversion was calculated in the following way: water conversion =
sum water input [kg / h ] - sum water output [kg / h] *loo [Yo] sum water input [kg I h]
It can be seen, that the water conversion decreases with increasing steam-fuel ratio. Also the influence of the temperature on the water conversion was investigated. As estimated before, the water conversion increases with increasing temperature. From these results it can be estimated, that the optimal steam-fuel ratio is below 0.5 kg steam per kg dry fuel. Above this ratio the water conversion is almost constant and more steam causes only more steam content in the product gas, without shifting the reactions to the desired side.
206
water conversion dependent on steam-fuel ratio
35
- 3025 c
.-0 f 20 >
5 15 L
9 m 10 3
l 0 0,OO
l
0,lO
0,20
0,30
0,40
0,50
0,60
0,70
0,80
0,90
steam-fuel ratio [kglkg]
CONCLUSION
It was shown in this work, that the product gas composition, the tar content and the water conversion depends strongly on the steam-fuel ratio. The dependency of the product gas composition on the temperature and on the steam-fuel ratio was investigated. With this results the model of the gasifier will be improved and a basis for scaling up was created. Also the dependency of the tar content on the temperature and the steam-fuel ratio was investigated. By using a low steam-fuel ratio a high tar content was produced. Increasing the steam-fuel ratio the tar content could be reduced essentially. As known from previous experiments a high gasification temperature causes a low tar content in the product gas. The influence of the steam-fuel ratio and temperature on the water conversion were also investigated in the experiments. It was shown that a steam-fuel ratio above 0.5 kgkg has only minor influence on the water conversion. Above this steam-fuel ratio only the steam content in the product gas increases, without a major influence on the equilibrium on the gasification reactions. In the 100kWmpilot plant two steam-fuel ratios for all further experiments could be defined. One is 0.25 kgkg, which causes a high tar content, but also a low steam content in the product gas, the second is 0.5 kgkg, which causes a lower tar content, but also a higher steam content in the product gas than the first one. With this two steam-fuel ratios all further experiments will be done and the separation efficiencies of the gas treatment system will be investigated. In this work all necessary investigations referring the temperature and steam-fuel ratio were done to improve the model of the FICFB-gasifier and to have a basis for the scaling up the pilot plant to demonstration plant. REFERENCES [ 11 Hofbauer, H.; Stoiber, H.; Veronik, G.; (1995). "Gasification of Organic Material in a Novel Fluidization Bed System", Proc. Of the Is' SCEJ Symposion on Fludization, Tokyo, pp. 29 1-299
207
[2] Fleck, T.; Hofbauer, H.; Rauch, R.; Veronik, G.; (1996). “The FICFB Gasifcation Process”, Proc. Of the IEA Bioenergy Meeting Banff, Canada May 1996 [3] Zschetzsche, A.; Hofbauer.; Schmidt, A.; (1994). “Biomass Gasification in an Internally Circulating Fluidized Bed“. Proc.of the 8‘h European Conference on Biomass for Agriculture and Industry, Vol. 3, pp. 1771-1777 [4] Fercher, E.; Hofbauer, H.; Fleck, T.; Rauch, R.; Veronik, G.; “Two Years Experience with the FICFB-Gasification Process “ 1Oth European Conference and Technology Exhibition, Worzburg (Jun e 1998) [5] Siplti, K.; (1995). “Research into Thermochemical Conversion of Biomass into Fuels, Chemicals and Fibres “. In Biomass for Energy, Environment, Agriculture and Industry, Proc. of the 8Ih EC Biomass Conference, ed. Chartier, Ph. et al., Pergamon Press, New York, Vol. 1, pp. 156-167 [6] Bridgewater, A. V.; (1995). “The Technical and Economic Feasibility of Biomass GasiJication for Power Generation“. Fuel, Vol. 74, No 5, pp. 631-653 [7] Hofbauer, H.; (1982) “Untersuchungen an einer zirkulierenden Wirbelschicht mil Zentralrohr “. Chem.-1ng.-Tech. 54, Nr. 5, pp. 528-529 [8] Gil, 3.; Aznar, M.P.; Caballero, M.A.; Frances, E.; Corella, J.; (1997) “Biomass Gasification in Fluidised Bed at Pilot Scale with Steam-Oxygen Mixtures. Product Distribution for Very Different Operating Conditions “. Energy & Fuels, Volume 1 1, Number 6; NovemberDecember 1997 [9] Gil, J.; Corella, J.; h a r , M.P.; Caballero, E.; (1999) “Biomass Gasijkation in atmospheric and bubblingjluidised Bed: Effect of the type of gasifiing agent on the product distribution “. Biomass & Bioenergy 1999
208
A Pilot Scale Circulating Fluidized Bed Plant for OrujiZZo Gasification P. Garcia-Ibafiez, A. Cabanillas, P.L. Garcia-Ybarra CIEMAT, Avda. Complutense 22. 28040 Madrid, Spain
ABSTRACT: A pilot plant for air-blown biomass gasification in an atmospheric circulating fluidized bed (CFB) has been constructed at CIEMAT. The plant with a capacity of about 0.3 MW,,, fuel input will be used to carry out gasification experiments using orujillo as a fuel. Orujillo is a by-product (residue) from the olive oil processing industry, which produces about 2 million t/yr. of orujillo in Spain. This paper provides a technical description of the CFB gasification plant and its main characteristics as well as the chemical analysis of these biofuels, which have been analysed by the standard analytical methods. Orujillo analysis results suggest that this kind of biomass will be among the more difficult biomass feedstocks. Like the straw and other agrobiofuels containing a high amount of chlorine and alkali metals, vapour condensation in the gasification plant cold side could lead to serious corrosion problems. Furthermore, this fact together with a high silica content is known to create ash sintering and deposition problems in fluidised beds. The final aim of this project is to demonstrate the feasibility of the orujillo gasification technology for power production. So, in a later stage, the process gas will be used to feed an internal combustion gas engine.
INTRODUCTION Biomass fuels for heat and power generation are of interest because biomass can be considered an environmentally and climatically sound option due to its almost closed carbon dioxide cycle, low sulphur dioxide emissions, good opportunities of ash utilisation and potential to save fossil energy resources. Though there is a considerable potential of biomass for energy in the EU of about 6 EJ/yr, currently about 27% are used [l]. Biomass energy accounts for about 15% of world energy supply; in some developing countries, however, it can represent more than 90% of total national energy supply P I . Starting from the first oil crisis in 1973, researches on biomass energy have been undertaken in many countries around the World and different technologies have been developed for the use of biomass for energy production. Apart from biological conversion processes, several thermochemical conversion routes are possible, depending on the specific application. Combustion, pyrolysis and gasification are the most well known technologies, being combustion the most developed alternative at the industrial level due to its reliability and simplicity of operation. Besides combustion, gasification is a very promising option for further use of biomass especially for electricity production by allowing to substitute steam cycles by more efficient cycles 209
involving internal combustion engines, for instance. Especially the gasification route is thought to have a large potential, due to relatively high overall system efficiencies. However the process and system technology is currently still at a development stage and further development is needed [3][4]. In gasification processes, biomass and other waste fuels are gasified to obtain so called ‘producer gas’ for power or electricity generation. Circulating and bubbling fluidized bed systems have been used extensively for thermochemical processing of biomass during the last decade [5-211. Circulating fluidised bed are generally considered for larger scale systems (> 10 MWh). Although operation of this type of gasifier is more complex than that of a fixed bed gasifier, interesting features are:
(1) Much greater fuel flexibility, i.e. a much greater tolerance towards feedstock particle size, mechanical strength and inorganic composition. (2) Good temperature control and high reaction rates. (3) High carbon conversion. Among the various types of biomass fuels that can be used for energy production agro-residues resulting as by-products of agricultural or agro-industrial activities are thought to be the most important. However, most of the agricultural residues have high ash contents whose behaviour determines the use of particular conversion technology and its efficiency. These biofuels can cause serious ash-related operational problems due to relatively large quantities of some inorganic elements in the biomass ash. The majority of these problems are associated with the low melting point ash of these biofuels causing ash agglomeration and sintering phenomena [ 19-22]. One type of agricultural residue, which is in abundance in rural Mediterranean regions, will be used in this project: orujillo, which is a by-product from olive oil production industry. In Spain, as by-product (residue) from olive oil production about 2 million t/yr of orujillo is produced. The fluidized bed technology is found to be the most suitable for converting a wide range of agricultural residues into energy, due to its inherent advantages of fuel flexibility, low operating temperature and isothermal operating condition [ 191[20]. A pilot plant for air-blown biomass gasification in an atmospheric circulating fluidized bed (CFB) has been constructed at CIEMAT. The plant with a capacity of about 0.3 MWh fuel input will be used to carry out gasification experiments using orujiffoas a fuel. For comparison purposes, two different solid fuels will be used as the feedstocks in the gasification tests of this project: orujillo and wood waste. The project goal is to optimise the operation of the process and interpret data, which will be used for the construction of a Process Development Unit-scale CFBG test facility with gas engine application.
OBJECTIVES The work presented in this paper is focused in the following objectives: (1) the characterisationof the orujillo and other fuels being to be used.
(2) technical description of CFB gasification unit.
210
RAW MATERIAL ANALYSES Physical and chemical properties were determined using standard procedures and are to be described below. The feedstocks of this project will mainly comprise orujillo as well as some woody biomass fuel. The fuels for the analyses were chosen to represent typically two significant biomass sources:
(1) Spanish oak. Wood chips prepared from small wood, branches and bark, without included green parts (needles). (2) Olive oil production residue, orujillo. Produced and collected in Andalucia, Spain. The wood-derived biomass was delivered dry in batches of about 2 m3 in chipped and dried to < 10% moisture content. Two orujillo samples were delivered, after they had been dried up to about 10% moisture content. One of them is from the direct olive oil production process (orujillo#l) and the other is washed (orujillo#2) in order to diminish the potassium content as a mean to prevent bed-agglomeration. This potassium-reduction pretreatment implies a fuel inorganic portion decrease so that a rise in the fuel heating value is caused.
FEEDSTOCK PHYSICAL CHAM CTERIZATION The biomass physical properties were measured without a additional fuel treatment.
Particle size distribution The particle size and moisture content were the two primary properties requiring a lot of pre-treatment of the samples to fit into the desired limits mentioned above. The particle size distribution was determined by normal sieving procedure. The particle size distributions of the fuels are shown in figure 1. ~~
~
Cumulative weight fraction, % 100
80 60
40 20
0 0 -orujillo#l
1
2 3 4 Particle size, mm
- - *-- *orujillo#2 -oak
I
Fig. 1 Particle size distribution of fuel samples.
21 1
5
6 wood
The mean particle size D50, described by the mesh of the sieve dividing the fuel sample in two equal parts, were for the feedstocks as follows:
D50 = 0.33 mm. D5o = 2.22 mm. DSo = 1.58 ~ITI.
Orujillodl: Orujillo#2: Oak wood chips:
Density of dust samples The density of the fuel samples was determined using the approach of the bulk density which corresponds to the overall volume occupied by a given mass of powder, including the pore and interstitial volumes. A one-litre volume was filled up with a measured amount of sample without tamping. Bulk density was determined according to ASTM E873-82. The determined bulk densities are shown in figure 2.
1
Bulk density, kg/m3 1000
800
600 400 200
0 Oak wood
Orujillo#2
OrujiIlo#l Bas recieved .dry
basis
Fig. 2 Bulk densities of fuel samples.
FEEDSTOCK CHEMICAL CHAM CTERISATION Representative samples were taken from each fuel. Oak wood sample was ground with a cutting mill and then with an additional grinding (disc mill). Both of the orujillo samples were ground with a disc mill. To make a comparison possible regarding chemical analysis, all biomasses were ground with the same equipment and the moisture content of the fuels was also relatively homogeneous, ranging from 5 to 15 %.
Proximate and ultimate analyses The ash and volatile matter contents of the dry samples were analysed according to ASTM D1102-84 and ASTM E872-82, respectively. The moisture content was determined according to ASTM D20 16-65. The C, H, N contents of the samples were determined with the Carlo Erba CHNanalyser and the sulphur content with the Leco Sulphur analyser SC.
212
The lower and higher heating values were determined according to ASTM D201596. Proximate and ultimate compositions and heating values are given on dry basis. The proximate and ultimate analyses for the fuel samples are presented in table 1.
Trace compounds The concentrations of alkali metals and chlorine were determined directly for the fuel samples using a inductively coupled plasma-atomic emission spectrometry apparatus in order to analyse sodium and potassium contents. Chlorine content was determined by Eschka method according to ASTM D2361-66. The concentrations are given on dry basis (table 1). Table 1 Feedstock analyses (dry basis).
Analysis
Oak Wood
Orujillo#l
7.9
13.9
11.7
79.2 17.4 3.4
64.6 19.2 16.2
74.4 17.1 8.5
51.0 6.2 0.4 0.05 0.016 42.33
47.3 6.0 1.9 0.09 0.36 44.35
52.7 7.1 1.6 0.07 0.37 38.16
Heating Value HHV (MJkg) LHV (MJkg)
19.1 17.8
18.6 17.4
20.6 19.2
Trace elements, ppm Na K
20 850
690 34200
1750 23500
Orujillo#2 ~
Moisture content (wt%) Proximate (wt %) Volatile matter Fixed Carbon Ash Ultimate ( d a j wt %) C
H N S
c1 0 (diff.)
Ash analysis
The elemental analysis of ash was carried out by ICP technique (inductively coupled plasma spectrometry). The biomass samples were ashed at 550°C to minimise losses. Typical ash analyses are shown in table 2.
213
Table 2 Ash composition of the feedstocks (wt 9%). Oak wood
Orujillo#l
Orujillo#2
Na20
0.45
0.47
2.08
K20
4.54
20.06
24.73
CaO
27.60
6.49
8.01
MgO
2.55
6.21
12.86
Si02
30.40
33.10
16.72
P205
3.40
3.70
1.45
A1203
2.44
4.29
0.89
Fe203
2.10
2.32
2.37
BED MATERIAL PROPERTIES Silica sand is to be used as main bed material in order to form an inert bed in the tests. The main characteristics and chemical composition of the silica sand are given in the table 3.
Table 3 Main characteristics of the silica sand. ProDertv
Value
Physical properties Particle density (kg/m3) 2600 Bulk density (kg/m3) 1500 Maximum particle size (p) 850 Average particle size (pm) 500 Minimum particle size (pm) 200 Chemical composition (wt%) Si02 99.50 A1203 0.25 Fez03 0.06 K 2 0 + NazO 0.10 CaO + MgO 0.02
EXPERIMENTAL FACILITY DESCRIPTION The CFB gasifier test facility is an atmospheric air-blown installation and consists of four main parts: (1) Feeding system; (2) Primary and secondary air supply; (3) Gasifier vessel equipped with a cyclone and a particle recirculation system; (4) Gas analysis equipment.
214
A schematic diagram of this pilot plant, which has a nominal capacity of 0.3 MW*, is shown in Figure 3 and technical and operational data are summarised in table 4.
PREHEATER
Fig 3 Schematic diagram of the atmospheric CFB gasification test rig.
The test rig is equipped with a feedstock hopper suitable to low-bulk-density biofuels and the biomass is fed continuously into the sand bed of the reactor via an injector screw which can feed fuel either into the bottom of the bed or to about 1 meter above the air distributor. The maximum fuel mass flow rate is 90 kg/h. The feeding system consists of two screw feeders in series separated by a rotary valve. The second screw has a higher feeding rate thus it will remain almost empty and therefore is not likely to be blocked by pyrolysis products. Steam can be added to the primary air as a gasification agent. The gasification air is fed to the bottom of the reactor via an air distribution plate. Primary air can be preheated to 500°C in the start up. Also secondary air can be fed to the reactor on different levels in the freeboard section. When the gasification air enters into the gasifier below the solid bed, the gas velocity is high enough to fluidise the particles in the bed. At this stage, the bed expands and all the particles are in rapid movement. The gas velocity is so high, that a lot of particles are conveyed out from the reactor into the cyclone. The heart of this facility is a refractory-lined reactor cylinder with 20 cm indiameter and a total height of about 6.5 m. The internal refractory forms two concentric cylinders (each one of thickness 6.5 cm) with different thermal conductivity and wearing properties. Additionally, the reactor tube, the cyclone and the recycling line are protected with outside insulation in order to minimise heat losses. The gasifier components (raiser, cyclone and standpipe) are equipped with ten thermocouples (Ktype) and pressure transducers. The thermocouples probes are connected to a computer based data acquisition and control system. In addition, the feed rates of gasification 215
agent are continuously measured and controlled by flowmeters and electronic pressure transducers are used for process evaluation. Table 4 Technical data on CFB gasification test rig of CIEMAT.
Inside diameter Reactor height Operation temperature Fluidisation velocity Operation pressure Fuel feed rate (max.) Gasification agent MaximumThermal capacity
0.20 m 6.5 m 600- 1000°C 2-6 m/s 1.O bar (abs) 90 kgh air about 300 kW
The operating temperature in the reactor is typically 600-1000°C depending on the fuel and the application. When entering the reactor, the biofuel particles start to dry rapidly and the first primary stage of reaction, namely, pyrolysis occurs. During this reaction fuel converts to gases, char coal and tars. Part of the char coal goes to the bottom of the bed and is oxidised to CO and C02 generating heat. After this, as these aforementioned products flow upwards in the reactor, the secondary stage of reactions take place, which can be divided into heterogeneous reactions, where char is one ingredient in the reactions, and homogeneous reactions where all the reacting components are in the gas phase. Due to these reactions, a combustible gas is produced, which enters the cyclone and escapes the system together with some of fine dust. Most of the solids in the system are separated in the cyclone and returned to the lower part of the gasifier reactor. These solids contain char, which is combusted with the air that i's introduced through the distribution plate to fluidise the bed. This combustion process generates the heat required for the pyrolysis process and the subsequent mostly endothermic reactions. The circulating bed material serves as heat carrier and stabilises the temperature in the process. Due to this circulation high temperatures can be maintained in the whole reactor with intensive gas-solids contact. In normal operation, the fuel feed rate will define the capacity of the gasifier and the air feed rate will control the temperature in the gasifier. The coarse ash is accumulating in the gasifier and will be removed from the bottom of the gasifier. Figure 4 shows the gasifier installation with the reactor in the middle, the feedstock hopper in the background right-hand side and the recycling line in the foreground lefthand side. After leaving the gasifier, the producer gas enters into a combustion chamber where propane gas can be injected to support combustion. The sampling point for the product gas analyses is located downstream the cyclone. The temperature of the whole sampling system is kept at about 250°C to avoid tar condensation. The main gas constituents, CO, C02, H2, C&, and the contents of hydrocarbon components ((22,) are to be analysed by on-line gas chromatography (GC). To this end, the chromatograph is equipped with both a thermal conductivity detector and a flame ionization detector. The lower heating value (LHV) of the product gas is expected to be in the range 3.5-5.0 MJ/Nm3.
216
For all tests the experimental measurements will be the following: (1) Temperature measurements along the length of the reactor and standpipe as well as at several other important locations such as the cyclone exit. (2) All input air flow rates (primary and secondary air), fuel feed rate and solid discharge rates from the bed. (3) On-line gas analysis in the flue gas for CO,C 0 2 , H2,Cb,C,H,.
Fig 4 Pilot scale gasification installation.
FUTURE WORK AND PERSPECTIVES Biomass CFBG will be studied in a pilot scale plant. Several test runs will be carried out with orujillo and wood waste as a fuel. The main objectives of these tests being to investigate the effects of the equivalence ratio, temperature and fluidisation velocity on the gas production rate from biomass CFBG, and to evaluate the quality of the producer gas as measured by the composition and higher heating value. Process variables to be analyzed are: 217
(1) Equivalence ratio (0.20-0.45), as used in this project, is defined as the air-to-fuel ratio divided by the stoichiometric air-to-fuel ratio; (2) Temperature of the gasifier bed (750-900°C); (3) Fluidisation velocity (2-6 d s ) ; (4) Use of secondary air (about 10%of the overall); ( 5 ) Addition bed additives. Although biomass gasification involves chemical processes that are thermodynamically favourable, a number of practical problems must be solved if a commercial breakthrough is to be achieved. Typical issues that are significant to overall process performance are those of fuel feeding and of gas cleanup (particulate, tar and alkali metal removal). Gas cleaning involves two important issues. The first is the removal of particulate matter from the product gas. The second is the matter of chemical impurities in the gas, such as hydrocarbon fractions of high molecular weight, trace metals (alkali), and nitrogen and sulphur-containing products that have to be removed or reduced to a level that is accepted by the downstream systems (gas turbines, engines, fuel cells). Tar formed in gasification is a complex mixture of organic compounds ranging from low molecular weight compounds like benzene to heavy polyaromatic hydrocarbons. Tar components can cause condensation and plugging problems in application where gas has to be cooled to below lOO"C, like gasification gas engine applications. Therefore, tar removal will be an objective of this project. Using advanced tar and gas sampling and analysis methods, the gas composition and tar content in the gas will be determined and their variation with the operation parameters will be studied. CFB exhibit very high particle contents and moderate to high tar levels in the producer gas [231. Key points in the present research programme are: optimisation of the gasifier with respect to thermal efficiency and gas quality; determination of operating limits (throughput, temperature, material characteristics); prevention of tar and acids formation. The global objective of the project is to evaluate, to laboratory scale, pilot plant and demonstration plant, the real possibilities (technical and economic viability) of gasifying several types of residual and cropped biomass that are (or could be) produced in our country in order to generate power or electricity by internal combustion in gas engines.
REFERENCES 1. Kaltschmitt M. & Dinkelbach L. (1997) Biomass for Energy in Europe. In Biomass Gasification and Pyrolysis (Ed. by M . Kaltschmit & A.V. Bridgwater), pp. 7-23.
CPL Press, UK. 2. Bain R. L., Overend R. P. & Craig K. R. (1996) Biomass-fired Power Generation. In the Engineering Foundation's Biomass Usage for Utility and Zndustrial Power Conference, Snowbird, UT. 3. Kaltschmitt M. & Bridgwater A.V. (1997) Research, Development and Demonstration Needs for Biomass Gasification and Pyrolysis. In Biomass Gasifcution and Pyrolysis (Ed. by M . Kaltschmit & A.V. Bridgwater), pp. 537550. CPL Press, UK.
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4. Bridgwater A.V. (1995) The Technical An Economic Feasibility of Biomass Gasification for Power Generation. Fuel, 63 1-653. 5. Narviiez I., Orio A., Aznar M.P. & Corella J. (1996) Biomass Gasification with Air in an Atmospheric Bubbling Fluidized Bed. Effect of Six Operational Variables on the Quality of the Produced Raw Gas. Ind. Eng. Chem. Res., 35 (7), 21 10-2120. 6. Maniatis K., Vassilatos V. & Kyritsis, S. (1994) Design of a Pilot Plant Fluidized bed Gasifier. In Advances in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater), vol. 1, pp. 403-410. Blackie Academic Proffesional, Glassgow, UK 7. Kurkela E. & Stahlberg P. (1992) Air Gasification of Peat, Wood and Brown Coal in Pressurised Fluidized Bed Reactor. I. Carbon Conversion, Gas Yield, and Tar Formation. Fuel Process. Technol, 3 1, 1-21. 8. Kinoshita C., Wang Y. & Zhou, J. (1994) Tar formation Under Different Biomass Gasification Conditions. J. Anal. Appl. Pyrolysis, 29, 160-18 1. 9. Schenk E.P., van Doom J. & Kiel J.H.A. (1997) Biomass Gasification Research in Fixed Bed and Fludised Bed Reactors. In Biomass Gasification and Pyrolysis, (Ed. by M. Kaltschmit & A.V. Bridgwater), pp. 129-138. CPL Press, UK. 10. Ising M., Holder D., Backhaus C. & Althaus W. (1998) Fluidized Bed Gasification of Biomass for Cogeneration. In Biomass for Energy and Industry, (International Conference, June 1998, Wiirzburg, Germany), (Ed. by H. Kopetz, T. Weber, W. Palz,P. Chartier, G.L. Ferrero), pp. 1549-1551.C.A.R.M.E.N.,Rimpar, Germany. 11. Gil J., Aznar M.P., Caballero M.A., Francis E. & Corella J. (1997) Biomass Gasification in Fluidized Bed at Pilot Scale with Steam-Oxygen Mixture. Product Distribution for very different operating conditions. Energy & Fuels, 11 (6), 11091118. 12. Gil J., Corella, J., Aznar M.P. & Caballero M.A. (1999) Biomass Gasification in Atmospheric and Bubbling Fluidized Bed: Effect of the Type of Gasifying Agent on the Product Distribution. Biomass and Bioenergy, 17 ( 5 ) , 389-403. 13. Pinto F., Franco C., Andrt R.N., Gulyurtlu I. & Cabrita, I. (1999) Gasification of waste materials and their co-processing with biomass. 3gh IEA Fluidised Bed Conversion: Fluidised Processing of Unconventional Fuels, (Ed. by A. Cabanillas & M. Miccio), pp. 123-135. Ciemat, Madrid, Spain. 14. Turn S.Q., Kinoshita C.M., Ishimura D.M. & Zhou J. (1998) The fate of inorganic constituents of biomass in fluidized bed gasification. Fuel, 77 (3) 135-146. 15. Garcia L., Salvador M.L., Arauzo J., & Bilbao R. (1999) Catalytic Steam Gasification of Pine Sawdust. Effect of Catalyst WeightBiomass Flow Rate and SteardBiomass Ratios on Gas Production and Composition. Energy & Fuels, 13 (4), 851-859. 16. Pan Y.G., Roca X., Velo E. & Puigjaner L. (1999) Removal of tar by secondary air in fluidised bed gasification of residual biomass and coal. Fuel, 78 (14) 1703-1709. 17. Rensfelt E. & Ekstrom C. (1989) Fuel Gas from Municipal Waste in an Integrated Circulating Fluidized bedGas Cleaning Processes. Energy Biomass Wastes, 12, 8 11-906. 18. Simell P.A. & Bredenberg J.B.-son (1990) Catalytic Purification of Tarry Fuel Gas. Fuel, 69, 1219-1225. 19. Rostn C., Zanzi R. Sjostrom K. & Bjornbom E. (1998) Pressurized gasification of olive waste in a fluidized bed reactor. In Biomass for Energy and Industry, (International Conference, June 1998, Wiirzburg, Germany), (Ed. by H. Kopetz, T. Weber, W. Palz, P. Chartier, G.L. Ferrero), pp. 1665-1668. C.A.R.M.E.N.,
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Rimpar, Germany. 20. Natarajan E., Ohman M., Gabra M., Nordin A., Liliedahl T. & Rao A.N. (1998) Experimental determination of bed agglomeration tendencies of some common agricultural residues in fluidized bed combustion and gasification. Biomass and Bioenergy, 15 (2), 163-169. 21. Ergudenler A. & Ghaly E. (1992) Quality of gas produced from wheat straw in a Dual-Distributor type Fluidised bed Gasifier. Biomass and Bioenergy, 3(6), 419430. 22. Ligthart F.S., van der Drift A. & Olsen A. (1998) Bed-Agglomeration in Fluidized Bed Conversion of Biomass. In Biomass for Energy and Industry, (International Conference, June 1998, Wlirzburg, Germany), (Ed. by H. Kopetz, T. Weber, W. Palz, P. Chartier, G.L. Ferrero), pp. 1765-1768.C.A.R.M.E.N., Rimpar, Germany. 23. Hasler P. & Nussbaumer Th. (1999) Gas cleaning for IC engine applications from fixed bed biomass gasification. Biomass and Bioenergy, 16,385-395.
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A two-stage pyrolysis/gasification process for herbaceous waste biomass from agriculture E. Henrich, E. Dinjus, S. Rumpel, R. Stahl Forschungszentrum Karlsruhe, Institut fur Technische Chemie, CP V, Hermann von Helmholtz Platz I , D- 76344 Eggenstein-Leopoldshafen
ABSTRACT: Cereal straw and strawlike residues from agriculture can contribute a small but significant share of about 5% to the world primary energy supply. Straw is among the cheapest renewables, because it is reliably available as a by-product of food-production without additional cultivation and harvest expenses. But straw is far more difficult to manage technically than wood and therefore hardly used for energy generation. As all fast growing biomass, straw contains more ash, potassium, chlorine and other heteroatoms; the low ash sintering temperature can cause reactor slagging due to balung and stickmg of the reaction material. Gasification technologies are expected to be superior to combustion in view to flexibility, efficiency and environmental compatibility. Wood gasifiers are not suited for straw. Straw gasifiers are not available on the market, but the large amounts of residual straw justify a special development. On the basis of a systematic analysis of various gasifier types and experience on straw gasification reported in the literature, we have selected a two-stage pyrolysis/gasification process for further development, whch especially takes care of the particular straw properties. The first step is an allothermal pyrolysis of straw chops in a fluidised bed of sand, followed by a second autothermal gasification step with oxygen or air at higher temperature. Various chemical and engineering aspects of the process and its many variants are being studied in a number of test facilities. The final aim of these experiments is the collection of design data for a pilot plant.
ENERGY FROM AGRICULTURAL BIOMASS RESIDUES
The consumption of fossil fuels is not sustainable. With business as usual, the proven and economically recoverable fossil reserves will be exhausted in little more than 100 years. Substitution of fossil fuels is therefore required anyway on the long term. If COz emission from fossil fuel combustion is responsible for adverse climate changes like global warming, substitution is desirable on the short term. The share of the renewable energy sources l k e wind, hydropower and biomass may be increased in a relatively short time. Each of these traditionally well known sources can contribute a
22 1
significant part to our primary energy supply. A crude estimate of the energy potential of biomass, especially for the agricultural biomass residuals like straw and strawlike materials is given in the following sections.
PRESENT AND FUTURE ENERGY CONSUMPTION An estimate of the world primary energy consumption can be obtained by
multiplication of the world population with the energy consumption per capita (see Fig. 1). The present world population of 6 billion consumes energy of 12 Gt of carbon equivalent [Energy conversion factors: 1 tC (tonne carbon equivalent) corresponds to o 1.12 tce (coal equivalent) x 0.77 toe (oil equivalent) = 2 t biomass (daf = dry and ash free)] (not coal) per year; this corresponds to an average of 2 tonnes of carbon per capita and year. The per capita consumption in most industrialised countries is usually 2 to 3 times higher than the average. The rapidly growing world population is predicted to arrive at about 9 billion in 2050 and to stabilise at around 10 billion towards the end of the new century. If the average energy consumption per capita and year increases only slightly from 2 to 2.4 t/y carbon equivalent - this is about half the present level in Europe - the world energy consumption will be doubled to 24 Gt carbon equivalent per year.
I0
10
5
5
C
s?i
0
+ 1800
0
I900
I
-1
1850
-
I
27
2050
I 12
I
24
2lM)
year
I 24+ m c i y
Fig.I: Development of the world population and the primary energy consumption
Almost 80% of our present energy mix comes from fossil fuels; the rest is renewable and nuclear energy. In Fig. 2, the biomass contribution is summarised among the solid fuel share, together with coal. Today, biomass still contributes about 10% to the world primary energy supply; commercially traded firewood amounts to almost half of this share. The Kyoto conference agreement in 1997 imposes legally binding obligations on its international signatories to reduce the greenhouse gas emissions. For the European Union (EU) this implies an 8% reduction until 2010 in comparison with 1990 levels. Emission of fossil COz is responsible for about half of the greenhouse effect. If we gradually consume all proven and economically recoverable fossil coal, 222
oil and gas reserves (with ca. 1400 Gt carbon content) as usual, the present C02-level of 367 ppmv in the atmosphere will be about doubled. If we continue to consume also some less economic ultimate reserves, a C02-level of 0.1 %vol. will be quickly approached.
-
12Gtcly
FUEL SHARES 1996:
400
nuclear power hydropower
5 300
5.0 % 6.3 %
gases
21.0 %
liquids
343 %
solid fuels: hardcod, Iignitc
32.2 %
a l
i
-al 200 a '7 100
, . . . . . . . . . . . . . . . . . . . . . . . .-. .K.I.
0 1970
commercial firewood
1975
iSS0
1985
1990
1995
5+
2000 windenergy, geothermal
year
0.1 %
Fig. 2: Breakdown of the world primary energy consumption It has been proposed as a transitional solution, to recover COz from the flue gas of large fossil fired power stations and to dispose it of in the deep sea or empty gas or oil fields [6]. But the development and implementation of such transitional technologies takes much time and money and the real application consumes a considerable fraction of the energy produced. There is some doubt, if such measures are really usefbl. The only sustainable solution to the C02-problem, is the development and use of all reasonable renewable energy sources. The EU for example intends to double the share of renewables from 6% (3% biomass, 90% from forestry (Ph.C.P Chartier et al. in Ref. 1 p. 62)) today to 12% until 2010, as an essential contribution to achieving the Kyoto objectives. The major part of this increase is expected to come from biomass.
ENERGY POTENTIAL OF BIOMASS The global biomass production in the various ecosystems of the earth is summarised in Fig. 3, based on data given in [7]. The total harvestable biomass production amounts to about 65 Gt C-equivalent per year. About 38% (ca. 25 Gt C/y) is produced in the oceans and not considered for energy generation for various reasons. About 62% of all biomass (ca. 40 Gt C/y) is produced on land. Since about 34% of the land area are deserts, the average productivity of the fertile land area is about 1 kg air-dry biomass (0.4 kg carbon equivalent) per m2 land area and year, with a large scatter of factor 3. l%s corresponds to a low photosynthetic efficiency between 0.1 and 1%. Though photovoltaic cells are 10 to 100 times more efficient per m2, harvest,
223
combustion or gasification of biomass is still a more cost-effective power generation possibility. Almost 70% of the land biomass is produced in forests (ca. 27 Gt-C/y). The major part of the production are falling and rotting leaves etc.; the valuable wood and woody parts make up only about a quarter or less. With a world forestry management including afforestation, a sustainable wood harvest of about half of the up-growth can be expected. This corresponds to an about three-fold increase compared to the present production of timber plus commercial firewood (ca. 2 Gt/y).
Fig. 3: Global biomass production in the various ecosystems Agriculture comprises arable land and grasslands with 10% and 15% of the land area respectively; the percentage contribution to the land biomass production corresponds to 11% to each part. (ca. 4.2 Gt/y C-equivalent). The main products are human and animal food. In the course of the next 50 to 100 years, food production has to increase in proportion to the population from 4.2 Gt/y C for 6 billion today to about 7 Gt/y C for an almost 10 billion population. This aim can be achieved by a combination of (1) an increase of the arable land area and (2) a productivity increase with the help of fertilisers etc..
ENERGY POTENTIAL OF CEREAL STRAW One part of the agricultural harvest are crops and the other part are by-products. Byproducts are only partly used and a substantial fraction will be available for energy generation without additional cultivation and harvesting expenses. Specially grown energy crops are more expensive than the agricultural by-products. A rather simplified breakdown of the total agricultural harvest of 4.2 Gt Cequivalent is outlined in Fig. 4. Half of the arable land (total 14 million km2) is cultivated for cereals: wheat, rice, maize and barley make up about 90% of all grain. The known world grain harvest in 1998 amounts to 2.1 Gt. From the average graidstraw ratio of ca. 1, the world straw harvest can be estimated to be about the
224
same. From present practice it is known, that about half of the straw harvest is not required as litter, coarse fodder or humus for soil Improvement. The long stubble and the roots left in the field are usually sufficient for humus formation. In a number of Danish district heating plants, about 20% or 1 Mt of the Danish straw harvest is already used for space heating. An extension to 50% is thought to be possible [M.G.Larsen in Ref. 1, p.11961. Therefore, about 1+ Gt/y of air-dry (15% water) cereal straw and unused residues will be available for energy generation. The energy content (LHV 4 kWh/kg) is equivalent to 0.44 Gt C or about 3.7% of the present world primary energy Consumption. products
by-products
hay-3% 1
for energy
grazing land, haymdring
Fig. 4: Agricultural by-products for energy The harvest on the other half of the arable land comprises a large diversification of different food crops and also some raw materials like cotton. For a crude first estimate, the energy potential of the total biomass harvest as well as the product /byproduct-ratio are assumed to be similar to the gradstraw harvest. Examples for byproducts are leguminous and rapeseed straw, sunflower and cotton stalks, sugar cane bagasse and many others. Many of these agricultural residuals are also suited for energy generation and some of them have a considerable local importance llke sugarcane bagasse in Brazil. Due to the large diversification of these herbaceous biomass by-products, we assume, that an energetic utilisation of only 16% of the total amount is justified for practical reasons. This is equivalent to 0.14 Gt C/y or 1.2% of the present primary energy consumption. A small additional contribution may be expected in form of dry grass, roadside and spoiled hay, which is unsuited as animal food. A small contribution from 3% of the grassland area results in 0.12 Gt/y C (3% from 4.2 Gt C total grass up-growth per year) or 1% of our present primary energy consumption.
225
Table 1: Estimate of the potential energy contribution from agriculturalby-products
World Cereal straw Herbaceous Unsuited hay Total population biomass contribution 6 billion 0.44 0.14 0.12 0.70 Gt C/y this is equivalent to 5.8% from 12 Gt/y primary energy consumption 10 billion 0.74 0.28 0.12 1.14GtC/y this is equivalent to 4.8% from 24 Gt/y primary energy consumption ~
I
I
The crude estimate for the energy potential of all agricultural by-products can be summarised as follows (see also fig. 4):They can contribute a small but sigmficant share of about 5% to the world primary energy supply. Different to the potential biomass energy from forestry, the contribution grows about in proportion to the population. Half of the cereal straw harvest makes up almost two third of this contribution. The potential energy contribution of wood from forestry is expected to be 2 to 3 times higher than the contribution from agriculture. In view to the energy potential of the C02-neutral biofuels, straw ranges in a second position behind wood. The large amount and the uniformity of straw justify a special technology development for an optimum energetic utilisation. SELECTION OF A STRAW GASIFICATION PROCESS
SPECIAL PROPERTIES OF STRA W In the following sections, “straw” is used as a synonym for all herbaceous, thinwalled and rapidly drying biomass with a low water content. Only dry biomass is storable without biological degradation. Compared to fuewood, straw and herbaceous by-products from agriculture have some unpleasant properties. Straw bales have a low storage density of about 0.1 t per m3 (straw chops ca. 0.05 t/m3),transport and storage become more expensive. The composition is a more serious and general aspect. The large organic lignocellulosic CHO-fraction of straw or wood is almost the same. This dry, ash- and heteroatom-free (dahf) biomass part can be represented by a simplified stoichiometric formula C3(HzO)2 (MW 72); formally an equal weight mixture of carbon and water with a slightly higher gross heating value (5 kWh/kg). But all rapidly growing biomass, such as grass, cereals, leaves etc. contains more ashes and heteroatoms such as N, S,C1, alkali and others, than wood. Ash constituents and volatile heteroatoms are needed in larger concentrations for a fast growth. Especially the ash-, K-and C1-contents in herbaceous biomass can be higher up to one order of magnitude than in wood without bark. A typical composition of wood and straw is compared in Table 2 [9,101.
226
Table 2: Typical percentage composition of straw and hay compared to willow wood. Dry biomass heating Val. CHO-fraction Heteroatoms MJkg C H 0 N S C1 Willow wood 18.4 49.8 6.5 41.8 0.3 0.03 0.01 dahP* (calculated) 50.8 6.6 42.6 Wheat straw 17.5 46.7 6.3 40.7 0.5 0.1 0.4 dahf (calculated) 49.6 6.8 43.6 Hay 16.445.0 6.2 39.5 1 0.08 0.26 dahf (calculated) 49.6 6.8 43.6 C3(H20)2 -18.2 50 5.6 44.4 *< 1% without bark, **dahf: dry, ash- and heteroatom-fie:e ~
Ash constituents K20 P2OS CaO ash
0.3 0.2
0.7 1.6*
1.5 0.3
0.4 5.3
4
0.8 8.0
~
1
~~~
In particular, technical difficulties of straw combustion or gasification are caused by the higher ash and heteroatom contents, especially chlorine and potassium: 0
0
The high HCl- concentration in the fuel gas is the cause of (1) corrosion problems, (2) catalyst poisoning in the gas cleaning system and (3) PCDDFformation on the surface of entrained dust particles during fuel gas combustion. K lowers the ash sintering temperature, eventually down to 700°C. T h ~ screates the risk of reactor slagging by sticky and glueing ash particles. Volatilisation of alkali compounds like KOH or KCl at temperatures above 600 to 700°C causes deposition, plugging and corrosion problems by desublimation during cooling down in the piping downstream.
Thermal treatment technologies for wood are usually not designed to control these special effects.
GASIFICATION VERSUS COMBUSTION A thermochemical conversion of a solid biofuel into a fuel gas is an additional operation prior to application, which can be justified by a more efficient and flexible use of the cleaner gas. The pollution and flue gas cleaning problems from direct combustion of dirty coal are expected to be reduced by more efficient and environmentally compatible integrated gasification combined cycle (IGCC) processes. This strategy for an advanced utilisation of fossil fuels is also suited for solid biofuels like the relatively clean wood, as it is even more suited for the dirtier cereal straw and related herbaceous bio-materials. Compared to a simple high temperature heat generation via combustion, a combustible fuel gas is suited for many different applications. The required fuel gas cleaning efforts depend on the final utilisation. Moderate purity requirements are connected to a special application such as a: reduction gas e.g. for ore reduction fuel gas for high temperature process heat generation in cement kilns; lime- or brick liunaces or during glass production. Higher fuel gas purities in view to dust, tar, HCI, S02, alkali etc. are needed for
227
0
0
power and electricity generation in turbines or diesel and otto internal combustion engines. The highest fuel gas purities are required for catalysed processes, since many catalysts are sensitive to a number of trace impurities in the gas syngas or hydrogen for the production of methanol, hydrocarbon fuels, ammonia or H2 for electricity generation in low temperature fuel cells (FC) ldce proton exchange membrane fuel cell (PEMFC) at 80°C and phosphoric acid fuel cell (PAFC) at 200°C. High temperature fuel cells like the solid oxide fuel cell (SOFC) and molten carbonate fuel cell (MCFC), operating at 1000°C and 600°C respectively, tolerate CO and hydrocarbon vapours, but are sensitive to various trace impurities like HC1, H2S, alkali etc.
-
-
EVALUATION OF STRA W GASIFICATION TECHNOLOGY
Wood gasifiers are state of the art. In view to a reliable long term operation of turbines or internal combustion engines with a fuel gas, there are still some remaining problems with the high tar content in the gas. Technology for straw gasification is not well developed. Straw gasifiers are not yet available on the market. In gasifiers originally designed for wood or waste, straw has been used occasionally and in most cases in a mixture together with other solid fuels. Furtlier details and references are found in [ 1-5, 18, 191. The suitability of the basic gasifier types for straw gasification has been evaluated in a FZK-study [ 111. The basic gasifier types considered are fixedbed, fluidised-bed and entrained flow gasifiers. Fixed bed gasifiers are simple and easy to operate with larger wood pieces but not with chopped straw. This will block the gas flow. The production of larger straw pellets or briquettes is a available technology, but pelletisation costs of 25 to 50 Euro per ton are too hgh. A sufficiently fast gasification requires temperatures 2 8OOOC and even for straw pellets, the risk of ash sintering and reactor slagging can not'be completely eliminated. The high thermal efficiency of an updraft gasifier (countercurrent flow) is connected with a low fuel gas temperature and a high tar content >10 g per m3. A downdraft gasifier (co-current flow) has a lower tar content < 0.5 g per m3 fuel gas, but is thermally less efficient. Gasification in the flame of an entrained flow gasifier proceeds in I 1s at high temperatures of 2 1300°C. Precondition for an efficient conversion is, that the fuel can be milled to a fine I 0.1 mm powder. This is easy for brittle coal, but complex and expensive for biomass with cellulose fibres, like straw. In the very hot fuel gas the tar content is rather low and a molten slag is obtained. The large amount of sensible heat must be recovered and recycled afterwards to achieve a good cold gas efficiency > 70% (LHV of cold fuel gas/LHV of initial fuel). Cooling-down of the fuel gas may also become necessary for efficient gas cleaning, since efficient hot gas cleaning procedures are not yet available. Gasification in a bubbling or circulating fluidised bed of sand must be conducted at 2 800°C in order to achieve a sufficiently fast gasification. At such temperatures, the low melting potassium salts in the straw ash were found to cause bed agglomeration and breakdown of fluidisation [ 121. None of the common gasifier types is really suited for straw chops. The main reasons are either the expensive straw preparations steps (pelleting, briquetting,
228
milling) or the high risk of reactor slagging, due to the low ash sintering temperature in the presence of much K.
TWO-STEP STRA W GASIFICATION CONCEPT In response to the considerations of the previous section, a FZK study group has selected the two-step straw gasification concept in Fig. 5 as a good basis for the fiuther development of straw gasification [ I l l . Only cheap straw chops are considered as feed, whereas reactor slagging by sintering ash is avoided with a twostage concept. option:
I
I heat input:
autothermal
at %TOO OC
at ca. loo0 O C
cha I
I
I
i
heatlosses
Fig. 5: Two-stage process concept for straw gasification In general, gasification and combustion are always preceded by a pyrolysis step. There is a number of newly developed gasification processes for municipal and hazardous wastes [13, 141 and solid biofbels [15, 161, whch also make use of a separate pyrolysis reactor prior to gasification. The first step of the straw gasification concept, is an allothermal pyrolysis of chopped straw in a fluidised bed of sand at temperatures below the ash softening temperature of about 700°C. Steam can be added to sustain bed fluidisation and to reduce tar formation. The allothermal reaction mode in combination with the excellent heat transfer in a fluidised bed, permits the recycle of heat to improve the process efficiency and simultaneously reduces the 02-consumption during fuel combustion in the following gasification step. Allothermal gasifiers are less common, but have been studied extensively within the framework for the use of nuclear process heat for gasification [ 171. The second step is an autothermal gasification of the pyrolysis gas and the solid pyrolysis char at temperatures above 800°C after partial combustion with air or oxygen. The pyrolysis char carries the whole ash and must be separated from the bed material prior to gasification. Straw char is a rather brittle material and is easily crushed to a fine powder, which gasifies faster. After a relatively fast pyrolysis step the pyrolysis gas contains > 50% of the energy.
229
CHEMICAL AND ENGINEERING ASPECTS MASS AND ENERGY BALANCES Mass and energy balances are obtained from stoichiometric reaction equations. For the preparation of the pyrolysis equation, the product spectrum must be known experimentally. The same is true for gasification, except for very high temperatures, where thermodynamic equilibrium is attained. Pyrolysis in inert atmosphere between 400 and 700°C produces water vapour, COz, combustible gases CO, H2, CH4 and a multitude of organic vapours from the biopolymers cellulose (C6(H20)s), hemicellulose (Cs(H2O)4) and lignin. An impression of the complex product spectrum especially of the condensable organic vapours is given in Fig. 6. The remainder is a black char, mainly consisting of carbon and inorganic ash oxides. CH30H
$pq
HCOOH
H C-0
methanol acetic CH,COOH formk acid acid CH2OHCHO hydmyacetaldehyde
cno guaiacol
Nc
4
HO
Ievoglucosane
levoglucosenone
5-hydroxydmethyl- 1,Sanhydro4deoxy2luraldehyde Dqlycaro-hexcn-2~iose
Fig. 6: Condensable organic CHO-vapours The following examples of pyrolysis equations are oversimplified, but show the essentials. The simplified formula C3(H2O)2 (MW 72) represents the dahfcomposition of the organic part in wood, straw or any other lignocellulosic biomass. CHO-intermediate: C3(H20)2
+
C + HOCH2-CHO
+ +
C + C H 4 + C 0 2 ; AH
=
-146kJ/mol
C + 2 C O + 2 H Z : AH
=
+102kJ/mol
biomass (dahf) char tars, oils Gas production can be exothermal with C 0 2 and CH4 as products or endothermal with CO and H2 as products. Since both reactions proceed simultaneously, pyrolysis can be almost thermoneutral. Most of the energy is then required to heat up the feed material to the pyrolysis temperature.
KINETICS OF STRA W PYROLYSIS AND GASIFICATION Heterogeneous kinetics of straw pyrolysis and straw gasification are essential data for reactor design. Pyrolysis is a relatively fast process. In view of the poor heat conduction of straw and straw char, the pyrolysis time is the time which is required to heat the center of the particle to the decomposition temperature. This is a rather simplified model, but allows a reasonable time estimate in view of the order of magnitude.
230
For an experimental determination of the pyrolysis time, a fluidised bed of sand and a large volume hot vessel have been used [20]. The method for reaction rate measurements in the fluidised bed is partly explained in Fig. 7. The pyrolysis gases in the fluidising nitrogen gas stream are combusted after 02-addition downstream from the vessel and the C0 2 plus some CO are monitored with NDIR-analysers. Thus, the course of a slow pyrolysis has been followed with a time resolution of several seconds. After pyrolysis, the O2 is added upstream from the fluidised bed and the combustion of the pyrolysis char can be followed in the same way. The equipment is calibrated by injection of a known C02-volume to the bed. Area and time response of the resulting calibration signal are needed for data analysis.
arburnup 4
0
200
400
600
800
------
1000
1200
time [SI Fig. 7: Concept for pyrolysis and combustion rate measurement in a fluidised bed of sand [20] Pyrolysis kinetics for cylindrical 12 mm diameter straw pellets in a fluidised bed of sand at different temperatures are shown in Fig. 8. The long pyrolysis times around 100 s are a consequence of the large particle dimension. This demonstrates, that pelletisation is not an advantage: it is expensive and destroys the high reactivity of untreated straw with thin ca. 0.5 mm thick walls. At comparable temperatures, the pyrolysis of straw chops proceeds more than 10 times faster than straw pellet pyrolysis. This is demonstrated in Fig. 9. Small cmsized single walled straw chops have been added to a large preheated steel vessel with a sand layer at the bottom. The pyrolysis gases quickly displace a certain amount of inert gas in the thermostated vessel. The released inert gas is collected in a burette, whose level is observed with a TV-camera. A time resolution of < 0.1 s has been obtained in this way. Fig. 9 shows that the pyrolysis times at 2 600°C are less than 10 s. Nodes in the straw stem halm must be squeezed to maintain short reaction times.
23 1
20
16
12
8
4
0
200
0
600
400
1000
800
time [s] Fig. 8: Pyrolysis kinetics for cylindrical 12 mm straw pellets in a fluidised bed [20]
7-7--+-+straw c ops
k
+
0
0
10
20
I
O
I
tlme(s)
30
I
,
,
,
,
;
,
,
,
,
40
Fig. 9; Generation of pyrolysis gases during straw chop pyrolysis
232
50
reciprocal temperature [I03/K] Fig 10.: Arrhenius diagram for combustion and C02-gasification of straw char in comparison to other chars (thermobalance data for a char particle size of 50 f 1OPm"l
In the Arrhenius diagram in Fig. 10 the straw char reactivity is compared to some other chars. Compared to the other chars the combustion reactivity of the straw char is relatively high, but this fact is not reproduced during gasification in C02. Ths is astonishing. The straw char contains the intrinsic ash with much potassium, which is reported to be a very efficient gasification catalyst. BEHAVIOUR OF POTASSIUM AND CHLORINE
Most of the particular difficulties of straw gasification are caused by the high K and C1-content. The behaviour of these impurities in the successive process steps is therefore of special importance. The selection of a method for their removal is a major process decision. Chlorine volatilisation starts at relatively low pyrolysis temperatures of about 200°C. About half of the chlorine can be removed into the pyrolysis gas up to about 500°C. The rest of the HCl is volatilised together with the potassium at higher temperatures. At lower pyrolysis temperature, K can be kept completely within the char particles together with the residual ash, but the chlorine distributes between char and gas. K and C1 can be completely leached from the straw with hot water in an additional pre-treatment step [8].We have found that the straw soaks somewhat more than its own weight of water; about 2 m3 of waste water per ton of straw with ca. 1% KCl will be generated even in an optimum countercurrent washing process. Concentration or distribution of the lean salt solution and drymg of the soaked wet straw chops are expensive additional operations. If the fresh yellow straw is left in swath on the field after harvest for about two month, L 150 mm of rain will wash out most of the K and C1. There are logistic 233
problems connected with this method and also the reduction of the K and C1 content is not reliable enough. The K and the residual half of the chlorine can also be leached from the straw char with hot water in the same way. If K and C1 are not completely removed prior to the high temperature gasification step, they are routed into the hot fuel gas in the form of volatile KCl and other volatile alkali salts. These salt vapours can be removed in a specially designed desublimation step in combination with the recovery of the sensible heat fiom the fuel gas. FLO WSHEET OPTIONS
Based on the two-stage concept with allothermal pyrolysis and autothermal secondary gasification, a large number of process variants can be generated. The selection of a simple and reliable process variant with regard to the type and design of the major components, suitable operating conditions and a good cold gas efficiency requires the accumulation of sufficient specific experimental experience with straw. The simplified process flowsheet in Fig. 11 represents just one of the favourite flowsheet options, which are still under investigation. rdepleted syngas air, oxygen to gasor water purificatioi
I
t
Fig. I I: Simplified flowsheet of a two-stage straw gasifier
Some options are summarised in the following. The basic two stage concept can be realised in two separate reactors or in a single reactor vessel with pyrolysis and gasification proceeding at two different sites. Pyrolysis gas and pyrolysis char may be gasified together, in two separate steps or the char may be separated and utilised for other purposes. Pyrolysis can be performed in a bubbling, circulating or a mechanically fluidised bed either with or without additional steam. Air or oxygen can be used for gasification. C1 and K can be removed prior to gasification by straw or char leaching with hot water in an additional pre-treatment step. Recuperators or
234
regenerators can be used to recover the sensible heat from the syngas. The pyrolysis bed can be heated internally or externally in a recycle loop. The possible flowsheet options are certainly not limited to the above examples. One central step considered in all flowsheet options, is an entrained flow gasifier operating at a flame temperature up to 1500°C. The hot tar-free syngas generated in the gasifier flame should be quenched with crushed char powder (not shown in the flowsheet) to increase the cold gas efficiency by chemical quenching [16]. CONCLUSION AND OUTLOOK Cereal straw, (wheat, maize, rice, barley) is among the cheapest renewables. The energy potential of straw and some related herbaceous biomass residuals from agriculture can contribute about 5% to our primary energy supply. This is a significant biomass contribution and straw holds a second place behind wood. Straw contains more ash, K and C1 than wood, therefore it is more difficult to manage technically. Compared to combustion, gasification is expected to be more flexible, efficient and environmentally compatible. Gasification technologies for straw are not well developed. Yet, the large amount of unused residual justifies the development of a special technology for optimum energetic utilisation. A gasification concept, which considers the particular properties of straw, has been selected as a suitable starting basis for further development. At present, thls concept is being investigated in the laboratory and on the bench scale. The work will be accompanied by systems analyses regarding technology, economy and market situation. The final aim are flow charts and design data for the components of a larger pilot plant. It is intended to construct and operate this pilot plant in co-operation with potential industrial users.
REFERENCES H. Kopek, T. Weber, W. Palz, P. Chartier, G. L. Ferrero; Proc. Of 10” Europ. Cod. Biomass for Energy and Industry, CARMEN, Runpar, Germany 1998 A.V. Brigdwater, D.G.B. Boocock (eds.); Developments in thermochemical biomass conversion, London 1997 DGMK Tagungsbericht 9802, Energetische und stoffliche Nutzung von Abfallen und nachwachsenden Rohstoffen, 20.-22. April 1998, Velen, Westfalen DGMK Tagungsbericht 2000-1, Energetische und stoffliche Nutzung von Abfallen und Biomassen, 10.-12. April 2000, Velen, Westfalen M. Kaltschmitt, A.V. Bridgwater (eds.), Biomass Gasification and Pyrolysis, CPL Press, Newbury UK, (1997), 24 W. Seifiitz, Der Treibhauseffekt. Hauser-Verlag Miinchen, Wien 1991 T. Kojima; “The carbon dioxide problem”, Overseas Publisher Association, Amsterdam 1998, ISBN: 90-5699-127-2 N.O. Knudsen, P.A. Jensen, B. Sander, K. Dam-Johansen in Ref. 1 page 244 H. Hartmann, A. Strehler; Die Stellung der Biomasse, Abschlussbericht fiir das BML, Landwirtschaftsverlag,Miinster 1995
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10 M. Kaltschmitt, G. A. Reinhardt (eds.); “Nachwachsende Energietrager”,
Vieweg, BraunschweigNiesbaden 1997 11 H. Bonmann, E. Henrich, L. Krebs, C. Leichsenring, E. Nieke, U. Richters, S . Rumpel, K. H. Scholl, E. Schroder, J. Seier, D. Wintzer; Marktanalyse zur Strohvergasung, FZK internal report (1997) 12 F. Zintl, T. 8hmann, in Ref. 1 page 1348 13 C. J. Fritz; Noel1 Konversionsverfahren zur Verwertung und Entsorgung von Abfallen, Emil-Freitag-Verlag,Berlin 1994 14 F. J. Schweitzer; Thermoselect Verfahren zur Ent- und Vergasung von Abfallen, Emil-Freitag-Verlag, Berlin 1994 15 W. Hahn, Vergasung nachwachsender Rohstoffe in der zirkulierenden Wirbelschicht, Dissertation, RWTH-Aachen 1994 16 B. Wolf, in Ref. 4 page 205 17 H. Jiintgen, K. H. van Heek, Kohlevergasung,Thiemig-Verlag, Miinchen 1981 18 C. Roesch, D. Wintzer; Monitoring ,,Nachwachsende Rohstoffe“, Vergasung wid Pyrolyse von Biomasse, Buro f3.k Technikfolgenabschatzung beirn Deutschen Bundestag, TAB-ArbeitsberichtNr. 49, Bonn 1997 19 A.A.C.M. Beenackers, K. Maniatis; Gasification technologies for heat and power from biomass, in: Chartier et al. (eds.), Proc. 9* European Bioenergy Conf., Copenhagen, Oxford (1996) 228-260 20 S. Rumpel; Die autotherme Wirbelschichtpyrolyse zur Erzeugung heizwertreicher Stiitzbrennstoffe, Dissertation Universitiit Karlsruhe (TH), FZKA-Bericht 6490, (2000)
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Gasification of “Liquidized” Biomass Supercritical Water Using Partial Oxidation
in
Yukihiky Matsumura’, Akitomo Kato2, Hiroyuki Sasaki2, Takuya Yoshida 1 Environmental Science Centel; University of Tokyo, 7-3-1 Hongo, Bunkyo-ku, Tokyo 113-0033 Japan 2 Department of Chemical System Engineering, University of Tokyo, 7-3-1 Hongo, Bunkyo-ku, Tokyo113-8656 Japan
ABSTRACE Pulverized cabbage was treated in an autoclave at 150°C to achieve “liquidization,” a conversion of solid biomass into slurry by thermal treatment. The product was a homogeneous slurry that could be pushed through a small diameter with small pressure. Thus, unlike untreated pulverized cabbage, which separates into liquid phase and solid sediments, the slurry can be easily sent into a high-pressure hydrothermal reactor. The liquidized cabbage thus obtained and cellulose, a model compound for biomass, were separately treated by partial oxidation in supercritical water at 4OO0C, 25 MPa. The effects of nickel and oxide catalysts on partial oxidation were investigated. Gasification assisted with partial oxidation was found to be effective for cellulose gasification, and a carbon gasification efficiency of 68% was obtained in 5 minutes using nickel catalyst. The liquidized cabbage was also gasified with the assistance of partial oxidation to produce a gas composed mainly of hydrogen, methane, and carbon dioxide. Oxide catalysts do not effectively enhance the partial oxidation or increase the carbon gasification efficiency, but they are good enhancers of the water-gas shift reaction, especially when mixed with the nickel catalyst.
INTRODUCTION WET BIOMASS GASIFICATION In view of their high growth rates and carbon neutrality, wet biomasses such as water hyacinth, algae, cattail, and banana tree are expected to become renewable energy resources for tomorrow’s sustainable society. However, the high moisture content of wet biomass prevents direct utilization or application of conventional thermochemical technologies. Biological technologies such as methane fermentation are applicable to wet biomass, but the treatment of fermentation sludge is often difficult, and a large fermentation reactor is needed since well over a week of retention time is required.
231
Thus, a hydrothermal conversion technique utilizing supercritical water gasification has been investigated as a promising technology for wet biomass utilization. Yu et al. [l] found that 0.2 M of glucose can be completely decomposed into hydrogen-rich gas in supercritical water at 600°C, 35 MPa. Xu et al. [2] succeeded in completely gasifying a 1.2 M glucose solution in supercritical water using a carbonaceous catalyst. Matsumura et al. [3] further investigated the carbon dioxide separation process applied to the supercritical gasification. A similar approach was conducted by Lee et al. [4]. Elliott et al. [5-81 successhlly treated organic wastes using nickel catalyst in both batch and flow reactors. Minowa et al. [9-111 also applied nickel catalyst for gasifying cellulose in a hydrothermal reactor. Kruse et al. [12] gasified polycatecol, a model compound of lignin, in supercritical water in the presence of potassium hydroxide. Boukis et al. [13] gasified glucose at 550 and 6OOOC using potassium carbonate catalyst. While hydrothermal gasification of biomass is a compelling field that attracts many researchers, two substantial difficulties remain in the supercritical water gasification process: solid feeding into the reactor and complete gasification at lower temperature. In this research, we attempted to achieve continuous gasification of wet biomass by applying biomass liquidization for solid feeding and partial oxidation for complete gasification at a lower temperature. BIOMASS “LIQULDIZATION”
Supplying solid biomass into a hydrothermal gasification reactor at high pressure is a difficult problem. The pressure in the hydrothermal reactor is much higher than the screw feeder can withstand. Moreover, it is difficult to apply a fluidized bed at the high temperatures and pressures encountered in supercritical water reactions. Nonetheless, the continuous feeding of solid biomass is indispensable. For an energy process using hydrothermal or supercritical reactions, the feedstock must be heated to bring it into a hydrothermal reaction field. The heat requirement for bringing the inlet flow to the reaction temperature is rather large. For example, assuming an inlet flow of pure water, as much as 2579 kJ kg-l must be added to raise the inlet flow temperature to 4OOOC under 25 MPa. This is almost the same as the enthalpy of evaporation of the same amount of water, which is 2676 kJ kg-I. Thus, the process efficiency is very low if there is no recovery of this heat from the outlet flow. For this heat recovery, continuous operation with the use of heat exchangers between the inlet flow and outlet flow of the reactor is needed. With a batch-type reactor, solid particles can be handled rather easily, but it is difficult to achieve heat recovery at a high efficiency. This is because the reactor has to be cooled down and depressurized for placement of the feedstock and recovery of the ash and char. Without continuous operation, the hydrothermal process would be inappropriate for an energy conversion process. The moisture content of wet biomass is sometimes as high as 95 wt%. A higher concentration of organic compounds is desirable for the feasibility of the process. If necessary, belt pressing of feedstock materials can be applied as pretreatment. Whether this pretreatment is needed depends on the efficiency of the heat exchanger. Several researchers have investigated the continuous supply of biomass feedstock to reactors at high pressures. By adding starch to a sawdust slurry, Dr. Antal’s group
23 8
at the Hawaii Natural Energy Institute successfully fed the slurry into their supercritical water reactor at 600°C, 25 MPa [14]. Specifically, it was the gelation of the starch that enabled the continuous feeding. In thls series of experiments, they successfully fed slurry with an organic content as high as 20 wt% into their reactor by gasifying the slurry with the help of a carbonaceous catalyst. According to their calculation, a 20 wt% organic content is required for their process to be economical. However, starch is not always a cheap by-product to use for sustaining an energy process. Later, the same group used potato waste from the food industry, a by-product that looks like mashed potato. Continuous operation of the reactor was also possible in this case [ 151. Organ0 Co. in Japan employed a pair of piston pumps for delivering sewage sludge into their reactor for liquefaction of the sludge [16]. The piston pumps themselves operated batch-wise, but the placement of feedstock into one pump while the other was feeding resulted in continuous operation. This system was effective for the treatment of sludge containing micro-organisms. Minowa et al. [171 at the National Institute for Resources and Environment, Japan, proposed treatment of city garbage at temperatures lower than 2OOOC to “liquidize” the garbage. This operation is different from conventional “liquidization” because the product material is not oil but a biomass slurry that can be obtained at much lower temperature. Specifically, they were able to liquidize a mixture of cabbage, steamed rice, clam shells, dried sardines, and butter employed as a garbage model by treating it in an autoclave. The product was a slurry, and solid components were found to precipitate with time. However, when they operated at 15OoC for 1 hour, the solid components could be suspended for more than several hours. They proposed that this process could be applied as an effective pretreatment for supercritical water gasification. For continuous feeding of biomass into a hydrothermal reactor, the approach by Minowa et al. [17] looks to be promising. Biomass is easily pulverized, but pulverization results in biomass slurry, whose solid phase easily settles down and separates fi-om the liquid phase. For example, Dr. Antal ’s group found that pulverized water hyacinth with a 20 wt% organic content did not have adequate fluidity, and when they attempted to feed this slurry into their reactor they ended up plugging up their feeding line [ 181. By applying heat to this pulverized biomass in the liquidization process, the solid structure is broken down and the affinity of the slurry to water is increased so that it can be kept homogeneous without settling of the solid content. Liquidized biomass can be continuously supplied to the reactor using piston pumps. Our present study was a hdamental investigation on the feasibility of biomass liquidization. Cabbage and water hyacinth were treated in an autoclave to determine whether liquidization takes place for only biomass, i.e., in the absence of steamed rice, clam shells, dried sardines, and butter, PARTLQL OXODATION OF BIOMASS
The biomass gasification processes so far proposed require operation at temperatures higher than 6OO0C or the use of expensive metal catalysts. Dr. Antal’s group achieved complete gasification by conducting supercritical water gasification at
239
6OO0C, 35 MPa [2]. When the operating temperature was lowered to 500 or 55OoC, the efficiency of the carbon gasification showed a dramatic decrease to around 50%. Lee et al. [4] applied temperatures as high as 700°C. Boukis et al. [13] applied a reaction temperature of 550-6OO0C. An operating temperature as high as 6OOOC is not desirable for the supercritical water gasification process. Since 100% heat recovery is not possible, a higher temperature usually results in higher heat loss. The heat supply to the inlet flow of the reactor is substantial, and this heat loss affects the process efficiency. Practically speaking, this high temperature poses problems in selecting the reactor materials. Because of the corrosive condition inside the reactor, a precious metal or alloy has to be used as the reactor material. Moreover, the wall of the reactor has to be thick enough to withstand the high pressure. The use of a precious material in large amount generates a high initial cost that makes economical feasibility difficult even over the long term. To lower the temperature, a precious metal catalyst is needed. Minowa et al. [9] could gasify cellulose at temperatures below 400"C, but only with the help of a nickel catalyst. Nickel is not cheap, and the application of a much cheaper catalyst is prerequisite for economical operation. In this study, we employed partial oxidation of biomass and investigated the effect of various catalysts on the gasification characteristics. As already found in the previous research, the efficiency of carbon gasification dramatically falls at lower temperatures around 4OOOC when no nickel catalyst is used [2]. The result is the formation of char or tarry materials. The addition of a proper amount of oxygen will increase the efficiency of the carbon gasification by decomposing these undesirable products. Partial oxidation will also form carbon monoxide, which will then be converted into hydrogen through the water-gas shift reaction. Dr. Antal's group found that the water-gas shift reaction plays an important role in the hydrogen productivity [19]. It is expected that the hydrogen thus produced will prevent the production of char and tarry materials. There may be worries about the cost associated with pressurizing oxygen or air. It is hue that in supercritical water oxidation, the delivery of oxygen into a high-pressure reactor is the main reason for the high operating cost. However, partial oxidation requires only one-tenth of the oxygen required for the supercritical water oxidation process, hence the process is still expected to be economically feasible. LIQUIDIZATION OF BIOMASS FOLLOWED BY PARTIAL OXIDATION
Thus, the liquidization of biomass followed by gasification supported by partial oxidation is expected to solve the problem of solid feeding into high-pressure reactors and enable complete gasification at lower temperatures. This study was conducted as a findamental investigation on the feasibility of this process. Specifically, we attempted the liquidization of biomass and examined both the liquidization characteristics and product fluidity. Further, we conducted supercritical water gasification supported by partial oxidation using cellulose as a model compound of biomass, and examined the effects of several catalyst in this process. Finally, we gasified liquidized biomass in supercritical water with the help of partial oxidation.
240
EXPERIMENTAL BIOMASS LIQUIDIATION EXPERIMENT Biomass liquidization In t h s study, we treated pulverized cabbage, a model biomass species, in the hydrothermal condition. Two kmds of reactors were utilized: a micro-reactor with an inner diameter of 7.52 mm and length of 150 mm, and a 200-mL autoclave. For the micro-reactor experiment, a cabbage was cut into 5-rnm squares, placed in the reactor, heated in an oil bath to 15OoC,and kept at that temperature for 60 min. At the end of the hour, the reactor was quenched by immersion in water and the product was taken out. For the autoclave experiment, about 60 wet-g of cabbage was pulverized in an electric juicer, placed in the reactor, heated by an electric furnace to 150°C, and kept at that temperature for 60 min. It took 10 min. to raise the temperature to150°C, and the pulverized cabbage and juice were stirred by a stirring propeller at the bottom of the autoclave during the whole 60-minute operation. After the operation, the autoclave was taken out of the furnace, cooled down, and the product was recovered. In both experiments, the pressure inside the reactor had to match the vapor pressure of the water coming out of the cabbage, i.e., 0.476 m a . The pressure indicator attached to the autoclave showed a value close to this pressure. The cabbage was pulverized by a home electric juicer, and the juice and residue obtained were used as a starting material for the liquidization. We used once-pulverized cabbage and thrice-pulverized cabbage. To obtain the latter, residue from the first juicing was pulverized two more times by the juicer. The pulverized cabbage was wet, and while its particle size distribution was not measurable, we could determine that a single pulverization resulted in particles between one and several millimeters, and three pulverizations resulted in particles of between several hundred micrometers and one millimeter.
Analysis of liquidization product The liquidized product was visually observed to assess its fluidity, and an Ostwald viscometer was applied as a trial to determine its viscosity. Some part of the product was put into a syringe to see whether the product could be smoothly purged form the syringe tip. Density measurements of the liquidized product were also taken. A fixed amount of well-mixed liquidized product was taken into a flask with a pre-determined volume, and then the flask was evacuated using a vacuum pump. Using a syringe, air was introduced into the flask in increments of 2 mL, and the increase in pressure with every 2-mL injection was measured using a pressure gauge. An equation of state for the ideal gas allowed us to calculate the volume of the gas phase in the flask, and subtracting thls value from the volume of the empty flask gave the volume of the liquid. The water content of the liquidized product and original cabbage was determined by drylng each sample at 6OoCand measuring the decrease in weight. FT-IR analyses were also conducted for the original cabbage and liquidized biomass for comparison.
24 1
GASIFICATION WITH PARTIAL OXIDATION Experimental operation for gasification with partial oxidation
Figure 1 shows the experimental apparatus used in this study. The reactors are stainless batch-type reactors (7.53mm ID x 100 mm in length, and 6.53 mm ID x 150 mm in length). A connecting tube made of 1/8 inch stainless tubing connects the reactor and the pressure gauge via a valve. This valve is used for oxygen supply for the partial oxidation and product gas sampling.
valve 7
Feedstock, catalyst, water, oxygen
__
3/8” stainless tubing
Fig. 1 Experimental apparatus.
After placing water, catalyst, and cellulose or liquidized cabbage into the reactor, the air inside the reactor was replaced by feeding in pure oxygen three times at 0.5 MPaG and releasing it back into the atmosphere. Then the reactor was fed with 0.3 MPaG of oxygen used for partial oxidation. Cellulose was placed in increments of 0.1 g, and catalyst was added in increments of 0.04 g. The amount of water was adjusted so that the water pressure at 4OOOC would reach 25 MPa. For gasification of the liquidization product, the amount was adjusted so that the water content of the liquidized product would equal the amount of water needed to attain a water pressure of 25 MPa at 400OC. Then a pressure transducer was attached to the reactor. The reaction was commenced by immersing the reactor into a molten salt bath preheated to 400OC. The reactor was kept in the bath for 5 min. and then quenched by immersing it in a water bath at room temperature. Once the reactor cooled, the pressure transducer was detached, the reactor was connected to the pre-evacuated gas sampling system (not shown in Fig. l), and the product gas was sampled by opening the valve. The amount of product gas ng was determined by the pressure increase in the sampling system, which was calculated as:
where, 8 and P,are the pressure before and after opening the valve, respectively; V, is the volume of the gas sampling system and V’ is the sum volume of the gas sampling system and reactor gas-phase volume in the reactor; R is the gas constant;
242
and T is the absolute temperature at the sampling system. The gas product thus collected was analyzed by a gas chromatograph equipped with a thermal conductivity detector and flame ionization detector. The pressure and temperature in the reactor were measured by a K-type thermocouple inserted into the reactor and pressure transducer, respectively. The molten salt bath was bubbled with air so to obtain a homogenous temperature inside the bath and effective heat transfer. The temperature in the reactor rose quickly, exceeding 390°C in 40 s. The pressure continued to slowly increase after the temperature reached 4OO0C,probably due to the gas formation. Materials The cellulose for column chromatography use (Macherey Nagel, MN100) was utilized in this study. The nickel catalyst was a commercial product (NEChemcat, Ni-5132P), and the oxide catalysts were prepared in the Takahashi Lab. of the Department of Chemical System Engineering, University of Tokyo. Table 1 shows the oxide catalysts used in this study. Table 1 Oxide catalysts used in t h s study. Catalyst
Chemical formula
ssc
~%.5I~~0.49~~1.0303
RESULTS AND DISCUSSION LIQUIDIZATION OF BIOMASS Fluidity of the liquidized product For the experiments conducted using a micro-reactor, effective liquidization did not take place. The product was composed of a water-like liquid and solid pieces that retained the shape of the original cabbage. The solid pieces separated from the water easily, and we could not make the whole product flow smoothly. This may have been due to a lack of stirring. When using a micro-reactor, the cabbage pieces placed in the reactor are not subject to any shear stress, preventing the effective dispersion of cabbage into the liquid phase. Autoclave liquidization successfully obtained products with good fluidity. Both the once-pulverized and thnce-pulverized cabbage yielded a homogeneous fluid slurry that was slightly more viscose than water. Figure 2 shows the thrice-pulverized cabbage and the corresponding liquidized product. The original pulverized cabbage easily separates into liquid phase and sediments of solid particles. On the other hand, the liquidized product is homogeneous, and the whole product can be fluidized. This result implies that the addition of steamed rice and other materials
243
applied in Minowa’s experiment to simulate garbage [ 171 is not necessary for biomass liquidization.
Fig. 2 Thrice-pulverized cabbage and the corresponding liquidized product.
Viscosity measurement of the liquidized product using the Ostwald visometer was not successful because the solid particles in the product plugged the capillary. However, the viscosity of the product was too small for the rotary viscometer. The pressure drop caused by the delivery of the liquidization product through the stainless tubing will be measured to determine the effective viscosity in the future. Although the capillary of the Ostwald viscometer was plugged, applying only a small pressure was sufficient to break apart the plugging structure and resume the flow of liquidized product through the capillary. A plastic syringe was used to roughly characterize the fluidity of the liquidized product. A sample was placed into a plastic syringe with an exit diameter of 1.25 mm and purged from the syringe tip by pushing the syringe piston. For the thrice-pulverized cabbage, the product was shaken to obtain homogeneity before the samples were taken. When the thrice-pulverized cabbage was purged from the syringe, the liquid was easily separated from the solid particles, and the solid particles were left in a packed bed at the end of the syringe. This may cause problems when the cabbage is delivered into the reactor. When the liquidized product was purged fkom the syringe, no phase separation took place and the solid particles were purged together with the liquid phase. Thus, liquidization yields a homogeneous slurry that can be treated as a homogeneous phase.
Composition of liquidizedproduct The moisture content of the liquidized product was 94.8%. This is close to the moisture content of the original cabbage which was found to be 94.4%. Thus, it appears that liquidization basically changes the solid form of the cabbage into fiuther
244
pulverized and softer particles, rather than involving chemical reactions which form or consume water. The density of the product was 1100 kg m-3. Judging from the h g h moisture content, the density should not be too different from that of water, and the observed value is reasonable in this sense. Figure 3 shows the result of FT-IR analysis for original cabbage and the liquidized cabbage. The original cabbage and its liquidized product give almost identical curves, firther supporting the proposition that liquidization changes the physical structure of the original cabbage rather than changing its chemical composition.
n
Fig. 3 FT-IR diagram of the (a) original cabbage and (b) liquidized cabbage.
245
GASIFICATION WITH PARTIAL OXIDATION
Efect of nickel catalyst amount Nickel catalyst is known to be effective for decomposition of cellulose, a model compound for biomass. As a reference to the oxide catalyst, the effect of nickel catalyst on gasification of cellulose with partial oxidation was measured by changing the amount of nickel catalyst added. The nickel catalyst was added in doses of 0, 0.01, 0.02, 0.03, and 0.04 g, and the amount of product gas was measured. As previously indicated, the temperature, pressure, and reaction time were set at 4OO0C, 25 MPa, and 5 min, respectively. Figure 4 shows how the addition of nickel catalyst affects the amount of product gas. The gas production increases linearly with the amount of nickel added. Without the addition of nickel catalyst, 1.2 mmol of gas was generated fiom 0.1 g of cellulose, and then an additional 0.7 mmol was generated for every 0.01-g of nickel catalyst addition. Carbon gasification efficiency, which is defined as the amount of carbon determined in the gas phase divided by the amount of carbon in the cellulose, increased fiom 31% to 68% with the change of nickel addition fiom 0 to 0.04 g. Table 2 compares the gas composition for the gasification conducted with and without nickel catalyst. The addition of nickel catalyst results in a higher concentration of hydrogen and methane, and lower concentrations of carbon monoxide and carbon dioxide. The increase in hydrogen and decrease in carbon monoxide suggests that nickel catalyst is a good catalyst for the water-gas shift reaction. Partial oxidation produces carbon monoxide, and then it is converted into hydrogen via the water-gas shift reaction. Nickel catalyst also enhances the rate of methanation. As a result, the addition of nickel catalyst enhances gasification of cellulose and the production of flammable gas.
0 .-E
"
.ba
a
u
L
n
0
0.01
0. 02
0. 03
0.04
0. 05
Weight of n i c k e l c a t a l y s t [ B 1
Fig. 4 Effect of nickel catalyst on gas production.
246
Table 2 Gas composition for gasification with and without nickel catalyst.
Without nickel catalyst With nickel catalyst
H2
co
0.089 0.301
0.251 0.033
CH4 0.019
0.138
co2
C2H4
C2H6
0.620 0.394
0.003 0.000
0.003 0.020
Catalytic eflect of oxide catalysts Oxide catalysts are known to be effective for oxidation reactions. In this study, we wanted to produce carbon monoxide through partial oxidation of the biomass, as this could be expected to lead to a conversion of carbon monoxide into hydrogen via the water-gas shift reaction. An oxidization of the tarry product is also expected. By these two effects, improvement of the efficiency of the gasification is expected. Oxide catalyst is expected to enhance the oxidation reaction needed for this scenario. Since oxide catalyst is considerably cheaper than nickel catalyst, its use would make the whole gasification process more economical. Hence, we decided to examine the effect of oxide catalysts on gasification with partial oxidation using cellulose as a model compound. Figure 5 shows the gas production with various catalysts. Compared to nickel catalyst, the effect of oxide catalysts is negligible. However, some catalysts such as SSC and LSM increased the hydrogen content and decreased the carbon monoxide content in the gas composition. This probably was due to the enhanced water-gas shft reaction. The carbon gasification efficiency was 68% with nickel catalyst, 3 1% without any catalyst, and 31% with the addition of SSC catalyst. Th~sshows that the oxide catalyst does not effectively enhance the gasification or the oxidation of tarry materials.
I . 1
none
SSC
LSM
BPC
SDC
Ni
Fig. 5 Gas production with various catalysts. Eflect of mixed catalyst Although oxide catalyst did not enhance the partial oxidation or increase the gasification efficiency, some of the other catalysts were found to be effective in enhancing the water-gas shift reaction. Therefore, added a mixture of nickel catalyst 247
and oxide catalyst to see whether the use of this mixed catalyst could help us reduce the h g h costs required for nickel catalyst alone. Gasification of cellulose with partial oxidation was conducted using a mixture of 0.02 g of SSC catalyst and 0.02 g of nickel catalyst. Figure 6 shows the product gas obtained with this mixed catalyst in comparison with amounts obtained with no catalyst, SSC catalyst, and nickel catalyst. The mixed catalyst resulted in a hydrogen production exceeding the average of levels obtained with 0.04 g of nickel catalyst and 0.04 g of SSC catalyst. Thus, the mixed catalyst is effective for hydrogen production. However, the carbon gasification eficiency with the mixed catalyst was 47%, which is just the average of the values obtained with the nickel catalyst and SSC catalyst, i.e., 68% and 32%. This mediocre carbon gasification efficiency again indicates that oxide catalyst is not effective for increasing gasification efficiency.
-p
4. 0
3. 0
E Y
c
.?
2. 0
c 3 0
z
*
1. 0
E . m
0.
0 none
ssc
NitSSC
Ni
Fig. 6 Effect of mixed catalyst on gasification of cellulose with partial oxidation.
Gasification of liquidized product
In gasification of actual biomass species, the presence of lignin and hemicellulose should cause different gasification characteristics from those normally seen with the gasification of cellulose. If partial oxidation is effective for lignin or hemicellulose, thus resulting in higher gasification eficiency, gasification with partial oxidation will be useful for practical operation. Liquidized cabbage was gasified with 0.04 g of SSC catalyst, a mixture of 0.02 g of SSC and 0.02 g of nickel catalyst, and 0.04 g of nickel catalyst. The weight ratio of organic compound to water was almost the same as gasification of cellulose (0.1 g of cellulose vs. 1.55 g of water; 94% water content). The reaction temperature, pressure, and reaction time were 400°C, 25 MPa, and 5 min, respectively. Fig. 7 shows the gas production for liquidized cabbage gasification. As in the case with cellulose gasification, nickel catalyst is the most effective catalyst for the gasification. However, the superiority of nickel catalyst over SSC catalyst is less prominent in liquidized cabbage gasification. This may be because of the lignin in cabbage, which is not suitable for gasification.
248
ssc
N itSSC
Ni
Fig. 7 Effect of catalyst on liquidized cabbage gasification with partial oxidation. CONCLUSION
With the intention of feeding wet biomass into a hydrothermal gasification reactor at high pressure, we experimentally investigated the liquidization of biomass followed by gasification of the liquidized product. The following results were obtained: (1) Pulverized cabbage was successfully liquidized using an autoclave with a stirring
propeller by keeping the feedstock at 150°C for 1 hour. (2) The liquidization product was homogeneous and could be purged fYom a syringe with an exit diameter of 1.25 mm, indicating that it may be possibly to smoothly feed such a product into a high-pressure reactor. (3) Gasification assisted with partial oxidation is effective for cellulose gasification and a carbon gasification efficiency of 68% was obtained for nickel catalyst at a reaction temperature as low as 4OOOC and reaction time as short as 5 minutes. (4) Liquidized cabbage was also gasified with the assistance of partial oxidation, and a product gas composed of hydrogen, methane, and carbon dioxide was obtained. (5) Oxide catalysts do not effectively enhance the partial oxidation or increase the carbon gasification efficiency, but they do enhance the water-gas shift reaction, especially when mixed with nickel catalyst.
ACKNOWLEDGMENT
This research was supported by the Proposal-Based New Industry Creative Type Technology R&D Promotion Program from the New Energy and Industrial Technology Development Organization (NEDO) of Japan. The authors are grateful for the help of Prof. K. Yamada, Department of Fine Materials Engineering, Shinshu University, and Dr. C.-j. Wen and Mr. M. Koyama, Department of Chemical System Engineering, University of Tokyo, for supplying the oxide catalysts.
249
REFERENCES
1. Yu D., Aihara M., Antal M.J., Jr. (1993) Hydrogen production by steam reforming glucose in supercriticalwater. Energy Fuels, 7,574-7. 2. Xu X., Matsumura Y., Stenberg J., Antal M.J., Jr. (1996) Carbon catalyzed gasification of organic feedstocks in supercritical water. Ind. Eng. Chem. Res., 35, 2522-30. 3. Matsumura Y., Minowa T., Xu X., Nuessle F. W., Adschiri T., Antal M.J., Jr. (1997) High-pressure carbon dioxide removal in supercritical water gasification of biomass. In Developments in Thennochemical Biomass Conversion (Ed. by A.V. Breidgwater & D.GB. Boocock), V01.2, pp864-77, Blackie Academic & Professional. 4. Lee 1.-G; Lee J.-S.; Kim M.-S. (1999) Hydrogen production by the gasification of biomass in supercritical water. Proc. 5th Korea-Japan Joint Symposium '99 on Hydrogen Energy, pp.365-72. 5 . Sealock, Jr. L.J.; Elliott, D.C.; Baker, E.G; Butner, R.S. (1993) Chemical processing in high-pressure aqueous environments. 1. Historical perspective and continuing developments. Ind. Eng. Chem. Res. 32, 1535-41. 6. Elliott, D.C.; Sealock, L.J., Jr.; Baker, E.G (1993) Chemical processing in high-pressure aqueous environments. 2. Development of catalysts for gasification. Ind. Eng. Chem. Res. 32, 1542-48. 7. Elliott, D.C.; Sealock, L.J., Jr.; Baker, E.G (1994) Chemical processing in high-pressure aqueous environments. 3. Batch reactor process development experiments for organics destruction. Ind. Eng. Chem. Res. 33,558-65. 8. Elliott, D.C.; Phelps, M.R.;Sealock, L.J., Jr.; Baker, E.G (1994) Chemical processing in high-pressure aqueous environments. 4. Continuous-flow reactor process development experiments for organics destruction. Ind. Eng. Chem. Res. 33,566-74. 9. Minowa T., Ogi T., Yokoyama S. (1995) Effect of pressure on low-temperature gasification of wet cellulose into methane using reduced nickel catalyst and sodium carbonate. Chem. Lett., 280, 285-6. 10. Minowa T., Ogi T.; Yokoyama S. (1995) Hydrogen production fiom wet cellulose by low-temperature gasification using a reduced nickel catalyst. Chem. Lett., 286, 937-8. 11. Minowa T., Fang Z. (1998) Hydrogen production fiom cellulose in hot compressed water using reduced nickel catalyst: product distribution at different reaction temperatures. J. Chem. Eng. Jpn., 31,488-91. 12. Kruse A., Danny M., Rimbrecht P., Sxhacht M., Dinjus E. (2000) Gasification of pyrocatechol in supercritical water in the presence of potassium hydroxide. Meeting Program for the 5th Int. Symp. on Supercritical Fluids, Poster Session pp.47-8, Apr. 8-12,2000, Atlanta. 13. Boukis. N., Schmieder H., Abeln J., Dinjus E., Kruse A. (2000) Gasification of high moisture biomass in supercritical water. Meeting Program for the 5th Int. Symp. on Supercritical Fluids, Poster Session p.5 1, Apr. 8-12,2000, Atlanta. 14. Angela G, Xu X., Antal M. J., Jr. (1997) Carbon catalyzed gasification of organic wastes in supercritical water. Proc. 4th Int. Symp. on Supercritical Fluids , Vol. 3., pp.875-80 (1997). 15. Allen, S. (1999) personal communication.
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16. Suzulu A., Nakamura T., Yokoyama S. (1990) Effect of operating parameters on thennochemical liquefaction of sewage sludge. J. Chem. Eng. Jpn., 23,6-11. 17. Minowa T., Dote Y., Sawayama S., Yokoyama S., Muakami M. (1995) Phase changing of garbage from solid to liquid slurry by thermal liquidization, J. Chem. Eng. Jpn., 28, 727-31. 18. Xu X. (1995) personal communication. 19. Nuessle F.W., Xu X., Matsumura Y., and Antal M.J., Jr. (1996) Hydrogen Production from High Moisture Content Biomass in Supercritical Water with Carbon Catalysis: The Role of the Water-Gas Shift Reaction, In. Hydrogen Energy Progress X I (Ed. by T.N. Veziroglu, C.-J. Winter, J. P. Baselt, & G Kreysa) Vol. 1, pp611-9, InternationalAssociation for Hydrogen Energy.
25 1
Pyrolysis and gasification of black liquors from alkaline pulping of straw in a fixed bed reactor G. Gea, R. Pukrtolas, M. €3. Murillo, J. Arauzo Chemical and Environmental Engineering Department, University of Zaragoza, Maria de Luna, 3, 50015 Zaragoza, Spain
ABSTRACT: New alternative processes such as low temperature gasification are currently being developed in order to use black liquor from pulp and paper mills for energy purposes. The development of these new processes makes necessary to study how black liquor behaves under different conditions. So, the present work is focussed on the study of pyrolysis and gasification of alkaline black liquor from pulping of straw, material for which few data are available in literature. Specifically, the influence of final pyrolysis temperature and pyrolysis heating rate on both, products (distribution and yelds) and char specific surface area have been studied. Additionally, the influence of gasification temperature on products has been also analysed.
INTRODUCTION Black liquor are industrial lignocellulosic wastes generated as by-products in the pulp and paper processes from wood, straw and other fibrous plants. Nowadays Tomlinson recovery boilers are used in these processes in order to generate heat from the organic compounds and to recover the cooking chemicals that are contain in the black liquor, nevertheless there are some safety and technical problems. These problems are enhanced in the case of black liquor coming from alkaline pulping of straw, rather than those coming from wood. This situation has caused the search and development of new alternative processes such as gasification. Alternative processes could be more energy efficient, safer and easier to control than conventional ones. So, in order to develop new processes, it is important a good understanding of the pyrolisis and gasification black liquor behaviour. Several studies about pyrolysis and gasification of krafi black liquor from cooking of wood can be found in literature, but there are scarcely data for black liquor coming from alkaline pulping of straw. The aim of this research is to provide more information about the thermochemical conversion of this residue. Studies [l] about COz gasification kinetics of alkaline black liquor from straw have been previously carried out in thermobalance by our research group. The present work is focus on the
252
study of pyrolysis and C 0 2 gasification processes for the same material, by means of a laboratory scale fixed bed reactor. In the pyrolysis experiments, the effect of temperature (T) and heating rate (p) on the product yields, gas composition and the energy recovery were studied by means of a fixed bed reactor. Additional measurements of the specific surface area of the resulting pyrolysis char were performed in order to determine its evolution at different pyrolysis conditions. This parameter is considered of great relevance in the reactivity of the char during gasification process. Finally several experiments of COz gasification have been performed in the same fixed bed reactor, in order to study the influence of the gasification temperature on the product yields.
EXPERIMENTAL As it has been pointed out in the previous paragraph, the aim of the present work is to provide more information about the behaviour, during pyrolysis and gasification of the alkaline black liquor from the pulping of straw. All experiments were performed in a fixed bed reactor on a laboratory scale plant. The studied variables were the heating rate and the final pyrolysis temperature. Their effects on the product yields, gas composition and energy recovery were analysed. The effect of the two cited variables on the specific surface area of the resulting char is also analysed. Furthermore, some C 0 2 gasification experiments were carried out at different temperatures in order to obtain some information about the influence of thls variable on the product yields.
MATERIALS AND EXPERTMENTAL SYSTEM The black liquor used in this research work comes from the soda pulping of straw. The main components of these black liquor are three: organic compounds, which come from the straw solution during the pulping, inorganic salts of sodium and potassium, and water. Moreover, it is important to indicate that the content of carbonates can reach 4.87 % wt. The water content in the studied material is approximately 90%. For the present research work the black liquor were dried up to 100% content of solids. This drylng stage needs to be gradual and carehlly controlled in order to reach an homogeneous distribution of the inorganic compounds inside the solid matrix. After drymg, the solids were ground, sieved under a particle size of 53 pm mesh and kept in a furnace at 105 "C. The laboratory scale plant is shown in Fig. 1. The reactor consists of a stainless steel tube, 90 mm inner diameter and 230 mm length, placed inside a concentric furnace. The particles of solid black liquor were placed in a tray of 40 pm mesh inside the reactor. Downstream the reactor exit, two ice traps and a cotton filter were installed to collect tars. Gas samplers and a CO/C02 I.R. analyser, permitted to follow the exit gas composition during the experiments.
253
Gas exit
9.
4.
"......I.
1.
8.
.........". 3.
I. I. Gas 2. Gas flow controller 3. Concentric furnace 4. Metalic tray holding sample of black liquors
5. Fixedbed reactor 6. Temperature controller
7. Ice traps and cotton filter 8. Gas sampling 9. LR. CO/Cq analyzer
Fig. 1 Scheme of the laboratory scale plant diagram.
Solid black liquor samples of 4 g were heated on a N2 flow of 100 d m i n (NTP) up to the final pyrolysis temperature, Tpb ("C) at a set heating rate, p ("C/min). Once the final temperature was reached, it was maintained for 90 minutes. All the performed pyrolysis experiments were as follows: (1) Influence of the final pyrolysis temperature, Tpb: 250, 300, 350,400, 450, 500, 550,600, 650,700,750, 800,850 and 900 "C respectively, at p = 5"C/min. (2) Influence of the heating rate, p: 5, 10, 15 and 30 "C/min., at Tpk= 850°C. As it has been previously pointed out, some C 0 2 gasification experiments were also carried out. The studied variable in this case was the gasification temperature. All these experiments were performed with the same initial pyrolysis stage, and a final pyrolysis temperature of 900 "C for p = 5"C/min. Once the final pyrolysis temperature of 900 "C was reached, it was kept constant for 90 minutes. After the pyrolysis stage the gasification stage took place by replacing the N2 gas flow by a N2-C02 gas flow of 1125 ml (NTP)/min with a 15% of C02. The studied gasification temperatures, T, were: 750,800,850 and 900 "C. SPECIFIC SURFACE AREA (Sg),PRODUCT YIELDS, GAS COMPOSITION AND ENERGY RECOVERY MEASUREMENTS. After pyrolysis, all the resulting char samples were ground and sieved to a particle size under 75 p.The specific surface areas (Sg) were obtained using BET adsorption method, by means of a Micromeritics Pulse Chemisorb 2700 apparatus and N2 at 77 K as the adsorbent gas. Due to the diffision problems of the Nz molecule inside the microporous structures, the BET specific surface area only gives
254
information of the macroporous area. The adsorption analysis provides the volume of Nz adsorbed with the relative pressure of NZ. Regarding to the study of the influence of the cited variables, Tpi and p, on the different product yields, gas composition and energy recovery respectively, total material and energy balances were determined. Moreover, material balances for gas products were also performed. Those calculations were determined using the gas concentration values measured by gas chromatography and by the I. R. analyser. The tar production was determined by weighting the tar-collecting device. The amount of char was determined by a similar method.
RESULTS AND DISCUSSION The pyrolysis experiments described above provided the following results about the effect of the two studied variables, Tpt and p, on both, the product yields and the specific surface area of the resulting pyrolysis char. PYROLYSIS RESULTS
Influence of the pyrohsis temperature. In thls section, it has been specifically studied the gas product dmibution along the thermochemical process, the final product distribution and the energy recovery as combustible exit gas. The experimental results were all obtained by GC analysis of the exit gas samples, which were collected at different temperature intervals during the pyrolysis process. The gas compounds detected were, HZ, CO, C 0 2 , CH,, Cz's (acetylene, ethylene and ethane) and C3Hs. In all the analysed samples the major gas products obtained were CO and COz. Fig. 2. shows the exit gas composition obtained in the TPi=60O0C experiment. As it can be noticed, the two compounds that first appear are CO and C02, which takes place in the range of 225 "C. The other gas products appear at around 300 "C for H2 and at 350 "C for the rest of hydrocarbons. The results obtained from the other experiments are not shown, due to the similarity to those shown in Fig. 2. From these results, it is generally observed that both CO and COz have a peak value at around 250 "C, and besides in the experiments corresponding to Tpi2 700 "C a second peak value higher than the first one appears in the case of CO. The maximum values for H2 are generally observed between 400 and 500 "C. For light hydrocarbons the peak appears about 400 "C. These results can be explained, considering that during the earlier stages of pyrolysis, due to descarboxylation CO and C02 are released from the material. On the other hand, H2 production occurs at higher temperatures, due to the fact that the breakdown of the aromatic rings requires more energy. Prior pyrolysis studies on alkaline black liquor from pulping of straw have been developed in our work research group [3]. In those studies the processed material was black liquor with 37 % in weight of solids. The experimental system used in that case was a fluidised bed reactor, and the results shown that the main gas product was H2 (about 65% in volume) followed by CO (about 30% in volume).
255
ji/
I 700
14 I
600 c. 500 400 300
'
tJ
4
200 E
2
100
0 0
50
150 time (min)
100
200
+ H2 0 co A c02 o CH4 x C2's + C3H8
8
0 250
Fig. 2 Pyrolysis gas composition versus time for Tph= 600°C.
Although these results were obtained from the same starting material, they point out an important difference if they are compared to those presented in this paper. In Fig. 3 it can be noticed that the main pyrolysis gas products obtained in the present research work were C02 and CO. These different results can be explained by the fact that in the present work the material processed was completely dry,while in the cited previous work [3] the processed material had an important humidity weight percentage (63%). This water leads up to the following reactions happen.
C (char) + H20+ CO + H2 CO + H20 + C02 + H2 CH, + H2O + CO + 3H2 tars + H203 gases Fig. 3 also shows an interesting result, the percentage of CO increases considerably as final pyrolysis temperature does it as well, which is specially remarkable at 600 "C. In the case of C02, its percentage remains practically independent of the final pyrolysis temperature up to around 750 "C. For temperatures over 750 "C, a certain increasing tendency is observed. The comparison of these results with those cited before [3] offers a new difference between them. In the referred case it was noticed that as final pyrolysis temperature increased, the percentage of each of the analysed compounds decreased (specially marked between 450 and 500 "C) except for H2. On the other hand, the global material balances were made for each one of the pyrolysis experiments, as they can provide important results related to the pyrolysis products distribution (distribution to char, tar and gases expressed as the sample initial weight percent). Fig. 4 displays these results. It can be observed that in general the tar yield is the lowest for all the final pyrolysis temperatures studied. Furthermore, it is important to notice that the char yield decreases as final pyrolysis temperature increases while the gas yield is increasing. For a final pyrolysis temperature of approximately 800 "C the gas production, expressed as the percentage of the sample initial weight, is higher than the char yield. Similar results are obtained in literature [2,4,51
256
45 40
I
I -
p=5OCmin-'
A
*H2 Q)
Z 25
0
A
A
-
co
A c02
CH4 x C2's 0
s
0
107
5L
0 0
I
0
+ C3H8
0
0
1000
500 Tpir
("C)
Fig. 3 Pyrolysis gas product yields versus (T,L).
80 A
70
p=5 "Cm in"
%gases rn %tars A % char
01 0
1000
500 Tpir
("C)
Fig. 4 Pyrolysis product yields versus Tpir.
The energy recovery is defined as the ratio of the low heating value of the generated gas formed and the low heating value of the black liquor. For this ratio, the following equation has been used: Energy recovery =
c
LHVgas i . mgas i
* 100
LHVliquor mliquor
(5)
Where: LHV,,, i: Low Heating Value of the gas i [kcal.g-'] LHVliquOr: Low Heating Value of the dried alkaline black liquor in study [kcal.g-
'I
mga,i: mass of the gas i generated during pyrolysis [g]
257
mIiauor:mass of the dried alkaline black liquor fed [g] The LHVliquor, determined in The Institute of Paper Science and Technology of Atlanta, is 3683 kcalkg. Fig. 5. shows both the percentage of energy recovery and the gas yield obtained for the different Tpk studied. It can be noticed that both, the gas yield and the energy recovery increase as final pyrolysis temperature increases too. So the energy yield varies from a 3% in the case of TPk= 250 "C up to 50% in the case of T,k = 900 "C. The main contribution to the energy recovery comes from the lower yield of pyrolysis gases: such as H2 and hydrocarbons. 80 1 70
4
,
+ gas yield
* .
50
u)
m
0
1
+
10 -
00
8 . .
200
I
400
600
800
1000
Fig. 5. Pyrolysis gas yield and final energy recovery versus Tpk.
Influence of the pyrolysis temperature on the char specific surface area As it was early introduced the char specific surface areas (Sg) were obtained by application of BET adsorption method in order to study the influence of the final pyrolysis temperature on this char property. In all cases the BET isotherm fits satisfactory with correlating index of about 0.99. The specific surface areas of pyrolysis chars measured are represented in Fig. 6. It can be noted that until Tpk about 600 "C the specific surface area values are lower than 10 m2.g-', corresponding with poor specific surface area chars. These results are probably obtained due to a macroporous surface generated from de thermal decomposition of the organic matter contained in the studied black liquor, [6].At about 650 "C a higher specific surface area char (Sg r~ 28 m2.g-') is generated. From 700 "C to 800 "C, the Sg of chars experiments an important increase (Sg r~ 100 m2.gI ). This increase in the Sg value is probably due to the thermal decomposition of the inorganic compounds such as sodium and potassium carbonates. This decomposition provides small molecular gases, CO and C02 , that create a microporous structure as they leave out the solid matrix. Finally, from Fig. 6 can be observed that at the higher temperatures studied, 850 "C and 900 "C respective1 the measured Sg value goes on increasing with Tpk, reaching the value of 300 m?.g- in the case of 900 "C.
'
258
300 250 200 150
I0 0 50 0
.
= 50~rnin-'
* * *
t,**w++ 200
600
400
800
1000
T p ~ r("C)
Fig. 6 Specific surface area (Sg) of the black liquor chars obtained at different Tpk. The reactions, which cause the thermal decomposition of black liquor, are varied. Several mechanisms can be found in literature, most of then are centered in sodium emissions from black liquor during their pyrolysis and gasification [6]. Some of them are the following:
There is no certain of which one is the dominant. It is known that carbonate decomposition and sodium release start about 675°C. These carbonates are from both, the chemical pulping of lignocellulosic materials and from the conversion of the sodium organic compounds contained in the black liquor during pyrolysis under 675°C. The registered CO and COz profiles versus time obtained by I.R. equipment during the process corroborates what it has just been explained. Fig. 7 shows these profiles measured as percentage in volume of CO and COZ in the exit gas for two different final pyrolysis temperature experiments, 650 and 750°C respectively. These two experiments have been selected for their graphical representation due to the fact that both correspond with the first leap in Fig. 6. Fig. 7 is expected to allow us to explain this leap. Both profiles, CO and COz are very similar up to 650°C. In the case of CO two peaks appear, one about 250°C and the other one about 500°C. For COz a peak also appears at about 250°C and reaches its maximum value at 400 "C. A second C 0 2peak appears at the same time that the second CO peak. These CO and COz peaks are due to the thermal decomposition of the organic matter contained in the black liquor as it has been explained before.
259
p = 50~rnin" 800
8 -%
g6
2 0
-Yo
C02 650 CO 750
600
4
400
s2
200
0
' A
L3
3 L
Q)
0 0
5000
10000 time(s)
15000
0 20000
42
Fig. 7 CO and COz time for different T,k At 650°C some differences can be found at T,i of 650°C and 750 "C. In the first case, during the isotherm stage it is observed a CO peak at the same time that a C 0 2 profile valley appears too. Later, at 10300 s a lower COZpeak appears. For a final pyrolysis temperature of 750°C the CO rises rapidly up to 5 3 % . Meanwhile the C 0 2 does not reach the 2%. All those results are concordant with other found in literature [6, 71 where the CO is the majority product from the carbonate decomposition in presence of C. Up to 650°C CO and C02 come from thermal decomposition of organic matter and from this temperature the inorganic decomposition starts mainly carbonate decomposition. As these reactions are endothermic they are thermodynamically favoured a higher temperatures. So, as temperature increases both peaks, CO and COz, are greater, mainly the CO ones. Influence of the heating rate on theproducts In this section, the experiments have been carried out in the same way, as it is describe on the experimental section for pyrolysis experiments. The studied variable has been the heating rate. Four different values have been used: p = 5, 10, 15 and 30"C/min. Fig. 8 shows the heating program followed in each case. Focussing the attention, firstly, in the gas products, Fig. 9 displays the percentage of initial sample weight transformed in the different gas products. It can be observed that H2 percentage increases slightly as heating rate does. In the case of CO it is observed an initial increase from p = 5"C/min up to p=IS"C/min, reaching a 25% value. From p = 15"C/min up to P=3O0C/min the value remains approximately constant. For COz the results are different, so at the lower heating rates, 5 and I0 "C/min, it reaches values up to 25%, but afterwards for higher heating rates the percentage of C02 decreases. HC's percentage follows a similar tendency to CO and Hzrso the HC's percentage increases as heating rate increases up to p = 15°C. It has been also considered interesting to study the influence of the heating rate on product distribution: tar, gases and char. Results are presented in Fig. 10. A small influence of the heating rate on the pyrolysis products yields was observed.
260
800
f!3 600
.cI
2
o)
E
g
400 200
I
-..-. B=lOoC/min .......B=lSoC/min - - - - B=300C/min
1
0
0
50
150
100
200
250
300
time (min)
Fig. 8 Heating paths during pyrolysis experiments at different heating rates.
g25; 20 =15E $10#. 5 -
0 ,
A
A 0
0
A
A
0
0
0 HC's 0
0
0
0
0
0
0
0 I
70 60
T,,,=800°C
4 4
4
A
A
4
A
.
1
04 0
10
.%tars A % char
1
20
30
40
fl (OClmin) Fig. 10 Pyrolysis product yields versus heating rate (p). With regards to the energy recovery, it can be observed in Fig. I I that the energy recovery calculated by equation (5) increases as heating rate does from p = 5 up to p = 15"C/min. From p = 15"C/min to p = 30"C/min it decreases slightly.
26 1
70 60 f g50 uE ~ 4 0
a 0
3 830 "
0
Gas yield
'5. $ 2 0 8 I0
8
Q
O
J 0
5
10
15
20
25
30
35
p (Wmin) Fig. 11. Pyrolysis gas yield and final energy yield versus heating rate (p). Influence of the heating rate on the char specific surface area
The measurements of char specific surface area presented in this part have been obtained following the procedure described above in EXPERIMENTAL. Fig. 12 summarises the results obtained. The graphic shows a Sg value (120 rn2.g-') more or less independent of the heating rate up to p = 15"C/min. From p = 15"C/min to p = 3OoC/min an important increase in the Sg value can be obserbed. Probably this increase is due to a change in the mechanism of black liquor decomposition that affects the released gas. Similar works on carbons have shown the lmportant increases of the char Sg value with the heating rate [8]. 250
-
200
-5 150
*
E m 100
*
u)
50 0
0
I
I
I
10
20
30
p (ocmin") Fig. 12 Specific surface area (Sg) of the char obtained at different heating rates (p). GASIFICATION RESULTS
Although this paper is focussed on the study of alkaline black liquor pyrolysis, additionally some gasification experiments have been performed, all of them with the same pyrolysis stage previous to gasification. The studied variable has been the gasification temperature (TW) at 750, 800, 850 and 900°Crespectively. The gas flow used (1 125 ml(NTP)/min) has been selected in order to avoid both, dikssional resistance and solids pulling during gasification.
262
It is important to take into account that the char C 0 2 gasification from alkaline black liquors is a very quick process. This fact can be observed in Fig. 13 corresponding to the C02 profiles (% in volume at the reactor exit) at different gasification periods. For all runs the C02 percentage achieves a value that is close to the C 0 2 content of the inlet gas (15%). 16
= 0
14 12
+G-750 - -D- - G-800 +G-850 - - - Q - - G-900
.-5' 108 C
i 0
6
4 2
0 0
5000
10000
15000
time (s) Fig. 13 C02profiles versus time for different gasification temperatures. More information about black liquor gasification can be obtained from the analysis of the CO profile, this is shown in Fig. 14. CO is a gasification product meanwhde C 0 2 is the gasification medium. As it can be observed in Fig. 14 all the curves presented a maximum peak, which appears about 180 s after starting gasification. On the other hand this results show a higher CO production as gasification temperature increases. Moreover at the lower temperatures (750 and 800°C) the CO in the exit gases is detected later than in the other cases (850 and 900 "C). 16
+G-800 --+-G-850
- 8 - G-900
1
10
1000
100
10000 100000
time ( 8 )
Fig. 14 CO profiles versus time for different gasification temperatures. These gasification results will be complete with a future study in order to understand the relationship between the char specific surface area and the gasification reactivity.
263
ACKNOWLEDGEMENT The authors express their gratitude to the Commission for Cultural, Educational and Scientific Exchange between the United States of America and Spain for providing financial support (Project 99017)
REFERENCES 1. 2. 3. 4.
5. 6. 7. 8.
Gea, G. Murillo, M.B. and Arauzo, J., Proceedings of the Fourth Biomass Conference of the Americas, 1999, vol. 2,969-975. Bhattacharya P.K., Shrinath AS., Kunau D., J. Chem. Tech. Biotechuol., 35 A, 1985,223-233. Sinchez, J.L., Ph. D. Thesis, Thermochemical processes at low temperature of black liquor from alkaline pulping of straw, 1999, University of Zaragoza. Frederick W.J., Hupa M., Uusikartano T., Bioresource technology 48, 1994, 5964. Sricahroenchailkul V., Phimolmas V., Frederick W.J., Grace T.M., Journal of Pulp and Paper Science, vol24,2, 1998,43-50. Li J., and Heiningen A.R. P., Tappi Journal, 1990,213. Yuh S.J., Ph.D. Thesis, University of Notre Dame, 1984 Maloney D.J., Jenkins R.G., Fuel, vol64, 1985, 1415-1422.
264
Effect of fuel size and process temperature on fuel gas quality from CFB gasification of biomass A.van der Drift, J. van Doom Energy Research Foundation (ECN), P. 0.Box I , NL-I 755 ZG Petten, The Netherlands
ABSTRACT A bench-scale CFB-gasifier with a capacity of max. 500 kWh has been used to study the effect of fuel size and process temperature. As also found by many others working on gasification processes, a higher process temperature (range tested: 750 to 91OOC) results in more air needed to maintain the desired temperature, a lower heating value of the product gas, a hgher carbon conversion and a net increase of cold gas efficiency of the gasifier. A higher process temperature also results in less heavy tars. However, light tars (measured using the SPA-method) show an odd behaviour. Some individual components within the group of light tars even increase in concentration when process temperature is raised. The main reason probably is that heavy tars decompose to these relatively stable light tar components. The particle size of the fuel does influence some product gas parameters considerably. The presence of small particles seems to increase the (heavy) tar concentration and decrease the conversion of fuel-nitrogen to ammonia. Small particles can also be responsible for large temperature gradients along the axis of the riser of a CFB-gasifier. T h s effect can be avoided by either mixing the fuel with larger particles or feed the small particles at the bottom of the reactor.
INTRODUCTION Converting biomass to useful products like electricity can be done in many ways. Gasification is considered as a good option with relatively high efficiency due to the possibility to use integrated cycles (gas turbines and steam turbine) for the larger scales and gas engines for smaller scales. Circulating fluidized bed (CFB) technology is considered as a flexible way to convert solid biomass fuel to a gaseous fuel. In this paper the effect of fuel size and process temperature on the quality of the product gas is presented. Experiments have been carried out with the bench-scale CFB test facility at ECN.
265
CFB-FACILITY
DESCRIPTION OF THE FACILITY The circulating fluidized bed (CFB) gasifier, called BIVKIN and situated at the Netherlands Energy Research Foundation (ECN in Petten, the Netherlands), is an atmospheric air blown facility of about 500 kWh [I]. Figure 1 shows a picture of the facility. It is equipped with various feeding systems, which can be used simultaneously in order to be able to feed fuel mixtures. The riser is a 20 cm diameter and 6 meter high refractory lined pipe. Generally the actual velocity (top of riser, actual temperature) is in the range from 6 to 8 d s . The circulation loop contains a bubbling fluidized bed seal. The feeding system is purged by air or nitrogen. The produced fuel gas is either flared or sent to a gas cleaning section and a gas engine to produce electricity. The bed material used is sand (97% silica) of approximately 500 pm in diameter.
Figure 1. CFB facility at ECN
MEASUREMENTS During each experiment, the concentrations of H2, CO, CH4 and C02 in the fuel gas are measured continuously using TCD (H2) and NDIR detectors respectively. The amounts of ethene, acetylene, ethane, benzene, toluene and xylene are measured discontinuously (every 10 or 20 minutes) using gas chromatography. The concentration of NH3 is determined by wet chemical methods. Tar is measured by two methods. “Light” tars are determined by the SPA-method [2]. These “light” tars comprise molecules from xylene to about 5-rings. “Heavy” tars are measured gravimetrically: a known volume
266
of gas is de-dusted by a hot filter at 250°C and led through a filter at 125°C. The increase in weight is measured. Note that our definition of heavy tars differs from the definitions used by others. Absolute values can therefore not be compared to data on heavy tar from other institutes. Our definition of light tars ("all compounds detected with the SPA method plus GC-FID analysis") is comparable to the definition used elsewhere.
FUELS Two kind of fuels have been used for the tests reported here: demolition wood and willow. Demolition wood pellets are purchased at Labee (Moerdijk, the Netherlands). This fuel does not show any feeding problems and for this reason is often used by ECN for research where he1 parameters are not the first interest. Willow is purchased from IMAG-DLO (Slootdorp, the Netherlands) where research is focussed on energy plantation. Table 1 shows the chemical composition of both fuels. Complete analyses can be found in [3].
Table 1 Composition of fuel used in CFB-experiments, daf: dry and ash-free basis, HHV: higher heating value. willow demolition wood 1.7 ash wt% dry 1.9 Wkgdaf 19797 20183 HHV 49.9 C wt% daf 49.7 5.95 H wt% daf 6.00 0 wt% daf 43.6 43.7 N wt% daf 0.63 0.42 S wt% daf 0.056 0.01 c1 wt% daf 0.0 12 0.025 F wt% daf 0.004 0.003 EXPERIMENTAL RESULTS PROCESS TEMPERATURE With demolition wood pellets as fuel, the process temperature has been varied between 750°C and 910°C. Figure 2 shows some fuel gas properties as a h c t i o n of temperaure. An increase of process temperature means that more air is needed to reach the desired temperature. The equivalence ratio ER is used to indicate the amount of air. ER is defined as the actual amount of oxygen divided by the theoretical amount needed to stoichometrically combust the fuel. The increase of ER with increasing process temperature is the result of the fact that more energy is needed to heat up the gases. Because the increase of air results in more fuel gas (to be heated to the desired temperature) due to the dilution of the nitrogen from the air, this effect is non-linear. The ER almost doubles when process temperature is raised from 750°C to 910°C. As can be expected, the carbon conversion increases with increasing temperature. This is the result of increased reaction rates at higher temperature. Also the increase of ER and the resulting increase of the amount oxygen in the system will have a positive effect on carbon conversion. This effect is so high that the cold gas
267
efficiency of the gasifier increases when process temperature is increased despite the extra loss due to more (diluted) fuel gas at high temperature. The concentration of heavy tar in the product gas reduces significantly when process temperature is raised. The reasons are similar to the ones mentioned above for the carbon conversion. The light tar concentration (SPA-method) does not seem to be lower at higher temperature. This probably is the result of the fact that (1) light tars are relatively stable molecules compared to heavy tar components and (2) heavy tar components decompose at higher temperature to smaller molecules, whch at least partly fall into the category of light tar. The last effect is illustrated by the increases of the concentration of some stable light tar molecules at increasing process temperature. For example, the concentration of phenanthrene, fluoranthene and pyrene increases with 35%, 80% and 60% respectively when process temperature is raised from 750°C to 910°C.This effect is even greater when realised that the concentration of inert nitrogen increases with increasing temperature. The heating value of the product gas decreases as can be expected due to the increase of ER and the resulting dilution by nitrogen. Apart from this, also the concentration of hydrocarbons like methane and ethene decrease due to hgher reactivity at higher temperature. These components have a relatively high heating value. A reduction of the concentration therefore results in a reduction of the HHV of the fuel gas. ,ZOO%
150%
50% 0%
L L
I
700
750
800
, 850
900
950
temperature [“C] I
Figure 2 The effect of process temperature on some fuel gas properties during CFB-gasification expressed as relative values, Results at 750°C = 100%.
FUEL SIZE Willow has been used to study the effect of fuel size. Different sizes have been prepared: saw dust with a maximum size of 2 mm, cylinder-shaped willow with a diameter of 10 mm and a length of 10 mm or 40 mm and a mix made by a commercial shredder. In figure 3 some product gas parameters are shown for both the abovementioned fuels and two mixtures. From Figure 3 it appears that the tar concentration is related to the presence of small particles. Small particles increase the tar concentration. The reason for this might be that small particles are entrained easily and therefore are present in relatively high
268
amounts at the top of the riser. Pyrolysis products and tar molecules therefore are generated in the top of the reactor and only have little time to decompose in the hot reactive atmosphere before leaving the reactor. In Figure 4, the temperatures along the axis of the riser are shown for three experiments. In all cases the fuel was fed at 1 meter above bottom. What strlkes most is the large temperature gradient measured during the test with saw dust. Over 150°C difference in temperature has been measured! On the other hand, the test with 10*10 mm cylinders show normal flat temperature profiles with maximum temperature differences of only 30°C. The reason for the low temperature homogeneity with saw dust as a fuel is that these small particles largely are blown upwards. Because these small particles are relatively quickly converted, almost no particles are circulated via the circulation loop. This means that in the first meter of the riser almost no fuel is present and hardly any gas is generated. Therefore, due to the lack of product gas and consequently a low gas velocity, little bed material will be blown into the section above the fuel feeding point. This means that the temperature homogenising medium (bed material) is only present in small amounts in the section where all the chemical reactions take place.
1 -
Ed carbon conversion
175%
-1
m heavy tar 0 NH3/fuel-N
150%
1
m HHV gas a
125%
a -
9 .-9
n
.-
100%
e
75%
50%
25% 2'2 mm
10'10 mm
10'40 mm
2'2 mm + 10'1 0 mm
10'10 mm + 10'40 mm
mix from shredder
Figure 3 The effect of fuel size on some product gas parameters during CFB gasification with willow, 2*2 mm: saw dust, 10*10 mm: cylinders with 10 mm diameter and 10 mm long, 10*40 mm: cylinders with 10 mm diameter and 40 mm long. Indicated mixtures are 50%/50% by weight. All values are relative to the saw dust test.
269
775
1 0
I
1
2
3
4
5
location above air distributor [m]
Figure 4 Temperature profiles in the riser of a CFB gasifier using different fuel sizes, fuel has been fed 1 meter above the bottom of the reactor.
By mixing saw dust with larger particles (see Figure 4), the disadvantages disappear and maximum temperature differences along the axis of the riser reduce to almost normal values. Not shown in the graph are the results of a test where the saw dust has been fed at the bottom of the reactor. Temperature profiles tum out to be flat for this test as can be expected. So, small fuel has to be either fed at the bottom of the reactor or mixed with larger particles to avoid problems related to temperature homogeneity. One of those problems is agglomeration of the bed particles [4,5].
CONCLUSIONS The gasification temperature largely influences the amount of air needed to convert the solid biomass into gaseous hel. With this, the heating value of the produced gas decreases dramatically when process temperature in raised from 750 to 910°C. At the same time the carbon conversion increases and the concentration of heavy tars decreases. However, light tars as measured using the SPA-method do show an odd behaviour. Some individual components within the group of light tars even increase in concentration when process temperature is raised. The main reason probably is that heavy tars decompose to these relatively stable light tar components. The particle size of the fuel does influence some product gas parameters considerably. The presence of small particles seems to increase the (heavy) tar concentration and decrease the conversion of fuel-nitrogen to ammonia. Small particles can also be responsible for large temperature gradients along the axis of the riser of a CFB-gasifier. This effect can be avoided by either mixing the fuel with larger particles or feed the small particles at the bottom of the reactor.
270
ACKNOWLEDGEMENT The Dutch agency for energy and environment (Novem) and EnergieNed, Federation of energy companies in the Netherlands, are greatly acknowledged for their (financial) contribution to the work reported in this paper. REFERENCES 1.
2. 3.
4. 5.
Description of the CFB gasification plant at ECN on the internet: http://www.ecn.nl/unitfb/bivkin/index.html(l998). C. Brage, Qizhuang Yu, Guanxing Chen and K.Sjostrom: Use of amino phase adsorbent for biomass tar sampling and separation. Fuel 76 ( 2 ) 137-142 (1997). Biomass and waste composition database on the internet: http://www.ecn.nl/phyllis/. (1999). A. van der Drift and A. Olsen: Conversion of biomass, prediction and solution methods for ash agglomeration and related problems. Final Report., Petten, ECN, ECN-C--99-090,62 p. (1999). A. van der Drift, H.J.M. Visser and A. Olsen: Prediction and solution methods for ash agglomeration and related problems during biomass conversion. In: First World Conference and Exhibition on Biomass for Energy and Industry, 5-9 June 2000, Sevilla, Spain (2000).
27 1
Biomass ash - bed material interactions leading to agglomeration in fluidised bed combustion and gasification H.J.M. Visser, H. Hofmans, H. Huijnen, R.Kastelein and J. H.A. Kiel Netherlands Energy Research Foundation (ECN), P.O. Box 1, 1755 ZG Petten, The Netherlands.
ABSTRACT: The present study has been aimed at improving the fundamental understanding of mechanisms underlying agglomeration and defluidisation in fluidised bed combustion and gasification of biomass and waste. To this purpose dedicated labscale static heating and fluidisation experiments have been conducted with carefully selected and prepared ashes and bed materials, viz. straw ashhand and willow ashhand mixtures, mullite subjected to straw gasification and artificially coated mullite. The main conclusion is that ash/bed material interaction processes are very important and often determine the bed agglomeration and defluidisation tendency. In the static heating experiments with both ashhand mixtures, partial meltingkegregation of ashcomponents and dissolutionheaction with the bed material are processes that determine the melt composition. This melt composition and behaviour can deviate considerably form expectations based on ash-only data. Artificially coated bed materials prove to be very useful for systematic studies on the influence of coating composition and thickness on agglomeration tendency. For the coated mullite samples, different stages in the defluidisation process are identified and the influence of coating properties (thickness, composition, morphology) and operating parameters is elucidated. The behaviour of the mullite appears to be dominated by a remnant glass phase. On the one hand, this glass phase accounts for an alkali-getter capability, while on the other hand it is mainly responsible for agglomeration at temperatures 2 800 "C.
1
INTRODUCTION
In (bubbling) fluidised bed combustion and gasification of biomass and waste, several potential problems are associated with the inorganic components of the fuel. A major problem area is defluidisation due to bed agglomeration. Based on existing knowledge, it is clear that in fluidised bed combustion and gasification with continuous fuel feeding, this defluidisation is a self-promoting process. At the onset of bed agglomeration, the tluidisation behaviour gets disturbed due to grain clusters and, as a result, uniform heat distribution is no longer possible. Continuous fuel feeding then leads to local peak temperatures, which promote further bed agglomeration and detluidisation. As a consequence, reactor operation must be stopped to replace the bed
272
material. In general, any measures that can postpone the renewal of bed material are important improvements for commercially operated reactors. The main objective of the present investigation is to improve the fundamental understanding of the mechanisms underlying bed agglomeration and defluidisation. An improved understanding of these mechanisms may lead to better predictive capabilities and to strategies for preventing/postponing bed agglomeration and defluidisation. Attention is focused in particular on the behaviour of alkali metals, since they are present in relatively large quantities in biomass ashes and they are of major concern with respect to bed agglomeration. At lower temperatures, alkali metals may react with chlorine and sulphur to form low-melting salts, while at higher temperatures they may react with the bed material to, e.g., silicate melts.
2.
APPROACH
The adopted approach has been developed based on morphological and chemical analyses of numerous agglomeration samples, obtained from lab-scale to commercialscale fluidised bed reactors operated with a wide range of biomass (waste) fuels. From these samples, two extreme types of agglomerates can be identified. One extreme results from "melt-induced" agglomeration. In this case, the bed material grains are "glued" together by a melt phase, which roughly matches the chemical composition of the ash. Generally, this kind of agglomeration can be attributed to the occurrence of (local) peak temperatures well above the targeted operating temperature. The other extreme type of agglomerate is more common and results from "coating-induced" agglomeration. Here, a (uniform) coating is formed with time on the surface of the bed material grains. At certain critical conditions (e.g. coating thickness or temperature), neck formation may occur between coatings of individual grains, which initiates the agglomeration. In general, this type of agglomerates is formed at longer operating times and the agglomerates have little mechanical strength. Le., they can be crumbled on squeezing or even fall apart on touching. If, however, after first neck formation partial defluidisation of the bed leads to local peak temperatures, melt formation may occur and a combination of the two extremes can be recognised in one sample. Although a limited amount of agglomerates can be characterised as being one of the two extreme types, the identification of these extremes has served as a useful concept for the design of experiments to study agglomeration mechanisms and associated interaction processes between bed material and ash-derived components. The following experiments have been performed: 1) Static heating of quartz sandwillow ash and quartz sandstraw ash mixtures to simulate melt-induced agglomeration. 2) Static heating of coated mullite (AI2SiO5)bed material obtained from a bubbling fluidised bed gasification experiment with straw to study coating-induced agglomeration. 3) Static heating of artificially coated mullite bed material. The coating is composed of K-silicate as a 2-component analogue of the coating observed after straw gasification. The advantage of the artificial coating is that important parameters such as coating thickness, heating period, temperature and starting composition can be varied independently. In addition, the simplicity of the chemical system
273
allows for a more detailed study of the interaction between bed material and ashderived components. 4) Fluidisatioddefluidisation experiments on coated mullite bed material. These experiments have been performed to study the effect of the coating on fluidisation behaviour and the mechanisms leading to defluidisation. The choice of mullite as bed material is not an arbitrary one. Since alkali components are the main problem causing elements in agglomeration, alkali-getters such as kaolinite are often used as additives in an attempt to postpone defluidisation. However, knowing that kaolinite is a low temperature precursor of mullite and assuming that the Al-Si chemistry determines the potential as an alkali-getter, mullite could work just as well. Moreover, its crystalline (nearly equi-axed) equilibrium shape, temperature resistance and hardness allow mullite to be used as a bed material. Therefore, mullite might be an attractive bed material with inherent alkali-getter capabilities. 3.
EXPERIMENTAL
3.1
MATERIAL PREPARATION AND CHARACTERISATION
3.1. I
Ashhand mixturesfor static heating experiments
Straw ash and willow ash were prepared by ashing finely ground raw material at 550 OC overnight in a box furnace. The bulk composition of the ashes was determined by lnductive Coupled Plasma - Atomic Emission Spectroscopy (ICP-AES) (see Table 3). Ash was then mixed with quartz sand in a ratio of 1.3:8.7 by weight (roughly 1:1 by volume). 3.1.2
Coated w l l i t e bed material from a straw gasiJcation test (WOB sample)
Coated mullite bed material was obtained from a lab-scale bubbling fluidised bed gasification test with straw at ECN. Before conducting the gasification test, the mullite starting material was characterised by Scanning Electron Microscopy (SEM). It appeared that a glass phase, being more silica-rich (82 wt% SO2, 18wt% A1203)than the crystalline bulk (32 wt% Si02, 68 wt% A1203),was present up to a few volume percent in nearly all the grains (see Figure 1). This glass phase is a remnant of the mullite fabrication process. The gasification test was conducted in the so-called WOB facility and actually was a standard agglomeration test as described in detail elsewhere (Drift and Olsen, 1999). During the 7-hour test, the first 4 hours comprised gasification with continuous fuel feeding at 750 OC. Then, the temperature was increased step-wise (25 "C/step, keeping the temperature constant for 45 minutes) until at 840 "C defluidisation occurred (before reaching the intended 850 "C). Afterwards the sample material was easily removed at room temperature and the bed material fell apart into a loose powder indicating a poor mechanical strength of the agglomerated bed as expected in case of coating-induced agglomeration. SEM analysis showed that the mullite bed material grains were covered by an evenly distributed coating with an average thickness of -3 pm and a composition of 77% SO2, 12% K20,8% AI2O3,3% CaO and 1% MgO (see
274
Table I). Since the composition of straw ash shows only an insignificant amount of A1203,the coating is thought to be a product of ash and bed material interaction. In addition, the SEM analysis showed that potassium had diffised into the remnant glass phase inside the mullite grains but not into the crystalline bulk (for compositions see Table 1 and 2). Both in the coating and in the remnant glass phase (up to the very centre of the grains) the K 2 0 content appeared to be 12-16 wt%. No Ca or Mg diffused into the remnant glass phase (or in the mullite). Figure 1 shows the different mullite sections before and after straw gasification. It can be observed that the volume of the glass phase has increased dramatically. From chemical analyses it is apparent that part of the mullite has dissolved in the glass phase. This shows that the remnant glass phase has served as a potassium-getter while concurrently reacting with the mullite crystalline material.
Figure 1. SEM images of mullite grain cross-sections. Left: material as received showing a minor amount of remnant glass phase. Right material after straw gasification showing a large increase in the volume of glass phase (lighter colour, due to K-content). The coating is not visible on this scale.
3.1.3
Artiycially coated mullite bed material
The K-silicate coating was intended to be a simple (2-component) analogue of straw ash (27 wt% K 2 0 and 73 wt% SO2) to be applied on mullite with an average grain size of 250 pm. The coating was produced using a sol-gel process (e.g. Hamilton and Henderson, 1968 and Visser, 1999). The bed material was coated by mixing the mullite grains through the prepared gel, drying it at 80 "C overnight, heat-treating it at 600 OC for 24 hours, and finally washing it with water to rinse of any excess K-silicate. The coating thickness was varied by applying multiple coating procedures resulting in a coating thickness ranging from 0.5 micron (1 coating) to -2.2 micron (3-5 coatings). While an increase in coating thickness was achieved up to 3 coatings, the average thickness after 3, 4 and 5 coatings remained the same (within the error margins), suggesting that through this sol/gel process a maximum coating thickness of -2.2 micron can be achieved. With respect to the coating composition, the intended K20/Si02 ratio could not be obtained unfortunately, because part of the potassium remained in the gel (see Table I). Noticeable is that the potassium content differs between the coating and the
275
remnant glass phase. There seems to be a preferential uptake of potassium by the remnant glass phase up to a certain (equilibrium) value whereafter the potassium in the coating further increases to a value probably dictated by the equilibrium composition of Si/AI/K at 600 "C. In other words, the remnant glass phase shows alkali-getter capabilities with respect to the coating as well. Table I . Coating composition and thickness. The mullite starting material contains 32 wt% Si02 and 68 wt% A1203.The coating of the WOB sample also contains 1 wt% MgO and 3 wt% CaO.
Table 2. Chemical composition of the remnant glass phase.
3.2
EXPERIMENTAL PROCEDURES
3.2. I
Static heating experiments
Static heating experiments were conducted with the ashhand mixtures, the WOB sample and artificially coated mullite. Samples were prepared by pressing sandash mixtures into cylindrical pellets (1.5 glpellet), then placing these pellets in small golden crucibles inside alumina containers with an alumina top plate. Starting dimensions of the pellets were -1Omm diameter and lOmm height. For the ashhand mixture, a good distribution of the ash between the sand grains was confirmed by optical and scanning electron microscopy (SEM). The experiments consisted of a heat-treatment followed by assessing the degree of agglomeration afterwards by visual inspection and SEM. The applied temperature/time combinations were 650 "C - 1 and 6 hours; 725 "C - 1 and 6 hours; 800 "C - I , 2.5, 5 and 6 hours; 900 "C - 1 and 6 hours. The alumina containers were
276
placed in a box furnace, already heated to the desired temperature and afterwards allowed to cool at room temperature. Allowing only minimal space above the samples (- 2 mm height) and using a ceramic top plate, evaporation was minimised. In general, weighing the samples before and after heat treatment showed no loss of material. 3.2.2
Fluidisation experiments
The fluidisation experiments were conducted in a 5 cm diameter fluidised bed reactor made of quartz glass and placed in a split furnace. This allowed for visual inspection of the fluidisation behaviour during experimentation up to 900 OC. Preheated nitrogen (300 OC ) was applied as the fluidisation gas. For all experiments, the bed height was 3 cm and the gas velocity was kept at 1.5 times the theoretical minimum fluidisation velocity. The bed temperature ranged from 600 to 900 "C. During the experimental runs the following visual observations were used as qualitative (de)fluidisation characteristics: 1) the formation of vertical channels, 2) first grains defluidise on top of the bed, 3) 90% of the bed surface is defluidised, 4) the total bed is defluidised. When these situations occurred, time, temperature and pressure differences were recorded. A disadvantage of these fluidisation experiments without continuous fuel feeding, is that in this case agglomeration is not a self-promoting process. On the contrary, when part of the bed defluidises, the gas flow through the remaining bed volume is increased, which helps postponing further defluidisation. Realising this, emphasis will be put on the early stages (initiation) of the defluidisation process.
4
RESULTS AND DISCUSSLON
4. I
STA TIC HEA TlNG EXPERIMENTS WITH ASHBAND MIXTURES
Analysis of the 20 samples after heat treatment, using scanning electron microscopy with chemical analysis by energy dispersive x-rays (SEM-EDX), shows the following results. For straw ashhand, the first occurrence of a melt phase appears in the sample heated at 725 OC for 6 hours. A transition from a partially molten to a fully molten ash phase occurs in the sample heated at 800 OC for 6 hours and all ash is also molten in the samples heated at 900 OC (see Figure 2). For willow ashhand, the first melting occurs in the sample heated at 650 OC for 6 hours, thus at a lower temperature then for straw ashhand. However, even at 900 OC not all the ash is molten (see Figure 3). Roughly 1/3 of the original ash content remains as a non-molten phase and SEM-EDX-analyses show that this fraction is highly enriched in Ca and P compared to the bulk ash composition (see Table 3). All samples show an increase in volume during agglomeration due to asWmelt material contraction in the grain-to-grain contacts and the presence of gas bubbles in the melt phase. The amount of gas bubbles decreases with increasing temperature. In general the straw ashhand samples contain a larger volume of bubbles than the willow ashhand samples. Morphologically a major difference between both mixtures is apparent. In the willow ashhand samples, the first melt phase tills up the cracks in the sand grains and other narrow spaces, but also forms a well attached, uniform coating on the
277
0
grain surfaces. This indicates a low viscosity and a good wetting behaviour. In contrast, the first melt phase in the straw ashhand samples contracts into the pores between the sand grains. A change in the distribution of the melt phase occurs with increasing temperature when the melt phase gets more "attached" to the sand grains which are hence effectively glued together. The large bubbles present in the straw ashhand melt phase and the contraction of this phase in the pores at 725 OC, suggests a high viscosity and a poor wetting ability, which however changes with temperature. Chemical compositions of the melt phases are listed in Table 3. The appearance of very irregular solidmelt interfaces (see Figure 2 and 3) suggests that part of the (SiOz) sand grains has dissolved in the melt. Especially for willow ashhand this leads to a remarkable melt composition with a very high SiOz content relative to the ash composition before experimentation.The chemical composition of the melt after heating at 650 "C is not in agreement with the morphological distribution requiring low viscosity and good wetting ability; a composition with 70-80% SiOz is highly viscous. This is additional evidence that the melt phase in this sample is a reaction product, with quartz either dissolving in a potassium-rich melt phase, or K diffusing into the cracks and forming a melt phase in situ. For both mixtures the melt phase is mainly a K-silicate and after heating at 725 and 800 OC the melt compositions are very similar. In case of the willow ashhand samples, a large fraction of the ash is not included in the melt resulting in a segregated Caphosphate phase. In the willow ashhand samples heated at 900 OC, part of the melt clearly is a product of a reaction between the K-silicate and the previously segregated Ca-phosphate phase. At other locations not adjacent to Ca-phosphate, the melt is still a K-silicate. The straw ashhand samples show some indication of quartz dissolving upon melt formation, but the melt composition changes little when increasing the heating temperature to 800 and 900 OC, where finally all ash is molten.
The main objective of the static heating experiments was to elucidate interaction processes between biomass ash and bed material grains with respect to the agglomeration tendency. From the observations, it can be stated in general that ash/bed material interactions are very important. Differences in interaction have been established for different ash/bed material mixtures, but also as a knction of heating temperature. In the temperature range 725-800 OC, the straw ashhand mixture shows a much stronger agglomeration tendency then the willow ashhand mixture. This is despite the fact that the first melting takes place at a lower temperature for the latter mixture. Apparently, the melting in the willow ashhand mixture involves only a small ash fraction with all Ca-phosphate remaining segregated from the melt. Above 900 OC, however, where Ca-phosphate partially reacts with the K-silicate, the melt volume increases and the agglomeration tendency may become more severe. In conclusion, the bulk ash composition and the temperature of first melting are not necessarily good indicators for a problematic fuel/bed material combination. The amount of melt, its composition and wetting characteristics, and the melt - bed material interaction are thought to be more important factors.
278
Table 3. Chemical composition of the ash before, and the different melt phases after, heat treatment as well as first melt characteristics.
Composition biomass ash after ashing at 550 OC (wt%)
Conditions for first melt First melt morphology
First melt composition (wt%)
Melt composition at 800 OC / 6 hours (wt%)
Melt composition at 900 OC I 6 hours (wt?!)
straw ashhand samples Si02 = 59.3 K20 =22.7 CaO = 6.0 P205 = 2.2 A1203= 0.1 MgO = 2.1 SO3 = 4.2 CI = 3.4 725 OC16 hours -many large gas bubbles -melt contracts to pores -+ non-wetting melt Si02 = 70 KzO = 16 CaO = 8 P2O5 = 4 SiOl = 71
willow ashhand samples Si02 = 5.8 KZO = 19.1 CaO =47.0 P205 = 14.4 A1203= 1.0
MgO= 6.0 so, = 4.5
-4
-small gas bubbles -melt fills cracks melt wets grain surfaces + ood wettin melt Si02 = 84 K20 = 16
(Ca-P hasese re ated) Si02= 78
K 2 0 =13
K20 = 16
CaO = 7 P205 = 5
CaO = 3 (Ca-P phase segregated) 2 different melts Si02 = 75 Si02= 37
Si02 = 74
K20 = 1 1 K20=2 CaO = 7 CaO=43 P205 = 1 PZOS = 14
K20 = 13
CaO
=
P205 =
7 4
279
Figure 2. SEM Back scatter electron micrograph of a straw ashhand sample heated at 900 "C for 6 hours. Dark areas are pores, middle grey are sand grains and light grey rims are,the molten ash.
Figure 3. SEM Back scatter electron micrograph ofa willow ashtsand sample heated at 900 "C for 6 hours. Dark areas are pores, middle grey are sand grains and light grey rims and small clusters are the ash components. Only part of the ash has molten and a segregation has occurred.
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4.2
STATIC HEATING EXPERIMENTS WITH COATED MULLITE
Most samples from full-scale installations show bed material grains on which (alkalirich) surface coatings have lead to agglomeration. These samples indicate that coating thickness and chemical composition are important parameters. Therefore, further static heating experiments were conducted with coated (mullite) bed material. 4.2.1
Static heating experiments with WOB samples
For the WOB samples, full agglomeration occurred at 800 OC after 2.5 hours and at 725
"C after 6 hours, indicating a dependence on heating time. This agglomeration temperature is 40-140 "C below the defluidisation temperature in the WOB gasification test. From Figure 4, it is evident that many of the locations where agglomeration has occurred, involve remnant glass phases emerging at the surface of the grains.
Figure 4. SEM image of a typical agglomerate section in a WOB sample after static heating. 4.2.2
Static heating experiments with artificially coated mullite
Static heating experiments were performed with mullite, that was subjected to 1, 2, 4 and 5 coating procedures respectively. After cooling, the degree of agglomeration in the samples was established as shown in Figure 5.
28 1
-\
G
I,
e
E! 5
-
-\ - - - - - il I
I;
800
--
t-"
I
1
WOB
\
t l
I
600
I
I
0
1
2
3
4
3
4
Coating thickness
0
I
2
Coating thickness (micron)
Figure 5. Degree of agglomeration for artificially coated mullite and the WOB sample as a function of heating temperature and coating thickness. Top: heating time 2.5 hours; bottom: heating time 6 hours. 0 = no agglomeration, I = number of few-grain clusters present, 2 = large number of grain clusters, 3 = full agglomeration. The two contours indicate the transition from 0 to 1 and from 1 to 2/3. The bars reflect the spread in the coating thickness.
The results illustrated in Figure 5 show that, in general, the degree of agglomeration is dependent on temperature, coating thickness and heating time. A single coating is not sufficient to initiate any agglomeration up to 1000 OC. However, above a critical coating thickness, agglomeration seems to become independent of coating thickness. The lower agglomeration temperatures for the WOB sample are probably due to a difference in grain surface composition; i.e. a larger amount of remnant glass phase and additional Ca and Mg in the coating. After heating, the coating morphology changes very little in all samples. Even after heating at 1000 O C , most of the coating remains as an evenly distributed layer around the grains with some evaporation occurring only at a few locations. However, heat treatment has a distinct effect on the morphology of the remnant glass phase inside the mullite grains. First, the volume of the glass phase increases, indicating that part of
282
the mullite dissolves into the glass phase. Secondly, the glass phase is redistributed to crystallographically determined locations similar to the morphology in the WOB sample as illustrated in Figure I . The increase in volume and orientation in bands result in more glass on the outside of the grains. These surface locations with remnant glass phase seem to be the main locations for agglomeration at temperatures 2 800 "C. The role of the glass phase may explain the strong temperature dependence and the weak dependence on coating thickness. 4.3
FLUlDlSA TlON EMERlMENTS WlTH COA TED MULLITE
In the fluidisation experiments, main differences occurred between the samples with 1 and multiple coatings, and between the first and second fluidisation run with the same sample. During the first run, the 1 coating sample showed the formation of channels in the bed at 860 OC, while for the samples with 2 , 4 and 5 coatings this occurred at 650 f 30 "C already. The first grains defluidising on top of the bed were observed at 720 It 20 "C for the 2 , 4 and 5 coating samples, while for the 1 coating sample this stage was not reached up to 900 "C. The 90% surface defluidisation stage, observed for the 2,4, and 5 coating samples, took place at 770 It 30 "C. Total defluidisation was observed to depend largely on bed history during the run with the temperatures of total defluidisation showing no consistent trend. The different behaviour of the I coating sample is in agreement with the results of the static heating tests, where for the 1 coating sample no agglomeration occurred up to 1000 "C. All three second runs, for samples with 2, 4 and 5 coatings, showed lower temperatures for channel formation (below 600 "C) and first defluidisation, while the temperatures for 90% defluidisation of the top of the bed were equal or somewhat lower. Again 100% defluidisation did not show a consistent trend. The lower temperatures indicate that permanent changes occurred in the sample during the first run. Afier each run, it was attempted to re-fluidise the bed by increasing the gas flow rate suddenly to double or triple values. For the first runs, the bed could not be refluidised at the high tinal temperature, but after cooling to room temperature refluidisation was at least partially possible. In case of partial re-fluidisation, complete re-fluidisation could be achieved by stirring the bed with a small spoon. This is in line with the behaviour of relatively loose agglomerates resulting from coating-induced agglomeration. For the second runs, the 2 coaling samples showed the same refluidisation behaviour as in the first run, but for the 4 and 5 coating samples the bed could be re-fluidised at high temperature already. Apparently, "sticky" (viscous drag) forces can be overcome to some extent by increasing the fluidisation velocity, effectively postponing defluidisation to higher temperature levels. It seems that this was confirmed by the fluidisation experiment with the WOB sample. Here channel formation started at 740 "C and all subsequent defluidisation stages occurred at this same temperature level. This temperature is 100 "C lower than in the WOB gasification test, while in that test the gas velocity amounted to 5.5 times the minimum tluidisation velocity compared to 1.5 times for the fluidisation experiments. However, the temperature difference should probably also partially be attributed to a time ef'fect involving similar permanent changes as suggested above for the artificially coated mullite.
283
SEM analysis revealed that these permanent changes are related to the behaviour of the remnant glass phase, as it was observed in the static heating experiments already. After the second fluidisation run, the coating of the artificially coated samples showed a more molten appearance at some point contacts between the grains, while at other locations the coating seemed unchanged. To be clear, molten appearance does not mean that the coating was really molten. The theoretical melting temperature of the connecting regions between the grains was always at least 100-150 "C higher than the applied operating temperature. From the chemical compositions of moltenhon-molten surface locations, as presented in Table 4, it becomes clear that the molten locations coincide with the remnant glass phase and a K20content of 2 8 wt%, while at the non-molten locations the K20content amounts to only 2-3 wt'??. The nonmolten locations show lower K 2 0 concentrations than before the fluidisation experiments, implying that the remnant glass phase has taken up more K from the coating (K-content depleted from 8 to 2-3 wt%). Apparently, the buffer capacity of the glass phase is higher at higher temperatures. In addition, the volume of glass phase has also increased and this remnant glass phase has become very mobile, sometimes even leaving holes in the grains (see Figure 6). The change in glass volume is irreversible which explains why the second runs show lower temperatures of defluidisation than the fist runs. These effects are more pronounced for the 4 and 5 coating samples than for the 2 coating sample due to the higher maximum temperature level. In conclusion, the results of these fluidisation experiments agree very well with the findings of the static heating experiments. Clearly, with the mullite bed material the remnant glass phase plays a crucial role. One the one hand, this phase is mainly responsible for the alkali-getter capacity, while on the other hand, surface locations with remnant glass phase are the main locations for agglomeration at temperatures 2 800 "C. Table 4. Chemical composition of molten and non-molten surface locations after two fluidisation runs.
Sample
Max. temperature ("C)
mullite + 2 coatings 750
"Molten" Si02 appearance (wt%)
K20
A1203
(wt%)
(wt%)
Yes
8
13
79
284
Figure 6. SEM image of a mullite grain from the WOB sample showing a hole in a remnant glass phase section.
CONCLUSIONS
In general, it is concluded that ash/bed material interaction processes are very important and often determine the bed agglomeration and defluidisation tendency. Therefore, the properties of both materials and the nature of their interaction should be considered when predicting these tendencies. This conclusion appeared clearly from the static heating experiments with sand/ash mixtures, which showed dissolution and reaction processes between both materials determining the melt composition. These interaction processes are temperature dependent, resulting in a temperature dependence of chemical and physical properties (e.g. viscosity, wetting ability) of the melt phase. The static heating experiments with mullite bed material subjected to straw gasification and with artificially coated mullite revealed that the artificial coating method can be applied successfully for more systematic studies on the mechanisms of bed agglomeration and detluidisation. For the coated mullite bed material, it appears that below 1000 O C a certain critical coating thickness is required to initiate bed agglomeration. At larger coating thicknesses, bed agglomeration becomes independent of coating thickness and temperature becomes the most important operating parameter. However, in the static heating experiments also time had a clear et'fect, even though no additional low melting (alkali) components were added, and in the tluidisation experiments some intluence of the fluidisation velocity was seen. The tluidisation experiments with coated mullite bed material revealed several stages of detluidisation, viz. I ) the formation of vertical channels, 2) first grains
285
defluidise on top of the bed, 3) 90% of the bed surface is defluidised, and 4) the total bed is defluidised. It is realised, however, that these stages may be passed through quickly in actual fluidised-bed gasification or combustion due to the continuous fuel feeding. Therefore, in fluidisation experiments with coated bed material to elucidate agglomeration mechanisms, attention should be focused on the initial stages of defluidisation. For the coated mullite, agglomeration was not found to be associated with (partial) melting of the coating. The theoretical melting temperature of the connecting regions between the grains was always at least 100-150 "C higher than the applied operating temperature. Mullite, as a crystalline material, showed no alkali-getter capability. However, the remnant alumino-silicate glass phase inside the mullite grains did. The buffer capacity is one of the dominant interaction processes reducing the agglomeration tendency at low temperatures (< 800 "C). The K-buffering capacity increases with temperature. However, by dissolving simultaneously part of the mullite the volume of glass phase also increases considerably with temperature. Above 800 "C, the increased area of glass surface locations in combination with increased softening of the glass phase results in the glass being the most problematic phase with respect to agglomeration. Hence, the potential,of this material in a combined bed material/ alkali-getter finction is limited to low temperature applications. REFERENCES Drift van der A. and Olsen A. (1999) Conversion of biomass, prediction and solution methods for ash agglomeration and related problems. Final Report EUproject JOR3-95-0079 (ECN-report ECN-C-99-090). 2. Hamilton D.L. and Henderson C.M.B. (1968). The preparation of silicate composition by a gelling method. Min. Mag., 36, 832-838. 3. Visser H.J.M. (I 999) Mass transfer processes in crystalline aggregates containing a fluid phase. Ph.D.-Thesis, No. 174 in Series Geologica Ultaiectina, Utrecht University. 1.
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CASST: A new and advanced process for biomass gasification H. den Uil Netherlands Energy Research Foundation (ECN), PO box I , 1775 ZG Petten, The Netherlands
ABSTRACT: Large-scale biomass gasifiers currently available produce a gas with a high tar content. For almost all downstream applications a substantial reduction of the tar concentration is required. CASST is a new biomass gasification process for the production of a fuel or synthesis gas with a low tar content. In the CASST process, biomass is fist converted into charcoal and volatiles. The charcoal is gasified with steam and the volatiles are combusted to supply the heat required for gasification. F’yolysis experiments showed that in the CASST process the charcoal will contain 30-50 wt.% volatiles. Gasification of this charcoal gives a gas with tar concentrations that are 10-50 times lower than for wood gasification. Furthermore, the tars produced by charcoal gasification have a lower molecular weight than the tars produced by biomass gasification. The energetic performance of a CASST/combinedcycle (=CASST/CC)system has been compared with an air-blown gasificatiodcombinedcycle (=ABG/CC) system. At a carbon conversion of 80% in the charcoal gasification step, the electrical efficiency of the CASST/CC system is lower than for an ABG/CC system, 39.0 vs. 4 1.7% (LHV). Increasing the carbon conversion in the charcoal gasification step to 95% almost eliminates the difkence in the performance of both systems. A cost estimate for a 30 MW, plant showed that the capital investment for a CASST/CC system and an ABG/CC system are comparable, -2700 ECUkW,. INTRODUCTION Bubbling or fluidised bed gasifiers are used for the large-scale, above 5-10 MW&, gasification of biomass. The gas produced by fluidised bed gasifiers has a high tar content, -10 g/Nm3 (1). Although the tolerance of downstream applications towards tars is a subject of discussion, it is clear that for almost all applications a considerable reduction of the tar concentration is required. Tar removal or conversion is considered as one of the major problems to be resolved before use of the gas produced by biomass gasification is possible (2). Several techniques can be used for removal or conversion of the tars: scrubbing, catalytic cracking or thermal cracking. However, all these techniques present disadvantages.
287
A biomass gasification process that produces a gas with a low tar content avoids the problems associated with tar removal or conversion. For large-scale gasification, no process is commercially available for the production of a gas with a low tar content. A new concept for the production of a he1 or synthesis gas with a low tar content has been developed by ECN: the CASST process (CASST=Clean Air-blown Sustainable Syngas Technology). PROCESS DESCRIPTION Gasification of a solid feedstock consists of three steps: 1. Drying; to evaporate moisture in the feedstock. 2. Pyrolysis; to devolatilise the feedstock giving: gas, vaporised tars and a solid char residue. 3. Gasification; to convert the tars and char formed during pyrolysis into gasses. The tars formed in biomass gasification originate fiom the volatiles released in the pyrolysis step. In the gasification step the tars are partially converted by gasification, but also give rise to the formation of secondary and tertiary tars (1). CASST is based on the hypothesis that a reduction in the volatile matter content of the gasification feedstock results in a decrease of the tar content of the gas produced by gasification. In the CASST process (see figure I ) biomass is first separated by low temperature pyrolysis at 350-400°C, into volatiles and charcoal. The charcoal is gasified with steam at about 800-900°C to produce fuel or synthesis gas. The volatiles are combusted with air at 1000-1100°C to supply the heat required for the endothermic gasification step. A circulating inert heat carrier, e.g. sand, transports the heat between the combustion and the gasification step. Since the CASST process is an indirectly heated process, no air is used in the gasification step. Therefore, the product gas will have a low nitrogen content giving a medium calorific fuel gas and enabling use of the product gas as synthesis gas. Air
Combustion
Charcoal Steam
Flue gas
Syngasl Gasification
Figure 1: Simplified flowsheet for the CASST process
CHARCOAL PRODUCTION To maximise the cold gas efficiency of the CASST process, heat supply by combustion of volatiles and heat demand by gasification of charcoal should be balanced. A preliminary system assessment study showed that charcoal yields of 0.38 kg/kg dry feed are required to balance heat demand and supply (3).
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In commercial charcoal production, charcoal yields of 0.33-0.41 kg/kg dry feed are possible (4). However, the residence times for the solids used in commercial charcoal production, 12-18 hours (4), would result in a rather large and expensive pyrolysis reactor. Shorter residence times of the solids in the pyrolysis process are required to make the CASST process feasible. To determine whether the required charcoal yields can be achieved at shorter residence times, pyrolysis experiments with willow wood have been performed. The pyrolysis experiments have been conducted in a horizontal tube with an internal diameter of 4 inch. The feedstock is transported through the tube by a screw. The residence time of the solids can be varied by adjusting the rotation speed of the screw. An electrical furnace supplies the heat required for the pyrolysis process, and the reactor is continuously purged with Argon. The process conditions used for the pyrolysis experiments are given in table 1.
Table 1: Operating conditions for the pyrolysis experiments. Feedstock Particle size Particle feed rate Pyrolysis temperature Argon purge flow Residence time solids
Willow chips -8~3x1mm 1.95 kgAu 350,400 and 550°C 10 NYmin 10 min at 35OoC,7 min at 400 and 55OoC
Charcoal yields and the volatile matter content of the charcoal produced decrease with increasing temperature (see figure 2). For willow wood the required charcoal yield of 0.38 kgikg dry feed can be obtained at a temperature of 350-400°C, giving a product with a volatile matter content of 30-50 wt.%. In this temperature range 46-60% (LHVbasis) of the energy content of the feedstock is retained in the charcoal. +Charcoal
04
300
yield
+Volatiles in charcoal
I
I
400 500 600 Pyrolysis temperature ( "C)
Figure 2: Charcoal yield and volatile matter content of the charcoal as a function of the pyrolysis temperature
289
CHARCOAL GASIFICATION
The objective of the CASST process is to produce a gas with a tar content that is low enough to make a tar removal or conversion step unnecessary. The charcoal production experiments revealed that the charcoal in the CASST process will have a volatile matter content of 30-50 wt.%. Gasification experiments have been performed to determine the influence of the volatile matter content of the feedstock on the BTX (BTX= benzene, toluene and xylene) and tar concentrationsof the product gas. The gasification experimentshave been conducted in a labscale bubbling fluidised bed gasifier with an internal diameter of 74 mm and a bed height of 200-250 mm. A screw conveyor continuously feeds the feedstock. The bed material used in the experiments is silica with a diameter of 0.27 mm and the particle size of the feedstock is 0.7-2 mm. The products leaving the gasifier pass a cyclone and are sampled for analysis. After passing a high temperature filter at 400°C the sample is analysed for BTX by gas chromatography and for tars by the SPA-method (5). Experimental conditions used for the gasification experimentsare summarised in table 2. Table 2: Operating conditions for gasification experiments
Feedstock Volatiles (wt.%) Fuel feed rate (g/hr) Steam (&) Air (NVmin) Nitrogen (NYmin) Gasification temperature ("C)
Willow 81 334 550
4,l 1,0 870
50 409
550 491 190 873
Charcoal 31 512 550 5,o
13 555 550
691
130
130
860
846
Results of the gasification experiments are summarised in figure 3. The BTX as well as the tar concentrations in the he1 gas decrease with decreasing volatile matter content of the gasification feedstock. For a gasification feedstock with 30-50 wt.% volatiles the BTX-concentration is 4-12 times lower compared to wood gasification, and the tar concentration 10-50times lower. +BTX
0
+Tars
20 40 60 80 100 Volatiles in feedstock (wt.??)
Figure 3: Tar concentrations in fuel gas as a hction of the volatiles content of the
feedstock
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Besides by a lower tar concentration, charcoal gasification is also characterised by tars see table 3. The shift in tar composition is important since requirements by downstream processes not only comprise concentrations, but also the type of compounds present in the fbel. Generally, requirements for lowmolecular weight compounds are less stringent than for high molecular weight compounds.
with a lower molecular weight,
Table 3: Composition of tars analysed by the SPA-method
Feedstock Volatile matter in feedstock (wt.%) 2-ring 3-ring 4-ring
Willow 81 64% 27% 9%
Charcoal 50 31 87% 100% 12% 0% 1% 0%
ENERGETIC EVALUATION OF A CASST/COMBINED CYCLE SYSTEM The energetic performance of a CASST/combined cycle (CASSTKC) system has been compared with the performance of an air-blown gasificatiodcombmed cycle (ABG/CC) system. The ABG/CC system uses a dolomite cracker for tar conversion. The configuration of the CASST/CC system is conceptually similar to the ABG/CC system described by Faay et. al. (6) and given in figure 4.
Figure 4: Simplified flowsheet for a CASST-combined cycle system
29 1
The CASST/CC system has been modelled within the process flow simulation program ASPENPLUS. The major assumptions made for the CASST process are given in table 4. For the combined cycle the assumptions described by Faay et. al. (6) have been used. Table 4: Major assumptions for the energetic evaluation of the CASST-combined cycle system Pyrolysis temperature Gasification temperature Steam/carbon ratio gasifier Steam inlet temperature gasifier Combustion temperature
35OOC 850OC 2 moVmol 45OOC 1050°C
The performance of the CASST/CC system is summarised in table 5. Preliminary gasifier design showed that for a conventional design a carbon conversion of 80% is feasible for the steam gasification of charcoal. Use of an advanced design allows the carbon conversion to be increased to 95%. The relatively low carbon conversion in a conventional design is a result of the continuous removal of char 6om the gasifier by sand circulating between combustor and gasifier. An internal separation of char and sand, similar to the C 0 2 acceptor process for coal (lo), prevents the continuous removal of char 60m the gasifier in an advanced design, and results in higher carbon conversiohs. At a carbon conversion of 80%, the cold-gas efficiency of the CASST gasifier (59%) is considerably lower than for an air-blown gasifier (74%). Due to the lower cold-gas efficiency, the output of the gas turbine, and therefore the gross power production, is lower in the CASST-CC system. The power consumption for fuel gas compression, and hence the internal power consumption, is lower in the CASST-CC system as a result of the lower volume flow of fuel gas. The lower internal power consumption only partly compensates the lower gross power output, giving a net electrical efficiency for the CASST/CC system that is lower than for an ABG/CC system (39.0 vs. 41.7%). Increasing the carbon conversion to 95% almost eliminates the difference in the net electrical efficiency of both systems. The cold-gas efficiency of the CASST gasifier is still lower than for an air-blown gasifier due to the fact that in the CASST process part of the feedstock is processed at a higher temperature (1050°C) than in air-blown gasification (9OOOC). The lower gross power production is almost completely compensated by the lower internal power consumption.
292
Table 5: Energetic performance of a CASST-CC and an ABG/CC system CASST
CASST
80 Willow
95 Willow
Air-blown gasifier 97% Poplar
59
70
74
Thermal input Thermal input
73.9 90.7
73.9 90.7
72.0 89.6
Gross power output Gas turbine Steam turbine Total
20.7 12.7 33.4
24.3 11.0 35.3
24.7 12.3 37.0
Internal use Air compression combustion Fuel gas compressor Dryer Others Total
0.4 2.8 0.9 0.5 4.6
0.3 3.2 0.9 0.5 4.9
5.6 0.9 0.5 7.0
Net power output
28.8
30.4
30.0
39.0 (W%) 31.8
41.1 33.5
41.7 33.5
C-conversion gasifier Feedstock
(%)
Cold gas efficiency gasifier
Net electrical efficiency Net electrical efficiency
ECONOMICS OF A CASST-COMBINEDCYCLE SYSTEM
In this paragraph the capital investment required for a 30 MW, CASST-CC system is determined, and compared with the capital investment for a 30 MW, ABG/CC system. The capital cost estimate has been made for a CASST-CC system with a carbon conversion of 95% in the gasifier. The hardware costs for most components in the CASST-CC system have been estimated based on similar work for air-blown gasification (6). Costs of individual components have been scaled based on throughput, and the scaling hctors given in (6). The investment cost for the pyrolysis process have been estimated based on a preliminary design of the pyrolysis process and a cost estimate for a similar installation for coal pyrolysis (7). The investment costs for the combination of gasifier and combustor have been estimated based on a preliminary design of the gasifier and cost estimates for the Battelle gasification process (8). The hardware costs are summarised in table 6.
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Table 6: Investments costs (MECU) for a CASST-CC and air-blown gasifier/CC System.
CASST CASST Air-blow with fuel gas with flue gas gasifier cleanup cleanup 7.0 7.0 7.0
Pretreatment section Gasification section - Pyrolysis - Gasifier - Tar cracker - Cyclones - Total gasification section Gas processing section - Gas cooling flue gas - Baghouse filter flue gas - Condensing scrubber flue gas - Gas cooling fuel gas - Baghouse filter fuel gas - Condensing scrubber he1 gas - Fuel gas compressor - Total gas processing section' Combined cycle section - Gas turbine - HRSG - Steam turbine
- others - Total combined cycle
Total hardware costs (=HC) General plant facilities (33% of HC) Total direct costs (=n>c) Engineering fees (15% of TDC) Project contingency (10% of TDC) Building interest 1st year (25% of TDC*interest rate) Building interest 2nd year (75% of TDC*interest rate) Fedoverheaddprofits (10% of TDC) start-up costs (5% of TDC) Total investment Investments~ e(ECUkWJ r
294
0.7 7.8
0.7
8.5
8.5
2.5 2.3 1.8 6.6
2.8 0.0 0.0 1.2 0.7 1.1 0.8 6.5
2.8 1.5 2.5 1.2 0.7 1.1 0.8 10.5
0.0 0.0 0.0 2.1 1.1 1.8 1.8 6.9
11.2 2.2 2.9 0.9 17.2 39.2 12.9 52.2 7.8 5.2 1.3
11.2 2.2 2.9 0.9 17.2 43.2 14.3 57.5 8.6 5.7 1.4
12.3 2.4 3.1 0.9 18.7 39.1 12.9 52.1 7.8 5.2 1.3
3.9
4.3
3.9
5.2 2.6 78.3 2574.9
5.7 2.9 86.2 2835.3
5.2 2.6 78.1 2692.3
7.8
The total hardware costs for the CASST-CC system depend on the necessity for cleaning of the flue gas produced by combustion of the volatiles. When no flue gas cleanup is required, the total hardware costs of a CASST-CC system are slightly lower than for a system based on air-bIown gasification. The hardware costs for the gasification section are higher for the CASST process than for air-blown gasification. The reason for the higher costs is the rather large dimensions of the gasifier. Obviously the low volatile matter content of the gasification feedstock for the CASST process results in a high fixed carbon content of the gasification feedstock. Since the gasification of char is a relatively slow process, a large gasifier is required in order to achieve high carbon conversions. The higher costs for the gasification section are compensated by slightly lower costs for the gas processing section and the lower costs for the combined cycle section. The lower costs for the combined cycle are a result of the lower gross power production of the CASST-CC system (33.4 MW, for the CASST-CC vs. 37 MW, for the air-blown gasifier based system). Furthermore it has been assumed that the gas turbine in the CASST-CC system does not require modifications for the combustion of low calorific gas. When cleanup of the flue gas produced by the combustion of volatiles is required, the total hardware costs for a CASSTKC system are slightly higher than for an ABGKC. The necessity for flue gas cleanup depends on a number of factors: 1. The distribution of contaminants over volatiles and charcoal in the pyrolysis process; the major contaminantspresent in the feedstock for pyrolysis process are chlorine, nitrogen and alkalis. Jensen et.al. (9) showed that during pyrolysis of wheat straw considerable amounts of chlorine are released during the pyrolysis between 300400°C. Potassium was only released at temperatures above 600700°C. 2. The behaviour of ash in the CASST process; at a relatively low carbon conversion in the gasifier, it will be necessary to use the energy content of unconverted carbon in the combustion process resulting in flue gas cleanup. At high carbon conversions in the gasifier the amount of ash transported to the combustion process will be negligible, and flue gas cleanup will not be required To obtain the total investment costs, the hardware costs have to be increased with costs for instruments, control systems, engineering, building interest, etc. Generally, these costs are calculated as a percentage of the total hardware costs. The costs factors used in this study are based on the work of Faay et.al. (6). Since percentages are used to calculate the additional costs, the trends observed in total hardware costs are also observed in the total investment costs and the investments costs per kW,. For the CASSTKC system the total investment ranges i3om 2575-2835 ECUkW,, whereas for the air-blown gasifier/CC system the investments were 2692 ECUkW,.
295
CONCLUSIONS CASST is a new and advanced process for biomass gasification. The cold gas efficiency of the CASST process is maximised if heat supply by combustion and heat demand by gasification is balanced. A charcoal yield of 0.38 kgikg dry feed is required to balance heat demand and supply in the CASST process. The required charcoal yields are possible at short residence times for a pyrolysis temperature of 350-400°C. The charcoal produced at these conditions has a volatile matter content of 30-50 wt.%. Gasification of charcoal with 30-50 wt.% volatiles gives BTX concentrations that are 4- 12 times lower than for wood gasification and tar concentrations that are 10-50 times lower. Furthermore, the tars produced by charcoal gasification have a lower molecular weight than the products produced by wood gasification. At a carbon conversion of 80% in the gasification step, the performance of a CASSTICC system is lower than for air-blown gasificatiodcombined cycle system (39.0 vs. 41.7%). The lower gross power output of the CASST/CC system is only partly compensated by a lower internal power use. Increasing the carbon conversion to 95% almost eliminates the difference in the performance of both systems. The capital investment for a CASSTKC system depends on the necessity to cleanup the flue gas produced by combustion of the volatiles. If no flue gas cleaning is required the investment for a CASST/CC system is slightly lower than for an ABG/CC system. If flue gas cleaning is required, the required investment is somewhat above the investment for an ABG/CC system. FUTURE WORK
The project will continue with detailed design of individual components of the CASST process and a more detailed capital cost estimate of the individual components. The capital cost estimate will be used for an economic evaluation of the CASST process for fuel gas production as well as for synthesis gas production. Further development of the process depends on the results of the economic evaluation. ACKNOWLEDGEMENTS The work described in this paper has been performed as part of the ECN-ENGINE program. The work has been co-fbnded by the Netherlands Agency for Energy and the Environment (NOVEM). REFERENCES 1. T.A. Milne, N. Abatzoglou and RJ. Evans, Biomass Gasifier "Tars": Their Nature, Formation and Conversion, National Renewable Energy Laboratory, NREL/TP-570-25357,1998 2. A.V. Bridgwater, The technical and economic feasibility of biomass gasification for power generation, Fuel, 74(5), 1995,631-653 3. H. den Uil, Secondary energy carriers fiom biomass. ECN-Report ECN-CX-99002, Petten, 1999 4. H.G. Brocksiepe, Holzverkohlung, In: Ullmans Encyclopedia der Technischen Chemie, Verlag Chemie, Weinheim, New York, 1976 5. C. Brage, Q. Yu, G. Chen and K. Sjostr6m, Use of amino phase adsorbent for
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biomass tar sampling and separation, Fuel, 76(2), 1997, 137-142 \ 6. A. Faay, B. Meuleman and R van Ree, Long term perspectives of Biomass Integrated Gasification with Combined Cycle Technology, EWAB Report 9840, 1998 7. COMPCOALm - A profitable process for production of a stable high BTU fuel fiom powder river basin coal, DOE-MC/30126-5 102,1993 8. K.R. Craig and M.K. Mann,Cost and performance analysis of biomass integrated gasification combined-cycle (BIGCC) power systems, NREL/TP-430-21657, 1996 9. P.A. Jensen, K. Dam-Johansen and B. Sander, Pretreatment of straw by pyrolysis and char wash, Proceedings of the 2"dOUe Lindst6m Conference on renewable energy - biomass, 9- 1 1 June 1999, Stockholm, 1999 10. D. Hebden and H.J.F. Stroud, Coal gasification processes, In: Chemistry of coal utilization, Ed. M.A. Elliott, John Wiley & Sons Inc., New York, 1981
297
Development of a Novel, Reverse-Flow, Slagging Gasifier B. Van de Belda, D. AssinkaJ. Brammerb,A.V. Bridgwaterb a BTG Biomass Technology Group BV, PO Box 21 7, NL-75OOAE, Enschede, the Netherlands Bio-Energy Research Group, Aston University, Birmingham B4 7ET, UK
ABSTRACT A novel biomass slagging gasifier has been developed under the European JOULE programme. The main objective of the project was to develop an innovative, reverse flow molten slag fixed bed gasifier concept for combined heat and power generation in a capacity range of 0.5-2.5 Mw,, and with improved thermal efficiency (increase of 5 - 10 %) compared to conventional fixed bed gasifiers. The reactor enables handling of relatively wet (up to 40 % mcdb) and ash-rich biomass fuels with a low ash-melting point. The gasifier produces a gas with a caloric value of at least 6 MJ/Nm3.This paper presents some experimental results obtained with the pilot-plant gasifier (20 kg biomass throughput per hour). The technical feasibility of the reverse flow gasifier has been demonstrated. The gasifier enables the production of a tar-free gas at low temperature, slag granulates and medium pressure steam. INTRODUCTION In April 1997 an EC-project started aimed at the development of a novel type of gasifier. The starting point for this development was an analysis of conventional, airblown fixed-bed gasifiers. Important (techrucal andor economic) disadvantages of these gasifiers were identified like: Producer gas has relatively low caloric value (4 - 5 MJ/Nm3),which results in low heating values of the air-fuel mixture, and consequently in the application of large engines. Furthermore, gas turbine application might be even impossible. Engine control will be improved with higher caloric values, whereas a smaller engine size will reduce the capital costs. Large amounts of inert nitrogen is flowing through all equipment, which results in the use of equipment with a large volume; Reduced equipment size will result in reduced capital costs. Inert nitrogen is also responsible for part of the sensible heat losses. Significant amount of the energy potential of the biomass is lost as sensible heat. Installation of gas-gas heat exchangers is economically not attractive at small
298
scale. Due to low heat transfer coefficients in gas-gas heat exchangers very large equipment is required for sufficient heat recovery. Moreover, due to high temperatures and corrosive gases, special materials might be required. Conversion of wet biomass is problematic; Conversion of biomass containing ash with a low ash melting point is problematic. During the last decade, a lot of effort was put into producing (new types of) energy crops. Most of these energy crops (e.g. Miscanthus) contain ash with a low melting point and can probably only be used in conventional gasifiers if mixed up with large amounts of, for instance, wood. Stand-alone operation of conventional gasifiers may cause large problems (clinkering and blocking). Besides energy crops, waste is an interesting and suitable feedstock, since waste streams often have low ash melting points. Compared to conventional waste conversion systems, the new gasifier should acheve much higher thermal efficiencies; Producer gas often contains tar, which has to be removed prior to use in engines or other prime movers. To reduce the tar content additional downstream tar removal or conversion units are often required. Based on the above mentioned points an idea on a new type of gasifier was born. The novel gasifier is based on the so-called reverse-flow technology. A general description of reverse-flow technology is given elsewhere [l]. Main feature of the system is the high degree of heat integration due to the reverse flow principle, which would result in: Improved thermal efficiency ( 5 - 10%) due to lower producer gas outlet temperatures; Conversion of relative wet biomass is possible; Increased Lower Heating Value of producer gas. Due to reverse-flow operation higher operation temperatures inside the gasifier are expected, which, it is hoped, it would result in:
0
Tar free producer gas; due to the high temperatures all tars will be thermally cracked, and downstream gas cleaning with respect to tars is not required. Enables handling of biomass containing ash with a low ash melting point; Ash will be removed from the reactor as liquid (slag). Due to slagging operation no carbon will remain in the ash, which improves the overall efficiency, and lmproves the opportunities for utilizing the asWslag.
The project started with experimental work in a small lab-scale unit (biomass throughput 2 kg/hr). The operation principle and preliminary results are presented elsewhere [2, 3,5]. The main conclusions which could be drawn was that operation of a reverse-flow slagging gasifier is technical feasible with promising performance. Based on these promising results of the small unit a mini-plant (20 kg/hr) was designed and constructed. EXPERIMENTAL SET-UP In Fig. 1 a schematic drawing of the experimental installation is shown. Roughly, the 299
installation can be divided in four section being 1. Oxygen enrichment; 2. biomass feeding section; 3. the reverse-flow gasifier, (RFSG) and 4. gas cooling. c
CONTROL
AIR ENRICHMENT
REVERSE FLOW SLAGGING GASIFIER
COOLER
FLARE
Figure 1 Flowsheet of the experimental set-up The air enrichment unit has been designed and constructed by KEMA and is based on membrane separation technology. The unit is capable to deliver air containing a maximum oxygen content of 40 ~01%.Besides the application of the membrane unit for enriching the air also pure oxygen from bottles was used to enrich the air. A screw feeder is installed on the top of the gasifier to feed the biomass. The maximum feeding rate is about 60 k&. Feeding to the gasifier is carried out batchwise at the maximum feeding rate; the gasifier is a fixed-bed type and does not require an accurate continuous biomass feed. The reverse-flow slagging gasifier has a maximum design capacity of 20 kg of biomass per hr. Four pneumatic ball valves are installed to control the flow direction through the gasifier. The reactor consists of two packed beds, the actual (crossdraft) gasifier and a slag removal system. The gasifier is made from a castable, high temperature resistant material, and was constructed by KARA Energy Systems. Between the packed beds and the actual gasifier the option is available to add additional (enriched) air to ensure the production of a tar free gas. The secondary air injection is controlled by Kammer valves in combination with temperature controllers. The producer gas is cooled before it is vented to the atmosphere or flared off. The product gas temperatures is reduced in a water-cooled, shell and tube heat exchanger. The water circuit is a closed system; the water is cooled in an external cooling tower. The process is controlled by a computer (VisDaq), and the mini-plant is to a large extent automated. To support the operation of the installation 5 video cameras have been installed at different location in the process. At several locations in the process gas, tar and water samples can be taken for 300
fiuther analysis. Gas is analyzed off-line by a gas-chromatograph. In one of the experiments a fast, on-line micro-GC was used to determine the gas composition, whch has the obvious advantage of small time intervals between measurements and obtaining the results very fast (within a few minutes). To determine tar content and composition in the producer gas the SPA method (developed by KTH, Sweden) has been applied. The final determination of tars were performed on a gas chromatograph equipped with a liquid autosampler.
RESULTS GENERAL Before the actual experiment is started both packed beds are preheated simultaneously by burning propane between the gasifier and packed beds. Complete preheating of the beds to about 1000 "C takes a few hours. At start-up the bottom part of the gasifier is filled with charcoal and on top of it a layer of biomass is placed. The gasification process is started by feeding cold, (enriched) air to one of the packed beds; the air leaving the packed bed does reach a temperature of about 900-1000 "C, whch is more than sufficient to start the gasification process. Subsequently the hot producer gas flows through the second packed bed, and heating it up W e r . Typically, the flow direction is reversed every 10 - 15 minutes. As a base case the air flow is 15 Nm3/hr, which results in a biomass consumption rate of about 15 kg/hr. In the top section of the gasifier a level indicator is installed, which gives an electronic signal when the biomass level becomes too low. At that moment a fixed amount of biomass will be screwed into the reactor (batchwise addition). During an experiment a number of temperatures are monitored, and gas samples are taken for analysis.
Balancing heat front velocities In a reverse-flow reactor the hot packed bed is cooled by the cold feed gas, whereas the hot producer gas is used to reheat the packed bed. The process of reheating and cooling of the packed beds should occur in approximately the same time to enable a 'steadystate' operation. A big difference with 'proven' reverse flow systems is the significant increase in gas flow due to gas production in the gasifier. For normal air - 21 vol% oxygen - the increase is about a factor 2, for enriched air the factor is even higher. As a consequence re-heating of the packed beds by hot producer gas is a much faster process than cooling the packed beds by incoming air. The net result will be a movement of the temperature profiles towards the reactor outlet. For a typical experiment it was observed that after about 1% hr the maximum outlet temperature is approximately 120 "C; after 5 hr the maximum outlet temperature was already increased to about 700 "C. Such high outlet temperatures give rise to operational problems, and limited the duration of the experiments. At the in- and outlet of the reactor a connection is made between refractory material and the stainless steel of the vessel. These materials exhibit a large difference in thermal expansion coefficient, which resulted in undesired internal gas flows. Obviously, this problem needs to be solved for long-term proper operation of the gasifier. Basically, the heat front velocities need to be approximately the same in both packed beds, whch leads to the following criteria;
30 1
ah, bed I =
v producer gas, bed 2
This equation can be written as (see for its deduction [6]):
In the temperature range considered the volumetric heat capacity is for both air and producer gas of the same order of magnitude, and the previous equation can be simplified to: air, bedl
-e
producer gas, be&
From this expression we can conclude that stable operation can be obtained when the (average) gas flows through both packed bed are approximately the same. Different process options have been evaluated to fulfill Eq. 3. The solution chosen is illustrated in Fig. 2. About half of the producer gas is removed from the gasifier via the slag removal system, the other half is used to reheat one of the packed beds. The advantages of this way of operation are 0.a. flow and temperature of gas withdrawn is insensitive for flow reversal, and conventional steam generation is possible, and, moreover an additional burner system for the slag removal unit is no longer required.
Figure2 Hot producer gas withdrawal via the slag removal system
Fig. 3 shows the effect of using the new system on the outlet temperature of the packed bed. For this experiment the reactor was started as usual (no gas withdrawal), after 6 cycles, gas withdrawal via the slag tap was started. It was observed that the maximum outlet temperature (at the end of a cycle) decreases to an acceptable level of about 75 OC.This decrease in maximum outlet temperature indicates that cooling of the bed is now faster than reheating of the beds. Fig. 4 shows the measured gas flows through the reactor. The inlet air flow is 15 Nm3/hr, and the total producer gas flow is about 30 Nm3/hr. The goal was to withdraw 50% of the flow via the slag tap, and to use the other 50% of the flow to reheat the 302
packed bed. For this experiment relatively too much producer gas was withdrawn via the slagtap, which is in agreement with the temperature profile shown in Fig. 3. However, by monitoring the temperatures in both packed beds the flows can be easily controlled to obtain optimum and steady-state performance. It should be realized that it is not necessary that at every moment the heat front velocities are equal, but average over a long period of time (which is due to the large heat sink, resulting in very slow reactor dynamics).
-250
E
F 3 200 2 al
Q
5 150 5al
3 100 -
0
0
5
10
15
20
CyCb
25
30
35
40
number[-]
Figure 3 Gasifier outlet temperature as a function of the cycle number
30
*
20
I
L
.K . cc)
E
E. 3
2 10 0 . 0
Producergar vat packed bed
I
I00
200 Runtime [min]
390
400
Figure 4 Different gas flows to and from the gasifier as a function of time
Experimental results
The operational problems related to undesired internal gas flows and the difference in heat front velocities limited the execution of a systematic experimental program. In Table 1 typical gas composition are given for some selected experiments; the number 303
of samples were limited for most experiments, since only an off-line GC was available. For the experiment 11 an fast, on-line micro GC could be used resulting in more reliable results. In Fig. 5 the CO, Hz, COz and CH4concentration in the producer gas are shown as a fbnction of time for experiment 11. Sampling started after about 290 min of operation. After a couple of minutes the real producer gas is sampled and actual concentrations are monitored. Experiment 11 was the first experiment where part of the producer gas was withdrawn via the slag tap. The reduction zone in the reactor might be too small to obtain full reduction of the gases to i.e. carbon monoxide and hydrogen. However, the reactor was not designed for the operation mode used in the last experiments. In particular the gas removed via the slag tap is rather close to the oxidation zone. This fact was confirmed by later experiments (see e.g. exp. 12) where separate gas samples were taken from both exit gas flows, see Table 1. Gas withdrawn from the slag tap contained a relatively high amount of carbon dioxide (and probably more water), and the LHV was much lower than the producer gas obtained at the exit of the gasifier.
25 20
0 250
350
300
400
Runtime [min]
Figure 5 Gas composition as a function of time At the end of the experiment the oxygen content was increased fiom 21 to 30 vol%, which resulted -as expected- in lower nitrogen contents in the producer gas, whereas the oxygen content remained very low. A surprising result was that after the increase in oxygen concentration mainly an increase in CO content was observed, whereas the hydrogen content remained nearly the same. Tar content and composition was determined by using the SPA method. Tar in the producer gas mainly consisted of lighter tars (as naphthalene and phenol), and the absolute concentration was rather low, but it is questionable whether these concentration are low enough for engine application. Heavier tars were not found in the product gas. For some experiments the gasifier cold-gas efficiency was determined, which varied between 65 and 82%. However, these figures are not very accurate.
304
Table 1: Results from a number of selected experiments (for experiment 12; a :producer gas from packed bed; b. Producer gas withdrawn via slag tap)
Excess heat produced in the RFSG can be utilized for biomass drymg. In the project specific attention was paid to thls item, because it was expected that the drymg might be influenced by the periodic operation of the gasifier. In the final mode of operation (hot gas removal via the slag tap) this particular problem of the RFSG was completely avoided. In principal, a hot gas is now available at a constant flow and temperature, and a conventional steam generator can be adopted.
DISCUSSION AND CONCLUSIONS The project was aimed at the development of an innovative gasifier based on the reverse-flow principle. It should result in an improvement in thermal efficiency, and enabling the conversion of biomass with high ash contents and low ash-melting point. Based on promising results obtained in a lab-scale unit the design was made of a mini-plant with a capacity of 20 kg of biomass per hour. The mini-plant included air enrichment by membrane separation technology. Experimental work started in second half of 1999, and a lot of effort had to be put in to the improvement of reactor construction and operability. To a large extend these initial problems have been solved, and a number of experiments have been carried out. Besides the development of the new gasifier also a technical and economic evaluation of the total process was carried out including biomass drymg, air enrichment, the new gasifier and application of the gas in a gas engine. The main conclusions from this evaluation were: Application of the reverse-flow principle will lead to an increase in thermal efficiency. Compared to conventional downdraft gasifiers the economics of the process are improved, provided that ambient air is used as gasifjmg medium
305
0
Application of air enrichment will lead to higher heating values of the producer gas, but will not improve the overall efficiency; from an economic point of view air enrichment is not attractive.
Within the project significant progress has been made with respect to the design and operation of a RFSG. Currently, the system has not yet reached a “market-introductionstage”, and additional work is required. NOTATION Cp V
e
heat capacity heat front velocity flow
[ J k KI [dsl [m3/s]
ACKNOWLEDGMENT This research is h d e d in part by the European commission in the framework of the Joule3 R&D programme (contract JOR3-CT97-0130). REFERENCES Beld, L. Van de, Wagenaar, B.M., Prins, W., Cleaning of hot producer gas in a catalytic adiabatic packed bed reactor with periodic flow reversal, in Developments in thermochemical biomass conversion, eds A.V. Bridgwater and D.G.B. Boocock, Banff 1996,pp. 907 920. Beld, B. Van de, Bridgwater, A.V., Assink, D., Development Brammer, J.G., of a novel reverse-flow slagging gasifier for small-scale cogeneration applications, Proceedings 4* biomass conference of the Americas, 1999, pp. 1119-1126 Beld, L. Van de, Technical feasibility of a reverse-flow, slagging gasifier to improve thermal efficiency, Novem-report 99 13,1999 Stassen, Method for the gasification of biomass-comprising material and installation therefor, Patent application WO 99/42540, 1999 Gasification of biomass with oxygen rich air in a reverse-flow slagging gasifier, Final report JON-CT97-0130, June 2000 Beld, L. Van de ,Air purification by catalytic oxidation in an adiabatic packed bed reactor with periodic flow reversal, PhD thesis, University Twente, 1995.
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306
Study of Biomass Gasifier-Engine Systems with Integrated Drying for Combined Heat and Power J. G. Brammer, A. V. Bridgwater Bio-Energy Research Group, Department of Chemical Engineering and Applied Chemistry, Aston University, Birmingham B4 7ET, UK
ABSTRACT: This study addresses biomass gasifier-enginesystems for combined heat and power (CHP) in the size range 0.5-2.0 dry tomes per hour (roughly equivalent to 0.5-3.0 MW,). By the use of whole-plant modelling, it seeks to identifyefficiency and cost optima from a wide range of system configurationsand operating conditions. Particular emphasis is given to the integration of biomass drying, with the type of dryer and the source of heat considered in detail. The whole-plant model is spreadsheetbased, with individual worksheets corresponding to sub-models of system components. Most of these are explicitly modelled on a continuous basis, rather than being represented as fixed processes based on user-supplied data. Wherever possible, data supplied by manufacturers or taken from real systems is used in the construction of the submodels, particularly in the derivation of cost functions. In addition to the emphasis on drying integration which considers both batch and continuous systems, the study also considers the alternatives of slagging and non-slagging gasification (the former allowing the consideration of air oxygen enrichment), both with suitable provision for the delivery of a low-tar product gas. In the case of slagging gasification, the potential for supplying oxygen-enrichedair to the engine as well as the gasifier is also examined. INTRODUCTION Much work has been published on the application of whole-plant modelling techniques to gasification-based bbenergy systems to estimate performance parameters such as efficiency and cost of electricity. However, to date most of this work has focused on large-scale systems in the region 20-100 MW electrical output, utilising gas turbine technology often in combined-cycle mode. Commercial flow-sheeting codes such as AspenPlus are usually employed to give detailed performance predictions and component sizing at a few specific design points (typically below 10) for a few specific system configurations (typically below 5). Smaller system sizes are much less frequently considered, and very little in the way of system design analysis can be found for capacities below about 5 MW,, the point often considered to be that below which gas turbine based systems give way to those based on internal combustion engines as the prime mover on both efficiency and economic grounds. Nevertheless, a very large potential market exists for biomass
307
gasifier-engine systems at these smaller capacities, particularly embedded systems operating in CHP mode. This study uses wholaplant modelling to examine the performance and economics of biomass gasifier-engine CHP systems supplying electricity to the grid and hot water for district heating, in the size range 0.5-2.0 dry tonnes per hour (roughly equivalent to 0.5-3.0 MW,). A spreadsheet model has been constructed incorporating sub-models for all the major items of plant; these sub-models contain sufficient resolution of the processes to allow the eff‘ect of the principal flow variables to be studied on a continuous basis, as well as cost functions based where possible on actual manu~cturers’data. A choice of component types is offered for both the gasifier and the dryer, drying being a highly influential factor for the viability or otherwise of smallscale systems. The work is an extension of a task carried out under an EC-JOULE contract (JOW-CT97-0130), the primary objective of which was to develop a novel slagging gasifier for biomass feedstocks, aimed at CHP in the range 0.5-2.5 MW,. The gasifier, known as the Reverse-flow Slagging Gasifier (RFSG), operates with oxygen-enriched air to give a relatively high heating value gas at a high cold gas efficiency [I]. This is one of the two gasifier types available within the model.
ASPECTS TO BE STUDJED A generic biomass gasifier-engine system is shown in Figure 1. The study addresses various aspects of the performance and economics of biomass gasifier-engine systems.
Figure I Biomass gasifier-engine system The principal areas of investigation are the following:
GasiJierVpe Two gasifier options are available. The RFSG represents an attempt to advance gasifier performance at small scale through the use of high temperatures, heat integration and air oxygen enrichment to achieve a high cold gas efficiency and avoid the production of tars. A “conventional” non-slagging option is also offered, represented by an updraft concept coupled with thermal and catalytic tar crackers. Drying medium Three different drying media are considered, each represented by a different dryer system. These are high-temperature engine exhaust gas (rotary cascade dryer); engine exhaust gas plus combustion products ftom the direct burning of undersize biomass (rotary cascade dryer with integral biomass burner); and low-temperature air heated fkom the engine cooling system (deep-bed band conveyor dryer). 308
Biomass moisture before and after drying The absolute moisture content of the biomass as delivered before on-site drying, the absolute moisture content of the biomass supplied to the gasifier after on-site drying, and the difference between the two - i.e. the extent of drying (note: moisture content is quoted throughout this study on a dry basis - kg liquid per kg solid, %). Biomass dty input The effect of scale in the range 0.5-2.0 dry tonnes per hour (approximately 0.5-3.0 MW,), particularly on cost of electricity. Air oxygen enrichment For the RFSG which can operate using oxygen-enriched air, the effect of different air oxygen concentrations to the gasifier. Also, for configurations with the RFSG, the option to supply oxygen-enriched air to the IC engine to restore power lost by operating on low-CV gas (compared to natural gas). Biomass cost at the plant gate The price for biomass charged by the grower on delivery to the plant gate, already comminuted. A range of costs are considered to reflect the range of present and fhture estimates in the literature. The process submodels contain a sufficient level of resolution to allow all of the above areas to be examined. All parameters not referred to above are held constant throughout the study. A single “generic” biomass type is considered, the data actually corresponding to chipped short rotation poplar wood. MODEL ELEMENTS The model was created using the proprietary spreadsheet package Microsoft Excel together with the programming code Microsoft Visual Basic. It runs on a standard personal computer (200 MHz, 64 MB RAM). Individual worksheets are allocated to each system component sub-model; these are described in the following subsections. Non-capital cost items such as labour, maintenance and overheads are based on functions derived by Toft [2]. All capital costs are corrected where necessary to a total plant cost basis using methods defined in the same work.
DRYER Two of the dryer system options incorporate a rotary cascade dryer (Figure 2). The performance is modelled using a straightforward mass-energy balance, where biomass inlet and target outlet moisture are input values, along with biomass flow rate. For a given gas inlet composition, the model then calculates the necessary gas inlet temperature for a specified flow rate. Gas outlet temperature has to be specified. For the rotary cascade dryer with integral burner, the combustion temperature is calculated at a fixed air-fuel ratio using an iterative heat balance. The burner fuel comprises the fi-action of biomass feed to the dryer which is undersize and therefore unusable by the gasifier. All ofthe surplus is burned and the hot gases divided between the dryer and the engine exhaust water heater so as to maximise hot water production, subject to meeting the dryer duty.
309
Moist gas out
Wet material in
Lifting flights
\
4
+ -. Hot gas in
Figure 2 Rotary cascade dryer
-
Moist gas out f Wet material in
material out
Figure 3 Band conveyor dryer Dryer cost data were obtained 60m manufacturers as a function of dryer volume, the latter being calculated by the model assuming the volumetric heat transfer coefficient to be a function of mean biomass moisture content. The form of the function was derived 60m manufilcturers’data. In the case of the deep-bed band conveyor dryer operating on wrm air (Figure 3), the gas exit temperature fkom the dryer is much more critical to the overall performance and must be calculated for each case rather than specified as constant. A simple massenergy balance will not therefore suffice. Instead, a program for bed drying of wood chips developed at the Silsoe research centre in the UK was utilised. This calculates the drying process within the bed as a function of time, bed position and inlet conditions. Rather than incorporate the program into the wholeplant model, a set of polynomial functions was derived from the program results which adequately represents the process over the range of interest. Dryer cost data were obtained from manufacturers as a function of dryer band area, which is one of the parameters calculated by the Silsoe program.
GASIFIER The RFSG (Figure 4) is modelled as an equilibrium reactor operating at the carbon boundary (i.e. the point at which all free carbon is just consumed), after the method of Baron, Porter and Hammond [3]. The equilibrium assumption is justified on the grounds that the high-temperatureoperating regime of the RFSG is intended precisely to approach equilibrium conditions and avoid the production of tars. Insufficient experimental data were produced under the EC-JOULE contract to confirm the extent
310
to which this was achieved in practice. To allow for non-idealities, an “equilibrium approach” of 50K is assumed - that is to say, for a given set of reactants, the product gas composition is the equilibrium composition, but the product gas temperature at exit is 50K above the equilibrium reaction temperature.
Packed bed regenerative heat exchangers
Gasification reactor
Oxidant in (product gas out)
Slag quench
Product gas out (oxidant in)
Figure 4 RFSG concept The RFSG has a pair of regenerative packed-bed heat exchangersat the air inlet and product gas outlet, with the gas flow reversing typically every 20 minutes. The heat exchangers are modelled using a one-dimensional finite difkrence technique, after Schmidt and Willmott [4]. The “conventional” gasifier is an updraft gasifier with external tar cracking (UGETC), shown in Figure 5 . Taken as a whole (i.e. including tar cracking which is assumed to reduce tars to insignificant levels), the process is again treated as an equilibrium reactor operating at the carbon boundary with an equilibrium approach of SOK. This yields a final gas temperature of approximately 600-70OoC, which is in good agreement with real systems of this design [5]. There is no option for air preheating or oxygen enrichment - this would lead to slag formation for which the gasifier is not designed.
CRACKING UPDRAFT GASIFIER
Figure 5 UGETC concept 311
CRACKING REACTOR
As with the band conveyor dryer submodel, rather than incorporate the full equilibrium and finite difference models into the wholeplant model, 'a set of polynomial functions was derived fiom the models which adequately represent the processes over the range of interest. Capital cost was based on a published correlation by Bridgwater [6], updated to 1998 prices: 0.698
TPC ci = 7627 m g where TPCG is gasifier total plant cost (€,OOOs) and mg is biomass dry feed rate (kg/s). The scope of this cost correlation is gasifier plus tar cracking, but without final gas clean up. This is taken to apply to both gasifier types. In the case of the UGETC, the correlation agrees well with data obtained commercially. In the case of the RFSG, the manufacturer of the pilot-scale installation built as part of contract JOR3-CT97-0130 has estimated that the capital cost of a full-scale RFSG will be similar to that of the UGETC, the added cost associated with slagging and reversing operation being broadly balanced by the lack of need fw external tar cracking. IC ENGINE
The IC engine is defined as a spark ignition gas engine. The sub-model calculates combustion product composition and flow-rate assuming complete combustion at an equivalence ratio (air-to-he1 basis) of 1.7, typical of modem low-NOx spark ignition designs in which lean combustion is employed. Engine brake thermal efficiency, energy balance between cooling water, exhaust gas and other losses, and engine capital cost are all calculated using correlations derived fkom data supplied by the Jenbacher company of Austria, a leading manuhcturer of spark ignition gas engines for low-CV gas. AIR OXYGEN ENRICHMENT PLANT
Flow rates to and fiom the air oxygen enrichment plant are calculated by mass balances. The plant is assumed to be a pressure-swing adsorption plant up to around 20 tonnes per day oxygen, and a vacuum swing adsorption plant above that level. The plant delivers oxygen at 90% purity, which is subsequently diluted as required. The plant is assumed to be leased. Lease charge and power consumption data were obtained fkom a leading supplier as a function of oxygen demand. HEAT RECOVERY FOR CHP The CHP case for the present analysis is the supply of hot water at 100°C and about 2 barg pressure to a district heating scheme. Heat can be recovered fkom three sources in the system to meet this requirement. These are the hot product gas afler the gasifier and before the final gas clean-up, the hot engine exhaust gas, and the engine coolant circuits. Counter-flow shell-and-tube type heat exchangers are assumed. For the engine coolant heat exchanger (liquid-to-liquid) a pinch temperature of 5°C is specified. For the two gas-to-liquid exchangers a hot fluid outlet to cold fluid outlet delta of 20" (min.) is specified. Flows are then calculated fkom mass-energy balances. Capital cost is from a published correlation with heat transfer area [7], which is estimated by assuming suitable heat transfer coefficients and counter-flow fkctors. 3 12
BALANCE OF PLANT Four other items of plant are modelled which require minimal or no process calculations. Biomass reception, handling and storage is assumed to have no effect on biomass feed rate; capital cost is based on estimates by Bridgwater [6]. A product gas quench scrubber system is included for both gasifier types; capital cost is again based on estimates by Bridgwater [6]. The quench scrubber is a recirculating type, so that only the water content of the product gas is discharged to the foul sewer. Charges are based on a water treatment company tarif€. For systems with the band conveyor dryer, the IC engine coolant radiator supplies warm air to the dryer; flow rate i s calculated ii-om a mass-energy balance assuming an air exit temperature of 50°C, and capital cost is included in that of the IC engine (a radiator will always be installed). Finally grid connection costs are taken from Toft [2].
MODEL OPERATION
For each case to be calculated, a number of variable input parameters are set which enable the model to perform the calculation. A large number of cases may be specified to run sequentially without W h e r user input. The variable input parameters specified for the present study were: Biomass d v feed rate to gasifer Biomass moisture to dryer (dry basis) Type of dryer Air oxygen concentration to gasifier
Biomass cost at plant gate Biomass moisture to gasifier (dry basis) Type of gasger Enriched air supply to IC engine?
In order to investigate the various defined areas of interest, a comprehensive matrix of cases was calculated. This is shown in Table I. Table 2 gives some key fixed parameters for all runs. Dry tonnes (zero moisture) are denoted as dt.
Table 1 Matrix of modelling cases band conveyor, rotary cascade, rotarv cascade + burner
dryer type
RFSG, UGETC
gasifier type
ambient air, oxygen-enriched air
IC engine configuration
I biomass input biomass cost at plant gate biomass moisture (delivered)
I
dth €/dt
I
Y O
I biomass moisture (after drying) I % I I air oxygen concentration I % vol. I
0.5, 1.0, 1 5 2 . 0 30,75,120
I
50,75, 100
10,20,35,50 2 1,28,40,60
I
I
The key model outputs are overall efficiency (qo, the net electricity supplied to grid annually in h4.I plus heat exported annually in MJ, all divided by the energy content of biomass delivered to plant annually in MJ on a lower heating value basis, expressed as 3 13
a percentage) and cost of electricity (CoE, the total annual operating cost less the revenue for the exported heat divided by the electricity supplied annually to the grid, expressed in cents per kilowatt hour or c/kWh).
Table 2 Key fixed parameters
I Biomass LHV (d. a. E)
I BC dryer air temp. in
1
I 18.6 MJkg
I
1
I RC dryer gas temp. out I I RC dryer max. gas temp. in I product gas temperature at engine inlet
10
,50"C
Heat recovery min. gas outlet temp.
120°C
I IOOC
CHP water delivery temp.
100°C
300°C
CHP water return temp.
50°C
4;~
Parasitic power demand (non-specific)
5%
Annual capacity factor
0.8
No. of engines per plant Generator electrical ef&.
Plant replication number
'97%
Annual availability
The study assumes that all cases have a drying stage. Therefore the cases with 50% biomass moisture content both to the dryer and to the gasifier (implying no requirement for a dryer) are omitted fiom the analysis. For the purposes of this study, the revenue from the supply of heat to the customer is calculated assuming that the customer will pay a price equivalent to the cost of meeting the heat load by installing a new natural gas boiler, less the cost of installing any piping infiastructure to bring the heat fiom the CHP plant to the customer's premises, less a discount of 10%. It should be noted however that heat may be available at prices below that fiom a new natural gas boiler (e.g. existing heat networks in Scandinavia); this would depend on the specific location. RESULTS
SUPPLY OF OXYGEN-ENRICHED AIR TO IC ENGINE In all cases, provided the oxygen enrichment unit type is the same, qo when supplying oxygen-enriched air to the engine is lower than when supplying ambient air, and CoE is higher. This is because when supplying oxygen-enriched air to the engine, the marginal improvement in engine power output is always much less than the marginal increase in electricity consumption of the oxygen enrichment unit. Also, the marginal increase in lease cost of the oxygen enrichment unit far outweighs the marginal reduction in engine annual cost of capital. To illustrate, a 30 kW, improvement in engine output is obtained from an increase in oxygen production of 6 tonnes per day, which requires 115 kW,. The annual plant lease charge rises by €35,000 whereas the engine annual cost of capital fills by just €4,500. Supply of oxygen-enriched air to the engine is not therefore justified, and is not considered further.
314
SELECTION OF DRYER For both the band dryer and the rotary dryer with burner, there is sufficient heat available fkom the system to achieve all specified drying duties. For the rotary dryer without burner, there is insufficient heat available where the initial moisture content is 100% and the desired moisture content is 20% or less. Systems with the rotary cascade dryer with integral burner give a much higher qo than either of the other two systems in all cases (by about 13% on average), because of the utilisation of the undersize biomass which results in additional heat export. Figure 6 shows average CoE over all pre- and post- dryer moisture contents plotted against biomass feed rate for each of the three dryer options. CoE is expressed relative to the average for all dryers at that feed rate.
Figure 6 CoE for different dryer types Band dryer systems give the lowest CoE at biomass feed rates up to about 1.2 dtJh, with systems incorporating the rotary dryer with burner giving the lowest CoE at higher feed rates. The band dryer has lower electrical power requirements and lower capital cost, and also utilises lower-grade heat than the rotary dryers. However, the increased revenue from heat associated with the rotary dryer with burner becomes increasingly important as economies of scale reduce the influence of capital cost. With the assumptions made for heat revenue, it is likely that the added complexity and capital cost of the rotary dryer with burner would result in the band dryer being selected up to at least 1.5 dt/h. However, the relative economics of systems incorporating the rotary dryer with burner would improve if greater revenue could be obtained for the heat than assumed here, for example through subsidy.
AIR OXYGEN CONCENTRATION TO GASIFIER Figure 7 shows CoE plotted against air oxygen concentration to the gasifier, for a system with an RFSG and a band conveyor dryer. The biomass feed rate is 1.5 dth, the pre-dryer biomass moisture content 75% and the biomass cost QO/dt. In all cases, operation without air oxygen enrichment leads to the lowest CoE. Any increase in air oxygen enrichment leads to an increase in CoE. The trends are repeated for systems with the band conveyor dryer, and at other biomass feed rates and costs.
315
This occurs because as oxygen enrichment is employed, the electricity generated increases, but the increase is outweighed by the electricity demand of the oxygen enrichment plant. At the same time, the oxygen enrichment plant lease charge far exceeds the small reduction in IC engine capital cost. Finally less heat is available for recovery from the engine cooling system due to improved engine brake thermal efficiency.
Figure 7 CoE versus O2concentration It would therefore be recommended not to supply the RFSG with oxygen-enriched air. However, should oxygen enrichment be necessary to ensure slagging operation (beyond the scope of this model to determine), then oxygen concentration should be kept to a minimum.
SELECTION OF GASIFIER In Figures 8 and 9, the RFSG at various air oxygen concentrationsis compared with the UGETC, for a system with a rotary cascade dryer with integral burner, at a biomass feed rate of 1.5 dt/h and a biomass cost of€30/dt.
Figure 8 qo for two gasifier types
316
The UGETC systems have a greater VO, by around I%, due to larger recovery of heat (mainly fiom the hotter product gas) outweighing reduced electricity output. This might be expected, as the RFSG aims to retain sensible heat as far as possible within the gasifier. However, CoE is lower for the RFSG system operating on unenriched air by 0.6-0.8 ckwh, reflecting the much lower value of heat compared to electricity.
Figure 9 CoE for two gasifier types Similar results are seen at other feed rates, pre-dryer biomass moisture contents and biomass costs. If the RFSG has to operate with oxygen-enriched air however, its superiority in CoE is rapidly eroded. For an oxygen concentration of 0.28 the difference in CoE is negligible; for an oxygen concentration above 0.28 the UGETC is usually superior. BIOMASS MOISTURE CONTENT TO DRYER Any increase in biomass content to the dryer results in a reduction in qOand a rise in CoE (other parameters remaining unchanged), for the obvious reason that more heat is required for drying and is therefore unavailable for export, leading to a fall in revenue and an increase in dryer capital cost.
Figure I0 vo vs. moisture to dryer
3 17
This effect is illustrated in Figures 10 and 11, for systems with an unenriched air RFSG and a UGETC. In each case, the system has a rotary cascade dryer with integral burner, a biomass feed rate of 1.5 dt/h, a biomass cost of €30/dt and a moisture content to the gasifier of 10%.
Figure 11 CoE vs. moisture to dryer BJOIU~SS MOISTURE CONTENT TO GASIFIER
With one or two exceptions, an increase in biomass moisture content to the gasifier results in an increase in qo. This is despite a fall in electrical output, which is outweighed by increased hot water production. The increase in hot water production comes fiom the product gas water heater, where a fall in gasifier cold gas efficiency gives rise to more sensible heat in the product gas. CoE however always rises with increasing biomass moisture content to the gasifier, as the net benefits of a smaller dryer and increased heat revenue are outweighed by reduced electrical output (see Figure 7). BIOMASS FEED RATE Figure 12 shows the effect of biomass feed rate on CoE for systems with an unenriched air RFSG and a UGETC. The system has a rotary cascade dryer with integral burner, a biomass moisture content to and from the dryer of 75% and lo%, and a biomass cost of €30/dt. The effect of feed rate on qo is minimal. The effect of feed rate on CoE on the other hand is major, with CoE falling sharply as feed rate increases, particularly at low f k d rates. A quadrupling of feed rate results in a halving of CoE. The trends are repeated at other biomass costs and initial moisture contents, and also for systems with the band conveyor dryer. Clearly at small scales such as those investigated here, there are substantial benefits in building as large as possible.
318
Figure 12 CoE vs. biomass feed rate
BIOMASS COST
The effect of the cost of biomass at the plant gate on CoE is illustrated in Figure I3 for systems with an RFSG with unenriched air and a UGETC. In each case, the system has a band conveyor dryer and a biomass feed rate of 1.5 dth, with biomass moisture content to and from the dryer 75% and 10%.
Figure 13 CoE vs. biomass cost The expected increase in CoE with biomass cost is clear. For every €lO/dt increase in biomass cost, CoE increases by about 0.8 c/kWh (RFSGwith unenriched air) and 0.9 ckWh (UGETC). These rates of increase are kirly consistent regardless of scale or biomass moisture content.
GENERAL Taking a biomass moisture content to the dryer of 75% as typical, and the largest size of plant considered (2.0 dth), the maximum qo obtained is 79.7%, for a system incorporating a UGETC and a rotary cascade dryer with integral burner, drying the 3 19
biomass to 50% moisture content. The system exports 2.6 MW, of electricity and 6.0 M W m of heat. The maximum qo for systems with an unenriched-air P S G is 78.7% with the same dryer and final moisture content. The system exports 2.8 MW, of electricity and 5.6 MWth of heat. The minimum CoE obtained is 8.36 c k w h , for a system incorporating an RFSG with unenriched air and a rotary cascade dryer with integral burner, drying the biomass to 10% moisture content. The system exports 3.1 MW, of electricity and 5.2 MW,h of heat. The minimum CoE obtained for systems with the UGETC is 8.93 c/kWh, with the same dryer and final moisture content. The system exports 2.8 MW, of electricity and 5.5 MW,h of heat. Table 3 gives some data on heat-to-power ratios. Systems with the rotary cascade dryer with integral burner give higher heat-ta-power ratios than those with the band conveyor dryer because of the use of the undersize biomass to provide heat. Systems with the UGETC give higher heat-ta-power ratios than those with the RFSG due to the former’s lower cold gas efficiency. Table 3 Heat-to-power ratios Band conveyor dryer
Rotary cascade dryer with integral burner
average RFSG (unenriched air)
maximum minimum
I
I
UGETC
2.4
2.0 I
I
average
I .6
2.1
maximum
1.2
1.7
minimum
2.1
2.6
I
CONCLUSIONS For systems with an oxygen enrichment plant to supply enriched air to the gasifier, additional supply of oxygen-enriched air to the IC engine to boost output causes a reduction in qo and an increase in CoE in all cases, and is not therefore justified. Systems with a rotary cascade dryer with integral burner give the highest qo, and also the lowest CoE at higher feed rates within the range of study. At lower feed rates, systems with the band conveyor dryer give the lowest CoE. Oxygen enrichment of the air supply to the RFSG results in reduced qo and increased CoE in all cases, and is not therefore justified unless the RFSG requires oxygen-enriched air for operational reasons. In the latter case, oxygen concentration should be minimised. Systems with the UGETC give a higher qo but also a higher CoE than those with the unenriched-air RFSG in all cases. If the RFSG required oxygen-enriched air
320
however, the UGETC systems would deliver equal or lower CoEs and would be recommended. As biomass moisture content delivered to the dryer rises, lofalls and CoE rises. As biomass moisture content delivered to the gasifier rises, qo (generally) and CoE both rise. Moisture content delivered to the plant should be minimised, but provided a drying stage is required the drying carried out should be maximised at the expense of heat exported. Increasing biomass feed rate has little effect on qO, but produces substantial reductions in CoE. The plant should be built at as large a scale as practical.
For every €lO/dt increase in biomass cost, CoE increases by about 0.8-0.9 c k w h .
For a biomass moisture content to the dryer of 75% and a feed rate of 2.0 dth, the maximum qo obtained is 79.7% (UGETC). The minimum CoEs obtained are 8.36 c k w h (RFSG) and 8.93 &Wh (UGETC). Typical heat-to-power ratios are 1.5 for systems with a band conveyor dryer, and 2.0 for systems with a rotary cascade dryer with integral burner.
REFERENCES 1. Brarnrner J. G., van de Beld L., Bridgwater A. V., Assink D. (1999) Development
2. 3.
4. 5.
6. 7.
of a Novel Reverse-Flow Slagging Gasifier for Small-Scale Cogeneration Applications. In: Proc. 4th Biomass Conference of the Americas, pp. 1119-1 126. Toft A. J. (1 996) A Comparison of Integrated Biomass to Electricity Systems. PhD thesis, Aston University (UK) Baron R. E., Porter J. H., Hammond 0. H. (1976) Chemical Equilibria in CarbonHydrogen-Oxygen Systems. MIT Press. Schmidt F. W., Willmott A. J. (1981) Thermal Energv Storage and Regenerarion. Hemisphere. Bridgwater A. V., Maniatis K., Masson H. A. (1990) Gasification Technology. In: Commercial and Marketing Aspects of Gasijiers, (Ed. by Buekens, A. G., Bridgwater, A. V., Ferrero, G.-L., Maniatis, K.), pp. 41-65. Cornmission of the European Communities. Bridgwater A. V. (1991) Review of Thermochemical Biomass Conversion. Report B1202, Energy Technology Support Unit (UK). Garrett D. E. (1989) Chemical Engineering Economics. van Nostrand Reinhold.
ACKNOWLEDGEMENT The financial support of the European Commission (JOULE Programme contract JOR3-CT97-0130) is gratefiilly acknowledged.
32 I
Fuel-bound Nitrogen Conversion: Results from Gasification of Biomass in two Different Small Scale Fluidized Beds Magnus Berg', Peter Vriesman2, Eloise Heginuz2, Knster Sjostrom2 and Bengt-Goran Espenas' 1
TPS Termiska Processer AB, Studsvik, SE-611 82 Nykoping, Sweden Kungl Tekniska Hogskolan (KTH), Department of Chemical Engineering and Technology/Chemical Technology, SE-100 44 Stockholm, Sweden
ABSTRACT: The content of nitrogen compounds in gas produced by gasification of biomass results in a significant NOx formation during combustion in gas turbines. This is a major drawback for the IGCC concept. Different concepts have been proposed to decrease the concentrationsof NH3and HCN formed in gasification, e.g. the use of nickel-based catalysts downstream the gasifier. The objective of this project was to evaluate some possibilities to reduce the conversion of fuel-bound nitrogen into NH3 and HCN and to shift the selectivity towards molecular nitrogen by measures taken in the gasifier. Experimental results from two small, externally heated, atmospheric fluidized bed gasifiers (AFBG) were compared in this work. The influence from the position of the feeding point on the fuel-bound nitrogen conversion have previously been shown in one of these gasifiers, which can be fed from the top or directly into the fluidised bed. This gasifier (inner diameter 0.05m) is practically isothermal, which makes studies of reaction kinetics during pyrolysis and gasification feasible. Top feeding of Miscanthus resulted in lower conversion of fuel-bound nitrogen into NH3.Any influences on the conversion of fuel-bound nitrogen to HCN and tar were not found. A main conclusion is that the feeding point and therewith the contact between the fuel particles and the surrounding oxygen-containing gas is of great importance for the formation and destruction of N H 3 . These results were partly unexpected and originally only determined for one fuel wherefore the results now have been compared with results from another, slightly larger AFBG (inner diameter 0.2 m), both using the same fuel but in addition also for a second fuel. NH3was measured downstreamthe gasifier in all tests, using ordinary off-line sampling and subsequent analysis. The new experimental results on gasification of Miscanthus have been shown to correlate well with the previously published results from the smaller gasifier. A comparison of different methods to determine the fuel nitrogen concentration indicates that variations in fuel nitrogen conversion between different studies can probably to a large extent be directly related to the uncertainty in the fuel nitrogen analyses. 322
INTRODUCTION Advanced power generation systems based on Integrated Gasification Combined Cycle (IGCC) offer an attractive alternative for improved efficiency in power production fiom solid fuels. Commercial units have been established for coal and the technique is presently at the threshold to be demonstrated for biomass fuels. In addition to the improved thermal efficiency, these systems offer opportunities for improvements in efficient cleaning to avoid hazardous emissions, such as nitrogen oxides, sulphur, greenhouse gases, heavy metals, organics and particulates. There are many methods that can be suggested for reduced NOx formation in IGCC, and measures can be taken at different stages of the overall process. These methods can be categorised as follows: 1 2 3
4 5
Diminishing of the conversion of nitrogen bound in the fuel to nitrogen compounds in the gasifier. Hot gas cleaning by catalysed conversion of nitrogen compounds to Nz, Low temperature gas cleaning by wet scrubbing. Low-NOxburners in the combustion chamber of the gas turbine. Cleaning from NO, downstream of the gas turbine.
Typical nitrogen contents of solid fuels are 0.5 to 3% where the lower range is typical for wood and some agricultural fuels and specific coals are representative for the higher values. The nitrogen is released during gasification mainly as ammonia, hydrogen cyanide, organic compounds and molecular nitrogen. A small part of the fuel nitrogen remains in the solid char. For biomass gasification ammonia is in general the main carrier of bound nitrogen at the exit of the gasifier. The fraction of fuel nitrogen that is converted to N2 in the gasifier varies within a broad range, and fractions from a few percent up to almost 90% have been reported. The content of bound nitrogen in the gas depends most of all on the nitrogen content of the fuel. Leppllahti [I] concluded that the chemical hctionality of the fuel-nitrogen is important to explain yields of NH3 and HCN fiom different fuels and at different process conditions. He suggested that fuels such as peat and wood, containing amino type nitrogen, release nitrogen in primary reactions to a large extent to NH3 and to a lesser extent through secondary reactions to HCN liberated from heterocyclic ring structures, if compared to coal. According to Leppalahti, HCN is converted to NH3 in the presence of Hz by reactions inside the pores of char particles. The same group has also investigated the effect of pressure during peat and wood gasification and the results indicate that increased pressure favours the formation of ammonia as well as the total amount of bound nitrogen [2]. Berg and Espenas made experiments to investigate whether the chemical functionality of fuel nitrogen would influence the yield of NH3 in atmospheric fluidised bed (FB) gasification [3]. A pelletised wood fuel (10 mm diameter) that had been enriched in nitrogen content using two different nitrogen carriers, melamin (50% heterocyclic N) and urea (amino type N) respectively, was gasified using a bed of dolomite. They found very similar conversions to NH3 (about 60%), irrespective of the type of nitrogen carrier. However, only 30% of the “natural” nitrogen content of the wood fuel were converted to NH3 in reference tests, using no N-additive to the fuel. These results seem to be in line with the results referred to above. At the experimental 323
conditions used, the formation of NH3 from HCN (HCN expected to be formed from melamin in primary reactions) should be favoured both by the dolomite and by the rather large fuel particle size and the moderate heating rate of such large particles. The results indicate that the fraction of the additives that is converted to Nz did not depend on the type of nitrogen binding. In spite of the similar overall conversion to N H 3 obtained with artificially added nitrogen of very different chemical nature, a significantly lower conversion to NH3 was obtained from the “natural” fuel nitrogen. The use of dolomite and lime in the gasifier favours the formation of N H 3 , probably by catalysing the conversion of HCN to NH3 [I]. This can be an advantage for the overall process since it facilitates methods for downstream cleaning that are specific for NH3. The results available from the different large-scale pilot and demonstration projects indicate a rather large variation in the data on conversion of fuel nitrogen to ammonia. Within the VEGA project different wood based materials (both bark-free and with a high content of bark) with a nitrogen content between 0.1 and 0.4% were evaluated and the conversion was typically between 40 and 75% [4]. Results from the Vamamo plant [5] for the conversion of fuel bound nitrogen to ammonia show approximately 60% and HCN levels that does not contribute substantially to the overall conversion. Both these processes are pressurised. Also from the test campaigns in the atmospheric pilot unit (2MWh) performed at TPS, values in the same range have been measured. Tests on Brazilian wood chips with a nitrogen content of 0.12% resulted in 70 to 95% conversion [6] and tests on pelletised bagasse with a nitrogen content of 0.2% resulted in 75 to 95% conversion [7]. However it is clear from these different tests that as long as a dolomite bed is used the values of HCN will account only for a very low percentage of the total amount of fuel bound nitrogen. Also in fluidised beds the HCN is normally only a few percent of the total fixed nitrogen but in fixed bed gasifiers HCN may contribute to a larger fraction. Generally, at least for coal, an increase of the gasifier temperature will lower the yield of NH3, although the size of this effect can vary a lot [l, 2, 8, 91. A plausible explanation to different influence from temperature is different contribution of catalytic effects, e.g. by constituent of the ash. Addition of secondary air to the free-board in fluidised bed gasification seems to have no or only small effect on the conversion to NH3 and HCN. Leppalahti et aI. [ 101 found no effect from secondary air neither at atmospheric or pressurised conditions using peat and coal. Berg et al. [3] found only a minor decrease of NH3 by increase of the secondary-to-primary air ratio in FB gasification of pelletised wood. It should however be remembered that it is very difficult to adjust the primary to secondary air without at the same time changing the temperature distribution over the reactor and this effect could therefore partly disguise the effect of changing the primary to secondary air ratio in these investigations. Earlier studies comparing feeding at the top or at the bottom of a gasifier have shown differences in the product distribution [ll]. However, this reference makes a comparison between the feeder positions in different reactors, where other reactorspecific parameters also can mfiuence the results, and does not include any measurements of the ammonia concentration. Results from more recently published experiments [ 121 performed in a lab-scale AFBG where feeding from top and bottom is possible in the same reactor have shown that top feeding of Miscanthus results in lower conversion of fuel-bound nitrogen into 324
NH3. Influences on the conversion of fuel-bound nitrogen to HCN and tar were not found. A main conclusion is that the feeding point and therewith the contact between the fuel particles and the surrounding oxygen-containing gas is of great importance for the formation and destruction of NH3.
OBJECTIVE
The objective of this project was to evaluate some possibilities to reduce the conversion of fuel-bound nitrogen into NH3 and HCN and to shifl the selectivity towards molecular nitrogen by measures taken in the gasifier. In this paper the conversion of fuel-nitrogen to ammonia as a function of equivalence ratio and fuel has been investigated. Results from experiments in two different gasifiers are compared and discussed.
EXPERIMENTAL THE TWO GASIFIERS USED
Two different externally heated air-blown fluidised bed gasifiers, cf. figure 1 and 2, have been used for the experimental work. The smaller gasifier, having an inner diameter of 0.05m, can be controlled independently of the oxygen content by external heaters, which makes studies of reaction kinetics during pyrolysis as well as gasification conditions feasible. This rig also has the possibility to use feeding of the fuel either fi-om the top or ordinary bottom feeding into the fluidized bed. A detailed description of this gasifier is given in [ 121. Top feeding
*
External beating
___, Product gas
Bottom feeding
4'1
Ceramic filter
1
m Gas pre-heater
Figure 1 The air blown fluidised bed gasifier at KTH, inner diameter 0.05m.
325
In the larger gasifier, which has an inner diameter of 0 . 2 ~ 1the , external heating is dimensioned to compensate for the heat losses and thereby a fuel gas with a composition and heating value that resembles the fuel gas from a large-scale unit can be produced. However, when reducing the air to fuel ratio in this gasifier the temperature drops even when the air pre-heater and the external heaters are increased to their full power.
Figure 2 The airblown fluidised bed gasifier at TPS, inner diameter 0.2m The bed material in the smaller reactor was a-alumina with a narrow size distribution around 94 pm. In the larger gasifier a mixture between dolomite and Olivine 33 was used as bed material. The total weight of bed material was 22.1 kg out of which 15.8 kg were Olivin and 6.3 kg dolomite. This corresponds to a bed weight of 19.1 kg and 17% dolomite calculated as calcined dolomite. The hel-feeding rate for all the experiments in the larger gasifier was 8 kg/h and based on this and the fuel analysis the required stoichiometric air was calculated. For the highest h (supplied air /stoichiometric air demand) used in the gasification experiments (0.3) the temperature was maintained at approximately 800°C. This temperature could be maintained also at h=0.25 but at the two lowest air to fuel ratios the temperature dropped significantly, for h=O.16 the lowest recorded temperature was 744°C. During the evaluation of the influence of the air to fuel ratio on the fuel nitrogen conversion, the total volume flow have been kept constant by replacing some of the primary air with nitrogen, thereby maintaining the linear velocity through the fluidised bed. FUEL ANALYSIS Two different pelletized fuels were used, Miscanthus and sawdust. The fuel analysis for both fuels is shown in table 1. In the smaller gasifier the Miscanthus pellets had been ground and sieved to a fraction of 1-1.35 mm which then was used in the experiments. 326
The nitrogen content of t h s fraction was separately analysed and compared with the analysis of the whole pellets but the small difference in nitrogen content is probably within the accuracy of the analytical techniques. In the larger gasifier the pellets were directly fed into the gasifier. The nitrogen content was determined with two different methods, elemental analyses and the Kjeldahl method. The Kjeldahl method resulted in sipficantly lower values for both fuels. All the graphs with results are plotted using the higher nitrogen concentration without making any estimation on the relative accuracy of the two methods. Table 1
C H N N Ash
Method Elemental analyses Elemental analyses Elemental analyses Kjeldahl
Fuel analysis Miscanthus
Sawdust pellets
48.3
47.3 6.4 0.12
6.0 0.67 0.51 (0.58)* 2.7
0.08 0.4
*The higher value is for the sieved fraction used in the smaller gasifier
GAS ANALYSIS
In the gasifier the gas sampling was done downstream of the ceramic filters. All major gas components (H2, 02,NZ, CO, CO2, CH4, C2H4, C2H6) were measured using a Hewlett Packard 5890 Plus gas chromatograph. The gases were separated on Porapak Q column and Molecular Sieve 5A. The Molecular Sieve had a pre-column arrangement with a Porapak column. The pre-column arrangement was to limit the consumption of the Molecular Sieve by back blowing all the other gases except H2, 02,N2 and CO. From the Porapak Q column CH4, C02, C2H4 and C2H6were separated and H2, 02,N2 and CO resulted in one peak. The carrier gas was argon and a thermal conductivity detector was used. Air and nitrogen fed to the gasifier was measured using thermal mass flow controllers. Thls, together with the GC analysis of the nitrogen content of the fuel gas, was used to set up a nitrogen balance and thereby determine the fuel gas flow. The tar content was collected by condensation followed by two subsequent impinger bottles with acetone. Quantitative determination was performed by addition of hexadecane as internal standard to tlus acetone solution. Non-polar components in the acetone-tar solution were extracted to a dichloromethane solution and subsequently analysed using gas chromatography (FID-detector and separation on a 30m, 0.32 mm ID, fused silica DB-5 column with 0.25 pm stationary phase). The ammonia content was measured using a standard procedure with absorption in sulphuric acid solution (1 M) and subsequent analysis using ion selective electrodes (Orion EA 920 Ion analyser). The sampling time used for ammonia was approximately 15 minutes and the mean values from the gas analysis and the bed temperature was always calculated for the same period.
327
RESULTS In table 2 the experimental conditions and results are summarised. The experiments were performed in the order listed in the table and the bed material was replaced after experiment 4, i.e. new bed material when the fuel was changed. Between each experiment the conditions were allowed to stabilise on the new settings for at least 20 minutes after which three consecutive ammonia samples were taken. Since the first samples for some of the experimental settings deviated from the following two it was concluded that the time before taking the first sample was too short to reach stable conditions and therefore only the last two samples are included in the table and used in the W e r evaluation. It is important to notice that the temperature has not been totally stable during the experimental series, for the lowest air to fuel ratio the external heating was not enough to maintain the gasification temperature. The maximum temperature difference is however only approximately 50°C and the effect of temperature on the nitrogen conversion has in previous published work been found to be weak [1,2, 121. The experiments in the larger gasifier were performed to facilitate a direct comparison with the previously published results on Miscanthus from the smaller gasifier [12] but widened to also include a second fuel, pelletised sawdust. The conversion of fuel bound nitrogen to ammonia is shown for both these fuels in figure 3 and for comparison a line for the experimental data for bottom feeding of Miscanthus in the smaller gasifier is included. The experimental results from the two different gasifiers are in good correlation when based on the same fuel but the pelletised sawdust results in significantly lower conversion. However, it should be remembered that the conversion values are calculated using the higher results from the fuel-nitrogen analysis. Due to the large uncertainties in the fuel nitrogen concentration, the difference between the two fuels cannot be considered as significant. However, the systematic error from the fuel analysis will not disturb the evaluation of the influence from air to fuel ratio on the relative conversion for each fuel or the results from the two gasifiers. In figure 4 the conversion of carbon to tar is shown as a function of h. Since the char content of the bed material was not determined for each individual experiments it is not possible to make a complete carbon balance. Based on the gas analysis and the tar conversion a rest fraction, which could be regarded as the carbon conversion to char, was instead calculated. The result as a function of h is plotted in figure 5 . With the assumption that the nitrogen content of the char is similar to the nitrogen content of the fuel it is clear that a significant amount of nitrogen is maintained in the char fraction, especially at low A.
328
w
lb
M
la
M
2a
M
2b
M
3a
M
3b
M
4a
M
**
M=Miscanthus, S=Sawdust
A 0.25 0.25 0.19 0.19 0.3 0.3 0.16 807 810 783 782 809 811 769 Temperature ("C) Gasflow(m3ih, drygas) 15.4 15.5 15.2 15.3 15.5 15.3 14.9 Carbon conversion (%) To gas 64 65 56 57 71 70 52 N.A. 5.2 N.A. 6.3 N.A. 6.1 N.A. To tar Rest N.A. 30 N.A. 36 N.A. 24 N.A. Nitrogen conversion (%) TONH3 48 22 45 50 59 61 47 Tar, benzene, toluene and xylene (mg/m3,d y gas) 5514 6453 6631 C6H6 C7H6 1630 1980 1888 CSHlO 173 224 200 CldIS 1047 1338 1327 Other tars 4465 5921 5395 Product gas composition, mean value during the period (vol.%, diy gas) H2 10.0 10.0 9.5 9.7 9.9 9.5 8.8 N2 60.9 60.5 64.4 64.0 57.8 58.5 67.0 co 11.0 11.2 9.4 9.6 11.9 11.6 8.4 CH4 2.8 2.8 2.8 2.9 3.0 2.9 3.0 co2 14.2 14.4 12.5 12.6 16.4 16.5 11.7 C2H4 0.92 0.93 0.92 0.95 0.97 0.96 1.01 C2H6 0.09 0.08 0.08 0.09 0.09 0.09 0.10 * No logging from the gas analysis system, NZcontent estimated from experiment 7
Fuel**
5a
S
5b
S
6a
S
6b
S
7a
S
7b
S
8a*
S
8b*
S
9a
S
9b
S
7594 2748 378 2066 9255 10.6 61.8 11.6 3.6 11.2 1.07 0.13
11.3 11.2 10.8 8.8 66.9 58.3 58.5 61.7 12.9 12.8 11.5 8.5 3.3 3.3 3.5 3.1 11.6 13.1 13.0 11.4 1.05 1.oo 0.99 1.05 0.10 0.10 0.10 0.13
6357 1797 187 1440 3891
32
63 9.0 28
7743 2714 372 2031 8600
7573 247 1 300 1781 8246
24
28
48
38
63 8.8 28
70 69 N.A. 5.7 N.A. 25
53 7.9 40
11.4 55.2 14.0 3.4 15.1 0.95 0.09
-
34
32
11.4 55.6 13.8 3.3 14.9 0.93 0.10
-
-
-
-
55.5 55.5
8106 2377 258 1959 6601
30
10.5 61.7 11.4 4.3 10.6 1.30 0.17
-
-
-
26
10.1 61.6 11.4 4.3 11.1 1.29 0.18
6752 2848 455 202 1 11069
26
N.A. N.A 65 66 N.A. 8.1 N.A. 9.7 N.A. N.A. N.A. 24
5514 1728 181 1148 4649 -
31
77 76 N.A. 5.5 N.A. 19
0.16 0.24 0.24 0.18 0.18 0.29 0.29 0.29 0.29 0.16 0.16 767 796 795 762 762 795 797 795 791 744 737 14.9 16.1 16.0 15.8 15.8 16.2 16.1 16.1 16.1 16.2 16.2
M
4b
Table 2. Experimental results
0 0
06 -
0
02
01
07 0
0.05
0.1
0.15
0.2
0.25
0.3
0.35
Lambda
Figure 3 Conversion of fuel nitrogen to ammonia for Miscanthus in the 0.05m internal diameter gasifier ( 0 ) and for Miscanthus ( 0 ) and sawdust (A) in the 0.2m internal diameter gasifier.
0.30
,
.._..ll..l..l.l_l...-._l.....l._......
........................
.....................................................
0.25 --
5
0.20 --
,o C
'Eal 0.15.
e
8 0 0.10 -.
0.05 -~
0.00 .I 0.00
0.05
0.10
0.15
0.20
0.25
0.30
0.35
Lambda
Figure 4 Conversion of carbon to tar for Miscanthus (0)and sawdust (A) in the 0.2m internal diameter gasifier.
330
3 b m
0.45 0.40
0.35 m
5
U
0.30
-
0.25
g
0.15
3
d
= 0.20 : T?
g 0.10 e
n
0.05
J
1
0.00
0.00
0.05
0.10
0.15
0.20
0.25
0.30
0.35
Lambda
Fzgure 5 Proportion of carbon not detected as gas or tar for miscanthus ( 0 ) and sawdust (A) in the 0.2m internal diameter gasifier.
CONCLUSIONS New experimental results on gasification of Miscanthus have been shown to correlate well with the previously published results from a smaller gasifier. Results from a second fuel, pelletised sawdust with a relatively low nitrogen content, have also been performed and compared with the results on Miscanthus. This fuel resulted in a lower conversion to ammonia, however the apparent difference can be due to uncertainties in the determination of fuel nitrogen concentration in the two fuels. This, and a comparison of the results achieved with the two different analytical methods for determination of fuel nitrogen concentration, clearly points out the problem on comparing experimental data from gasification experiments performed at different laboratories with different fuels. Variations in fuel nitrogen conversion can probably to a large extent be directly related to the uncertainty in the determination of fuel nitrogen concentration. For both the previously published results and the new experimental data there is a clear trend on increasing ammonia concentrations with increasing h. However, taking the relatively low carbon conversion into account and assuming the same nitrogen concentration in the produced char as in the original fuel the trend is less clear. The difference in top and bottom feeding does however remain significant since there were no differences in char yleld between these two situations. The results with respect to the conversion of fuel-bound nitrogen to nitrogencontaining compounds in the gasifier cannot be explained by the theories at hand. The location of the fuel feeding has a definite impact on conversion, feeding from the top gives less ammonia than feeding the fuel directly into the bed. One theory is that oxygen influences the larger tar molecules weakening the nitrogen bonding, thus promoting ammonia formation. When the fuel is fed from the top the pyrolysis gas does not encounter oxygen to the same extent as when feeding directly into the bed. We are
33 1
thus inclined not to believe in the more established theories of today. Our - findings could also have a bearing on NO,- formation during combustion.
ACKNOWLEDGEMENT This work has been performed in collaboration between TPS Termiska Processer AB and KTH within the national program “Fluidised bed combustiodgasification” under contract no P5551-6 and contract no P3452-617. The support from STEM, the Swedish National Energy Administration, is gratefidly acknowledged.
REFERENCES 1. Leppalahti, J. Behaviour of fuel-bound nitrogen in gasification and in hightemperature NH3 removal processes. Ph. D. dissertation. VTT publication 369 (1998). 2 . Leppalahti J. and Kurkela E. Behaviour of nitrogen compounds and tars influidised air gasification ofpeat. Fuel, 70,491-497 (1991). 3. Berg M. and Espen;is B.-G. Minskning av emissionen av kvaveforeningar frdn biobranslebaserade IGCC-processer. TPS Report, TPS-99/44 (In Swedish). 4. Liinanki L. and Karlsson G. VEGATest and Verification - Pressurised gasification of biomass. Vattenfall Report 1994112. (In Swedish). 5 . Goldschmidt B. and Hansson S. Brunsletester i fdrgasningsanluggningen i Varnamo. Sycon. Notes distributed at the Swedish Energy Administration conference in Toftaholm, March 8-9, 2000. 6. Blackadder W.H., Gustavsson P. Svensson 0. and Waldheim L. BIG-GTproject, Summary report 1993-1994. Confidential internal TPS-report TPS-95/27. 7. Waldheim L., Fredriksson C. and Svensson 0. GEFIUNDP Bagasse project, Summary report. Confidential internal TPS-report TPS-99/37. 8. Kilpatrick M. Coal gasification environmental data summary: Sulfur and nitrogen species. Final report EPA/600/7-86/015b, Radian Corporation, Austin TX, April 1986. 9. Mann R.M., Harris G.E., Menzies W.R., Simonson A.V. and Williams W.A. Environmental, health and safety data base for the KR W coal gasification process development unit. Report GRI-85/0123, Gas Research Institute, Chicago IL (1985). 10. Leppalahti J, Koljonen T. Nitrogen evolution from coal, peat and wood during gasification: Literature review. 11. Corella, J., Herguido, J. and Alday, F. J. Research in Thermochemical Biomass Conversion, Eds. A. V. Bridgwater and J. L. Kuester, Elsevier Applied Science, London, 1988, p. 384. 12. Vriesman P., Heginuz E. and Sjostrom K. Biomass gasification in a laboratory-scale AFBG: influence of the location of the feeding point on the fuel-N conversion. Fuel 79, 1371-1378 (2000).
332
Modelling a CFB Biomass Gasifier. Part I: Model Formulation J. Corella and J. M. Toledo Faculty of “Quimicas”, Department of Chemical Engineering, University “Complutense‘‘ of Madrid, 28040 Madrid, Spain
ABSTRACT This communication concerns the basic principles and formulation of a model for a circulating fluidised bed (CFB) biomass gasifier. The main problems are the lot of unknowns concerning the set of kinetic equations needed and the lack of data from commercial CFB biomass gasifiers to check the model. The reacting network to be used in the model is deeply analysed. It is concluded that a self learning or self tuning adaptive model is the best solution
INTRODUCTION There are many papers published about modelling of both bubbling (BFB) and circulating fluidised bed (CFB) coal combustors but not so many on modelling on coal gasifiers. The work of Yan et al. (l), Rhinehart et al. (2), Weimer and Clough (3), Piao et al. (4), and Zhong et al. ( 5 ) concerns modelling of BFB coal gasifiers. Modelling of risers or entrained flow coal gasifiers has been studied by Kojima and co-workers (6, 7, 8), and of CFB coal gasifiers by Fang (9) and Kim et al. (10). Regarding fluidized bed biomass gasification, there have been several attempts on modelling of BFB gasifiers such those of Belleville and Capart (1 l), Chang et al. (12), van der Aarsen (13), Jiang and Morey (14), Corella et al. (15), de Souza-Santos (1 6) and Mansaray et al. (17). About modelling of CFB biomass gasifiers there is nothing significant published in the open literature. The short, but promising, work of UMSICHT in Oberhausen, Ge, (18) and the brief references or notes fiom Lurgi GmbH (19) and from ECN in The Netherlands (20) are the only publications on CFB biomass gasifiers modelling. Unfortunately they do not provide most of the important details about their models which make such publications to be scarcely useful for readers. UCM’s main task in this project is to develop a model for such CFB biomass gasifiers.
SOME DETAILS, BASIS AND ASSUMPTIONS OF THE CFB BIOMASS GASIFIER TO BE MODELLED There are very few CFB biomass gasifiers in operation in the world at commercial, and even at pilot, scale. Such existing gasifiers and previous (since 1980) experience by Corella and co-workers on biomass gasification in fluidised beds already allow a
333
“picture” to be made of a good gasifier (bad FB gasifiers are out of the scope of this modelling). Such good gasifier to be modelled has the following characteristics: 1. 2.
3.
4.
5.
6.
Biomass feeding is made near the bed bottom (such feeding point generates less tar content in the produced gas (21). Both biomass fed into the bed and the gases instantaneously generated from it are homogeneously distributed along the diameter or cross sectional area of the gasifier. In fact this assumption is far from reality. Feeding zone cannot be well modelled nowadays. There exist only some indications of an important heterogeneity in such feeding zone (22). There is not enough data about the “plume” of gases foxmed from the biomass in gasifiers of big size (above 1 m in diameter). There is a secondary air injection. It decreases the tar content in the produced gas (23). Total air is split thus in primary and secondary air flows (there would be a possibility for a tertiary air injection). Besides, it is considered that both primary and secondary air flows are well (homogeneously) distributed along with crosssectional area of the gasifier. Gasfier shape, dimensions and/or topology are, at this stage, unknowns. Inside the CFB gasifier there are continuous gradients of composition, temperature and of concentration of solids (i.e. ref 8). The gasifier was then divided in 20 zones or stripes but, since there was not enough information in kinetics of reactions involved in the gasifier, this first attempt was simplified to only 2 zones. So, only two main zones will be considered in the gasifier: a high concentration of solids at the gasifier bottom (which will be called gasifier bed, with a temperature Tbd), and a low solids density zone, at the gasifier top (with a temperature called Trap). This number of zones can be fiuther increased when good kinetic equations be available. Bed material: A mixture of silica sand (fluidising solid) and calcined dolomite (OCaaOMg) or limestone (OCa) are used as main bed material. As Gil et al. (24) and Olivares et al. (25) have definitively demonstrated a 20-30 wt% of dolomite in the gasifier bed improves enormously the produced gas quality (by decreasing the tar content because of its catalyhc elimination) and avoids bed agglomeration. This 20 % in-bed dolomite content has been recently (March 2000) patented by VTT (25) as “own invention”. Most, if not all, of the existing biomass FB gasifiers use nowadays in-bed dolomite because of its advantages. Nevertheless its use is going to make more complex and difficult the modelling of the CFB gasifier. Besides the silica sand and the calcined dolomite, in the gasifier bed there are the reacting biomass (which very quickly disappears after its entrance to the gasifier, ref. 37), its char and, when all carbon in the char has reacted, its ash. In the gasifier there are 4 or 5 different solids. Fluid dynamics of this mixture is very complex. There are two exits for the solids: Fine fly ash going with the exiting flue gas, to be cleaned, and coarse particles, such as small stones, irons, agglomerated particles, etc, exiting at the bed bottom.
REACTING NETWORK After 20 years working without stop in this process, it has been considered that the reactions which have to be taken into account, because all of them play an important role, are those shown in Table 1. Discussion about these reactions is as follows:
334
Fast pyrolysis at gasifier inlet. Biomass when entering into the gasifier generates the following products: gases: H2, CO, C 0 2 , C a , C2H, and H20 (from biomass moisture and as reaction product too), char, primarytar. T h s pyrolysis step is very fast. For biomass particle diameters of 1 to 10 mm and when pyrolysis is made at 800 "C in fluidised bed, it has been recently calculated (37) to be of 5 to 22 seconds. The gas generated in this step clearly increases the net gas flow in the gasifier (1) and deeply modifies the fluid dynamics near the biomass feeding point. Both facts should have to be taken into account in modelling the biomass gasifier but, again, there is not enough accurate data on these phenomena. There are several hundreds of papers studying the kinetics of reactions involved in the pyrolysis step (27). Unfortunately for the gasifier modelling there is neither a unified approach nor an overall kinetic equation(s) describing the pyrolysis step for all possible biomass under all possible circumstances. Besides, most of the kinetic studies have been made in thermobalances or related apparatuses with low heating rates (around 25 "C/min), far from the high heating rates in fluidised beds of up to a claimed 1000 "C/s or even 10000 "CIS. Kinetic equations obtained in thermobalances would provide results very different from the ones in fluidised bed, which is the present case. Pyrolysis of biomass has been studied by many authors and nowadays it is well known how its product distribution depends, at least, on the following variables: nature or type of the biomass, including its chemical composition and physical properties (28-35). Among all these physical properties, two of them have deep importance and have to be mentioned in a separate way: size and shape (e.g. length, thickness,) of the biomass (29,30,32,36,37), and (biomass) moisture (16,33) heat and mass-transfer processes, heating rate of the biomass (28, 31, 32, 37, 38, 39) (final or bottom bed) temperature (27,28,33,38,39,40,41) residence time (in the reactor) of the biomass particle and of the vapor generated (28,35,39) biomass conversion or reaction extent (34,37,41,42) quenching rate of the vapour and char generated (39). The kinetics for the pyrolysis step are very complex thus. Besides, only data generated in fluidised beds should be used in this modelling but there are only a few of works (i.e. refs. 32,35, 39) made with such a fast heating process and, even more, such papers do not provide kinetic equations. They would have been very useful for the present modelling. For all the above said reasons, in the present model the approach used for the pyrolysis step has been to use the following empirical but easy equation:
335
This approach, equation 8, has been used by many other authors too (2, 11, 12, 14, 72) when modelling the pyrolysis step. Fi and Gi parameters have been calculated in this work from different sources: a) b) c)
Own experimental data (43), from specifically designed tests for this modelling. These tests were made at small pilot plant scale feeding small pine wood chps. They will be published elsewhere. Pyrolysis tests in fluidized bed made by van der Aarsen (13), Laguerie et al. (35), Sjostrom et al. (32), and Wu et al. (39) Fi and Gi depend on the biomass type and size, of its heating rate, (of the feeding device of the gasifier thus), etc ... These two parameters will have to be changed from biomass to biomass and from gasifier to gasifier. They are adjustable and empirical parameters, thus. The values of these parameters obtained by Corella and Gonzalez-Saiz (43) are shown, just as an example, in Table 2. These values are for pine wood chips pyrolysed in a fluidised bed (15 cm i.d.) of silica sand, feeding the biomass near the bed bottom. They will be used in the forthcoming calculations of the overall model. Table 2.- Fi and Givalues of eqn. 8 for pine wood chips (43). COMPONENT Gas H2
co co2 CH4 c2H4 Tar Char
Fi -96.37 59.75 -13.22 28.67 26.63 0.977 93.5 102.9
Gi 0.163 -0.034 0.059 -0.0 15 -0.015 0.0027 -0.081 -0.082
Obviously, using eqn. 8 is (another one more) empirical approach, without a sound kinetic scientific basis. Authors are not happy with it but, in favour of it, it can be said that: 1st) there is not yet an universally accepted kinetic equation for the pyrolysis step, 2nd) it works and is u s e l l for the gasifier modelling, 3rd) it can be replaced for another more scientific approach or subroutine when it be available. From the flow rate of biomass fed to the gasifier, pyrolysis temperature (those of the bottom bed), and biomass formula (composition), eqn. 8 will provide the amounts (kg/h) of gas, char and tarl after the pyrolysis step. In one of the three different models developed by Corella and co-workers for FB biomass gasifiers, the elemental analysis or formula (CcREs H m s Oom~)of the pyrolysed char and tarl (to be called residue, RES) formed in the pyrolysis step has to be known. It can be easily calculated by mass balances for each element (C, H, 0)from the elemental analysis of the biomass (C, HB Oc) and the amounts of gases, char + tarl, and H20 generated in the pyrolysis step, calculated from eqn. 8 and values in Table 2 or similar ones. From such easy mass balances, the amounts of C, H and 0 in the residue (char + tarl) are:
336
mBx12xA - mc0x12 - mCH,x12- mco,x12 - mClH,x24 28 16 44 28 PMB
C,RES =
m xB mcH,x4 - mClH,x4 H,RES = B 16 28 PM B
0,RES=
mBx16xC - m,,x16 28 PMB
-
(9a)
mHl
mcolx32 44
In-bottom-bed reactions Just after pyrolysis, three different types of reactions start. They will be analyzed here one by one. Concerning tar reactions. Tar just formed from biomass pyrolysis, whch is called tarl, suffers a lot of catalytic and not-catalytic reactions, as published by Chan et al. (30) and Cozzani et al. (36,44). Tar1 reacts with HzO, COz, 02,H2 and suffers thermal degradation reactions too, such as cracking. A secondary and “more-refractory-to-be destroyed" tar, which will be called tar2, is produced together with H2, CO, CH4,,Stochiometryof these reactions is not well established yet. Cozzani et al. (36, 44) have evaluated the homogeneous (not-catalytic) tarlcracking time (to convert 90 % of the primary tar at temperatures of 800-850 "C) in 0.7 to 1.8 seconds (44). It is a very fast reaction then. Activation energy for this not catalytic tarl-craclung has been evaluated in 102 kJ/mol(44) or in 144 kJ/mol by Chan et al. (30). Some homogeneous reactions concerning tar evolution in gasification gas have been studied by Heppola and Simell(45) but it was considered by UCM not enough for the gasifier modelling and consequently a lot of fbrther research has been made by UCM in this subject for thls project. Tests have been made feeding small pine wood chips at the bottom of a fluidised bed. Fluidising agents have been Nz(pyrolysis tests) and air with an equivalence ratio (ER) of 0.24 (gasification tests). Gasification tests were of two types: 0
with silica sand (only) in the fluidised bed, without dolomite (ref. 46) with 20-30 wt% dolomite in the bed, mixed with 80-70 wt% silica sand (ref. 24).
Remember that the calcined dolomite (0Ca.OMg) is a well known tar-elimination (in gasification gas) catalyst by steam and dry (COZ) reforming reactions (24, 25,47). The three said exp. conditions generate a very different tar yield and composition: 0
0
-
-
pyrolysis tarl pyrolysis + not catalytic reactions tar2 pyrolysis + not catal. reactions + catalytic reactions --* tar3
-
A fourth case (a real situation in the CFB gasifier) has still to be considered when modelling the upper zone of the gasifier: 0 pyrolysis + not catal. reactions + catalyhc reactions + 2nd air flow tar4
337
Experimental yields of these tars are shown in Figure 1. There can be appreciated how tar yields at the gasifier exit without in-bed dolomite (tar 2) are much lower that tar yields from the pyrolysis alone, tar 1, and given by eqn. 8. For instance, working at 800 “C,tar veld from the pyrolysis step of around 90 g tarlkg biomass daf is decreased till 13 g tar2kg daf if there is gasification. Oxygen from the primary air burns some tar. Not-catalytic reactions in the gasifier bed must be taken into account, thus. Catalytic tar elimination reactions (tar1 -*tar2, and tar2 + tar3) must be considered too, if in-bed dolomite is used, as can be seen in Figure 1. With such in-bed dolomite the above said tar yield is still decreased till 7 g tar3kg daf. Apparent activation energy for the tar2 + tar3 reaction has been calculated many times in 100 kT/mol(47). But not only in-bed dolomite is the catalysts for tar elimination reactions. The same produced char and ash act as catalysts in several tar-elimination reactions (62,63,64,65). The secondary air flow still reduces this yield to 0.4-4 g tar4/ kg daf (23,48). Tar yield predicted by eqn. 8 is reduced then by about 100 times because of the homogeneous and heterogeneous (catalytic) in-bed reactions and by the secondary air flow. Unfortunately for our modelling, there are not any good kinetic equations available, which fit data shown in Figure 1, that’s to say, to predict the tar (tar4) yield at the gasifier exit which is one basic objective of the present modelling. Tar1 yield from eqn. 8 has to multiplied by a factor of around 0.01 (from Figure 1) to know the realistic tar4 yield at gasifier exit, but such Figure 1 has not an universal validity. In this point or submodel we found again a lack of good kinetic equations to calculate tar2, tar3 and tar4 yields and their corresponding compositions.
Benzene
Naphthalene
2-rings
3&4rings
Molecular Weight (Mw)
Fig 2.- Evolution of the tar, as a continuous mixture, along with time of reaction Very recently’Corella and co-workers (49) have determined the kinetics of the catalytic disappearance of tar as a “continuous mixture” by using the SPA method 338
developed by KTH in Stockholm. They have sampled and characterized the tar present in the flue (gasification) gas after several residence times of the gas in a catalytic reactor. The evolution of the tar as a continuous mixture with gas residence time (7,in seconds) for a concrete case is shown in Figure 2. The kinetics of the catalytic tar elimination are described for this case (and for 6 0 . 1 ) by:
Being 0’ the variance of the distribution of molecular weights (MW) of the tar components. For a bed with a nickel catalyst:
,,k
=
170 g/mol.s
= 7,100 g2/molz.s and n’ = 0
klapp
n
These values do not have universal validity, they are for a concrete case only. They can not be used for a general model, thus. Concerning the variations of the H2, CO, COz, CH4, CzH, and HzO contents in the flue gas because of the tar elimination reactions, not only the kinetics of the reactions involved (eqns. 2 and 3) are not known but even their stochiometry. For instance: both H20 and COz are reactants, which disappear by reaction with tarl, tar2 and tar3 by steam and dry reforming reactions, and products too in the tar-elimination reactions. Their overall variation by reactions given by eqns. 2 and 3 are other unknowns. Yields of Hz, CO, COz, CH4, CzH, and HzO obtained by eqn. 8 have then to be multiplied for some modifpng factors whch are unknowns. Concerning char reactions Char is not C. Its formula is complex. So, we will write and use “char” and not C for the following reactions. Char generated in the pyrolysis step can react with the following reactants present in the flue gas in the bottom-bed zone: Char + O2 Char + COz Char + H20 Char + Hz
+ + +
+
coz + co + ... ? co+... ? CH,+CO+ CH4 + ... ?
... ?
These reactions have been studied one-by-one by several authors and their kinetics have been obtained usually under pure (with 0 2 or COZ or HzO or Hz) atmospheres. Just as an example, some kinetic equations published for these reactions are: for eqn. 4a: (-r) = 5.3.105.[exp(-125.103/R.T)].PO~53.( 1-X)0.49
339
(ref. 53)
for eqn. 4b: various eqns. in refs. 54 and 57, and (-r) = 0.25.exp(-E/R.T) (-r) = 1.7.10-' [exp(-14.1O6/R.T)].PCo2
(ref. 5 5 ) (refs. 56,58,59) (ref. 50) (ref. 43)
for eqn 4d: various eqns. in ref. 59 and (-r) = 4.8.1O'.T.[exp(- 179oo/T)]'cH2 k = 0.75~1O3~exp(-2.3~1O'/R~T)
(ref. 60) (ref. 12)
Many other kinetic equations and references could be cited too but it would be, to these authors, an academic show of how extensives are their files in this field. No more equations and references are here given thus because of 1st)
Most of the chars used in these studies come from coals, and not from biomass which is our only interest here. 2nd) Most, if not all, of the kinetic equations have been obtained under simple gas surrounding atmospheres (one, two or three compounds). Nevertheless in the gasifier there are at least four simultaneously gasifying agents (02,C02, H20 and H2) and some other species (CH4, C2Hn,tars, ... particulates) that affect the overall gasification of the char. 3rd) Very often these kinetic equations were obtained under exp. conditions (i.e. thermobalances) quite different than those existing in CFB biomass gasifiers. These authors consider that a system of good kinetic equations for the biomass-char gasification reactions under the true gasifier atmosphere do not exist yet. This conclusion can be discussed and even criticised, of course. The authors are open to a discussion and experimental checking with the person(s) who considers that holds a system of kinetic equations to correctly describe the network of reactions 4a-4d. These authors do not know such system of kinetic equations and have again to look for an empirical and easy-tome approach. In such a simplified approach the basic question is the "char partitioning" with all existing and competitive reactants. Since each reaction, eqns. 4a to 4d, has different reactants and products, to calculate the product distribution after the b.2 step, the fractions of char reacting with 02,with C02, with H20 and with H2 would need to be known. Since there are not good kinetic equations that would provide a scientific answer to this question, the pgoblemremains open. Another empirical approach (one more!) has to be used for this step. It will be shown later. Concerning gas-phase reactions Concerningprimaiy air (03 A situation very similar to those just explained for char occurs for the oxygen introduced by the primary air. There exist at least the following reactions (not adjusted):
340
0 2
+ co
+
+ H2
4
+ CH4 +GHn +tarl,2,3 + char
co2
+
+ +
H20 C02 + CO + H20 + H2? C02 + CO + H20 + H2? C02 + CO + H20+ H2?+ ... ?
(5a) (5b) (5c) (54 (5e)
c02
(50
+ co + ... ?
Remember that in gasification conditions, the air is clearly substoichiometric (ER from 0.19 to 0.40; recommended around 0.25) and that reaction products, besides C02 and H20, are Hz, CO, etc, ... There could be here cited several kinetic equations for some of these reactions (61) but the same comments above said for char gasification reaction here applies. A system of good kinetic equations describing the said reacting network (eqns. 5a to 50 under gasifier operating conditions is not known yet. Even the stoichiometry is not well known. So, there appears a problem about the partitioning of the oxygen (whch is substoichiometric). In which percentage does the oxygen react with each one of the competitive and simultaneous reactants present in different amounts along the height of the gasifier? The authors have not an answer to this question and are obliged to think in another empirical solutiodapproach. Concerning water (H20) An identical situation occurs for H20. Maybe the problem is now still more difficult because H20 can come from 3 different sources:
biomass moisture airmoisture intermediate product in several in-bed reactions. The H20can disappear by the following reactions:
co
4-
+
+ CH4
+
H20 + C2Hn + tarsl, 2, 3 + coke
-
4
+
C02 + Hz CO + C02 + H2 CO + C02 + H2 + ... ? tars2, 3 + H2 + CO + ... ? CO + H2 + ... ?
(64 (6b) (6c) (64 (6e)
The partitioning of H20 among all simultaneous and competitive reactants is not known at all. A system of kinetic equations to describe well this network is not known again and empiricism is the only solution known by the authors. To know the axial profile of H20 along the height of the gasifier is very difficult and even not possible, thus. Upper part of the gasifier All the above problems can be applied to the upper part of the CFB gasifier. With an accurate fluid dynamic model the axial profiles of biomass, char, calcined dolomite and silica sand could be calculated, but the lack of accurate kinetic information on the effect of the concentration of these solids on the rate of all reactions involved avoids to calculate the axial profiles of the gaseous compounds (H2, COYC02 ... and tars). So, it
34 1
has no sense for these authors to consider 2, 3, ... 20 ... 100 zones for the gasifier if there is not an accurate knowledge of the kinetics of reactions involved. Even more, if the concentration profile for H2, CO, CH4, tars can not be calculated, the temperature profile inside the CFB gasifier will not able of being calculated with a minimum accuracy because each reaction in the network has a different heat of reaction and the overall heat involved (generated or consumed) in each zone depends on the unknown extent or conversion of each reactions. There will be needed again a lot of empirical assumptions. It has to be remembered that in the upper part of the gasifier there has to be considered, besides reactions given by eqns 2,3,4 and 6, the secondary air ( 0 2 ) flow: + co + (74 + H2 + (7b) C02 + CO + H2 + H20 + ... ? ( 7 ~ ) O2(Secondary air) + CH4 + CzH, + (74 + tars2, 3 ---f (74 + char --+ (70
-
The partitioning of this secondary air among all simultaneous and competitive reactants is unknown. It would be excellent that the secondary air (0,) would react only with the unwanted tars and char, leaving unreacted the H2, CO, and CH,. It would increase the LHV and quality (tar content) of the produced gas and the overall effectiveness (by decreasing the char yield) of the gasification process. Unfortunately nobody knows how “to direct” this secondary air towards the tars and char only. SUMMARY
Summarising the reacting network in the CFB biomass gasifier: a) b) c)
it is full of unknowns. There is not a good system of kinetic equations describing the reacting network under realistic gasification conditions. Consequently, A lot of empiricism has to be accepted and introduced in the model (concerning the reacting scheme and its kinetics).
The final approach for the overall model has to be empirical, thus. Since the addition of a lot of unknowns (as it is the present case) gives or generates ONE UNKNOWN, if this “unknown” could be obtained from the running gasifier, as a self learning or self tuning adaptive model, the problem would be solved. This is a solution, model, approach andor situation not so strange: it is used in industry in a lot of complex processes. So, the final and adopted solution is not so bad, “sad” and new. Just the present degree of knowledge on biomass CFB gasifiers and the principles of Chemical Reaction Engineering and of Control of Processes have to be known, accepted and used. NOMENCLATURE
d.a.f.
dry,ash free. 342
particle diameter, by sieving, mm. equivalence ratio, defined as the air-to-fuel ratio used in the gasifier divided by the air-to-fuel ratio for the stoichiometric combustion, dimensionless Hbed height of the bed at the gasifier bottom bed, bulk fixed bed conditions, c m LHV lowest heating value of the produced gas, MJ/m3,, dry basis. m3n cubic meter, normal conditions (OOC, lat or 101 Ha). Mi; Mdaf weight of product i formed, and of biomass daf fed, g or kg. Tkd, Ttop temperature in the gasifier bottom bed and in its upper part, respectively, “C. tar* tar referred to the tar sampling and analysis methods described in ref. 4. YG,Yi gas yield and yield to product i at the gasifier exit, m3, dry gas / kg biomass fec daf and g i/kg daf respectively. Greek symbols 70 space-time for the gas in the gasifier bottom bed, defined as Hbed/&, s.
dP ER
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53. Janse, A. M. C.; Prins, W.; van Swaaij, W. P. M., Ind. Eng. Chem. Res., 1998, 37, 3909-3918. 54. Zhong, T.; Jianmin, Z.; Yang, W., J. Combust. Science and Technol., 1996,2, 3137. 55. Sotirchos, S. V.; Crowley, J. A., Ind. Eng. Chem. Res., 1987,26, 1766-1773. 56. Adschiri, T.; Furusawa, T., Fuel Process. Technol., 1987, 15,135-144. 57. Matsui, I.; Kunii, D.; Furusawa, T., Ind. Eng. Chem. Res., 1987, 26,91-95. 58. Adschiri, T.; Furusawa, T., Chem. Eng. Sci., 1987,42, 1313-1317, and 1319-1322. 59. Goyal, A.; Zabransky, R. F.; Rehmat, A., Ind. Eng. Chem. Res., 1989, 28, 17671778. 60. Haynes, H. W., AIChE J., 1982,28,517-521. 61. Kulasekaran, S.; Linjewile, T. M.; Agarwal, P. K., Fuel, 1999, 78,403-417.
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Influence of the Reaction Atmosphere on Gas Production and. Composition in the Catalytic Conversion of Biomass L. Garcia, M.L. Salvador, J. Arauzo, R. Bilbao Chemical & Environmental Engineering Department, University of Zaragoza, 50015 Zaragoza, Spain.
ABSTRACT The objective of this work is to study the influence of the reaction atmosphere on the gas product obtained in the catalytic conversion of biomass. Pyrolysis, steam gasification and CO, gasification experiments were carried out in a bench-scale installation based on Waterloo Fast Pyrolysis Process (WFPP) technology. Pine sawdust was the biomass used. At the reaction temperature, 700 OC, the presence of the catalyst exerts a significant effect transforming tars into gas and improving the quality of product gas. A Ni-AI coprecipitated catalyst prepared in our laboratory was chosen because of its mechanical strength and activity. However, the catalyst has the inconvenience of its deactivation, mainly caused by carbon deposits. A comparative analysis focusing on H,, CO, CO,, CH, and C, yields generated in the three processes has been carried out using a similar catalyst weighthiomass flow rate (W/mb) ratio. For the W/mb ratios studied (0.32 and 0.16 h) the following results have been observed. H, yields were similar in pyrolysis and in CO, gasification, while the highest yield to this gas was obtained in steam gasification. The highest CO yield was generated in CO, gasification. CH, and C, yields were higher in pyrolysis than in CO, gasification and steam gasification. The highest H,/CO (vol/vol) ratios were obtained in steam gasification with values at the beginning of the experiment at around 4 for an steam/biomass ratio of 1.5. At the beginning of the experiment, the HJCO ratio was around 1 for pyrolysis and 0.6 for CO, gasification, using a CO,/biomass ratio of 1.4. INTRODUCTION
Biomass can be gasified to obtain gas for power generation, hydrogen production or chemical synthesis. The reaction atmosphere can influence very significantly the gas composition and its posterior use. For instance, steam gasification of biomass (1 - 5 ) and catalytic steam reforming of liquids from pyrolysis (6,7) produce a rich-hydrogen gas and a CO, residual stream. The present work has the novelty of introducing CO, as a gasifying agent. Pyrolysis and steam gasification are processes currently being investigated to obtain
346
gases from biomass. However, few works in thermochemical conversion of biomass have been carried out using CO, (8-12). Nowadays there is a great interest in the conversion of CO, into useful gases. CO, as a gasifying agent for biomass produces a synthesis gas with a low H,/CO ratio which can be useful in 0x0-synthesis reactions (13). CO, biomass gasification could contribute to the environmental problem of CO, release. The possibility of recycling C02 from the product gas and converting it into useful gases appears to be a promising option. The reaction temperature is one of the most influential variables in thermochemical biomass conversion. Pyrolysis, steam gasification and CO, gasification are endothermic processes, and an input energy is required. Therefore, a decrease in the operating temperature allows a significant saving in energy and a consequent decrease in costs. However, this temperature decrease causes lower gas yields and more tar production. The use of catalysts can solve this problem by transforming tars into gas. The catalyst, however, presents the problem of its deactivation. Commercial and research nickel-based catalysts have been widely used in pyrolysis and gasification of biomass (14-1 7).
Pyrolysis, steam gasification and CO, gasification experiments were carried out at 700 "C using a Ni-AI coprecipitated catalyst prepared in our laboratory. This catalyst
was chosen because of its mechanical strength and initial activity. These three processes were compared using a similar W/mb ratio because of the important influence of this ratio on catalytic processes (5,18). EXPERIMENTAL EXPERIMENTAL SYSTEM The experimental system used to carry out this work is a bench-scale installation based on Waterloo Fast Pyrolysis Process (WFPP) technology (19,20). The biomass feeder supplies a continuous biomass flow rate of between 5 and 100 g/h. Thermal decomposition of biomass occurs in a fluidized bed reactor with an inner section of 4.35 cm2. In the experiments, two streams are introduced into the reactor. One is nitrogen and transports the biomass into the fluidized bed. The other enters at the bottom and reaches the bed through the distributor. This latter stream is nitrogen in pyrolysis experiments, mixtures of steam and nitrogen in steam gasification, and C 0 2 and nitrogen in CO, gasification experiments. The water is supplied in liquid state by a syringe pump that allows a constant and accurate flow rate of between 0.1 and 99.9 mL/h. The steam is generated while the water flows towards the reaction bed passing the electric furnace. The total flow rate into the reactor is 17 15 cm3(STP)/min,of which 1215 cm3(STP)/min transports the biomass to the reaction bed while the rest enters through the distributor. The product gas is cleaned of char particles using a cyclone. The liquid products and water are retained in a system of two condensers and a cotton filter. The gas flow rate is then measured using a dry testmeter, and the CO and CO, concentrations are continuously determined by an infrared analyzer. In addition, gas samples are taken at regular time intervals and analyzed by chromatography to determine the percentages of H,, CO, CO,, CH4and C, (C,H,, C2H4, C,H,) in the product gas. Reaction temperature,
347
total gas flow, and CO and CO, concentrations are registered by a data acquisition system. A scheme of this experimental system with minor modifications can be found elsewhere (5,18). The experiments have been carried out at 700 "C and at atmospheric pressure. Two catalyst weighthiomass flow rate (W/mb) ratios of around 0.32 and 0.16 h have been employed. The catalyst weight or the biomass flow rate was changed to obtain these ratios. The reaction bed was composed of catalyst diluted with sand. All the experiments were performed with the same bed volume (34.39 cm3). The steamhiomass (g/g) (S/B) ratio in the steam gasification experiments was calculated dividing the total mass of inlet steam by the total mass of fed biomass. This ratio was varied from 0.53 to 1.54. The CO,/biomass (dg)(CO,/B) ratio in the CO, gasification experiments was calculated dividing the total mass of inlet CO, by the total mass of fed biomass. This last ratio ranged from 0.57 to 1.39. The biomass used was pine sawdust with a moisture content of about 10% and a particle size of -35W150 pm. The results of the elemental analysis (in YOmass) of the pine sawdust are carbon 48.27%, hydrogen 6.45%, nitrogen 0.09%, and oxygen (by difference) 45.19%. CATALYST The Ni/AI catalyst used was prepared in the laboratory by coprecipitation following the method of Al-Ubaid and Wolf (21). Nickel nitrate hexahydrate and aluminum nitrate nonahydrate were dissolved in distilled water in the appropriate proportions to obtain a Ni:AI ratio of 1:2. This solution was precipitated with amonium hydroxyde until the pH reached 7.9. The precursor was calcined in air atmosphere, with a low heating rate, reaching a final temperature of 750 "C during 3 hours. The catalyst employed was not reduced prior to the biomass reaction due to the results obtained in previous works (5,22) in which it was concluded that the gases generated in the thermal decomposition of biomass at 700 "C have the ability to reduce the Ni/AI catalyst prepared by coprecipitation and calcined at 750 "C. The precursor and the calcined catalyst were characterized by various techniques such as nitrogen adsorption, mercury porosimetry, X-ray diffraction (XRD), atomic emission spectrometry by inductively coupled plasma (ICP), thermogravimetric analysis, and temperature-programmed reduction (TPR). More details about the catalyst preparation and characterization can be found in a previous work (22). RESULTS AND DISCUSSION This work is focused on analyzing the influence of the reaction atmosphere on gas yields and the H,KO (vol/vol) ratio of the product gas. With this aim, the results of pyrolysis, steam gasification, and C 0 2 gasification experiments, with a similar W/mb ratio, are represented versus time. The change in gas yields over time indicates catalyst deactivation and the effect of the reaction atmosphere on the loss of catalyst activity. Figures 1 to 5 represent H,, CO, C02, CH, and C, yields respectively. Figure 6 shows the H,/CO ratio of product gas. Figures (a) correspond to the experiments carried out with a W/mb ratio close to 0.32 h and include the results of pyrolysis, steam gasification with S/B ratios of 0.53 and 1.54, and CO, gasification using a C0,hiomass ratio of 1.39. Figures (b) show the experimental results obtained using a W/mb ratio
348
around 0.16 h. These results include pyrolysis, steam gasification with a S/B ratio of 0.93 and CO, gasification with CO,/biomass ratios of 0.57 and 0.94. Figure 1 represents H, yields versus time. For all three processes, a decrease in the H, yield is observed over time. This fact is assumed to be due to the catalyst deactivation.
(a) 0.16
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-
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0.08
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W/mb FJ 0. I6 h
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Fig. 1. Influence of the reaction atmosphere on H, yield. (a) W/mb = 0.32 h, (b) W/mb = 0.16 h.
The H, yield obtained in pyrolysis is similar to that generated in CO, gasification for both W/mb ratios. The CO,/biomass ratio does not influence the H, yield as can be seen in Figure lb. The highest H, yields are obtained in steam gasification for both W/mb ratios. The S/B ratio influences the H, yield very significantly. An increase in the S / B ratio causes an increase in the H, yield (Figure la).
349
CO yields are shown in Figure 2. The decrease in the CO yield with time might be caused by the loss of catalyst activity. The highest CO yields are obtained in CO, gasification. The CO,/biomass ratios used do not influence the CO yield as can be observed in Figure 2b. A comparison of Figures 2a and 2b show that the CO yield is higher at short reaction time in pyrolysis and CO, gasification, when using the highest W/mb ratio. This result could be related to the influence of W/mb ratio on catalyst deactivation (1 8).
- (a) - >p(xxX 1.2 -
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Time (min) Fig. 2. Influence of the reaction atmosphere on CO yield. (a) W/mb c 0.32 h, (b) W/mb c 0.16 h. Some of the CO yields generated in CO, gasification using the W/mb ratio of around 0.32 h are higher than 1. This fact can be explained by the conversion of CO, into CO. The S/B ratio also influences the CO yield very significantly. As can be seen in Figure 2a, an increase in the SIB ratio decreases the CO yield. Figure 3 represents CO, yields over time. The smallest CO, yields are obtained in pyrolysis, with values lower than 0.1 g CO,/g sawdust. In CO, gasification, the CO,
350
yield is strongly influenced by the CO,/biomass ratio used. The difference between the CO,/biomass ratio and the CO, yield in the product gas is related to the conversion of CO,. This conversion increases when the COJbiomass ratio increases. An increase in CO, yields over longer reaction time is observed in the CO, gasification experiments. This result must be due to catalyst deactivation.
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A decrease in the CO, yield with the reaction time is observed in steam gasification. This evolution could be related to the loss of catalyst activity. The S/B ratio also influences the CO, yield. When the S/B ratio increases, the CO, yield increases very significantly (Figure 3a). The CO, yield at the beginning of the experiment is around 1 when using a S/B ratio of 1.54 and 0.4 with a S/B ratio of 0.53. Figure 4 shows CH, yields. The highest CH, yields are obtained in pyrolysis for both W/mb ratios. The increase in CH, yields when the reaction time increases is assumed to be caused by catalyst deactivation. The S/B ratio also influences the CH,
35 1
yield (Figure 4a), an increase on CH, yields occurring when the S/B ratio decreases. For the W/mb ratio of 0.32 h, similar CH, yields are observed in CO, gasification as in steam gasification using a S/B ratio of 1.54.
0.07
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For the W/mb ratio of around 0.16 h (Figure 4b), higher CH, yields are obtained in CO, gasification than in steam gasification. The evolution of CH, yields with time is not clear for CO, gasification. A slight effect of the CO,/biomass ratio on CH, yields can be observed. Higher CH, yields are generated using the smallest CO,/biomass ratio (Figure 4b). Figure 5 represents C, yields. A similar performance as for CH, yields is observed when using the W/mb ratio of 0.32 h. For the W/mb ratio of 0.16 h (Figure 5b), the C, yield obtained in pyrolysis presents the highest values while the smallest values correspond to steam gasification. No influence of the CO,/biomass ratio is observed.
352
0.07
m
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0.00
0
20
40
60
80
100 120 140 160 180 200
Time (min) Fig. 5. Influence of the reaction atmosphere on C, yield. (a) W/mb * 0.32 h, (b) W/mb = 0.16 h. The results obtained can be explained by considering the reactions involved in the processes. We can assume that the main reactions in catalytic pyrolysis are catalytic cracking of tars and light hydrocarbons, which will explain the increase in gas yields when the catalyst is present in the reaction bed (1 8). Steam reforming of tars (reac. I), methane (reac. 2) and C, (reac. 3), and the water-gas shift reaction (reac. 4) can explain the final gas composition generated in catalytic steam gasification. Tar+H,O s nCO + m H, CH,+H,O
(reac. 1)
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G
(reac. 2) (reac. 3)
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(reac. 4)
353
The increase in the S/B ratio causes an increase of steam in the reaction atmosphere that produces the increase of CO, and H2 and the decrease of CO. These tendencies are observed in Figures 1 to 3. The influence of the S / B ratio on the gas yields of H,, CO and CO,, is evidence of the participation of the water-gas shift reaction. CO, reforming of tars (reac. 5), CH., (reac. 6), C, (reac. 7), and the inverse watergas shift reaction (reac. 8) can be involved in catalytic CO, gasification. All these reactions contribute to the high CO yield obtained in CO, gasification. Tar+CO, a p C O + q H ,
(reac. 5 )
CH,+CO,
(reac. 6)
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(reac. 7)
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(reac. 8)
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+
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20
40
60
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80
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Time (min) Fig. 6. Influence of the reaction atmosphere on HJCO ratio of the product gas. (a) W/mb SS 0.32 h, (b) W/mb I 0.16 h.
354
Dark brown heavy oil and solvent soluble material were obtained in this manner. The amount of formed gas was measured with a gas meter and the composition of evolved gases was determined by gas chromatography with TCD (silica gel column of 60/80 meshes). Conversion to gas, oil and carbonized solid were calculated by the following equation: Weight of material obtained fiom each phase x 100 Conversion (%) = Weight of drymg biomass The moisture content of rice straw was measured after drymg at 105 "C for 5 hrs and the measured value was 8 %. Ash was analyzed after heating at 600 "C for 5 hrs and its value was 4 % (ASTM D1102-84). PREPARATION OF CATALYSTS
The nickel catalyst (about 50 wt% nickel on kieselguhr) was prepared by an ordinary precipitation method. Sodium carbonate solution was added to a slurry of kieselguhr and nickel nitrate solution at 70 "C and precipitate was obtained. This precipitate was washed with water thoroughly and then was dried at 105"C for 12 hrs, crushed to 60150 mesh, calcined at 350°C for 4 hrs. This was activated with 100% hydrogen at 200, 300 and 350 "C for 4 hrs. These prepared catalysts were stored in nitrogen atmospheric bottle and desiccator. CHARCTERIZATION OF C A T L Y S E REDUCTION (TPR)
TEMPERATURE PROGRAMMED
For TPR studies, 50 mg of the calcined catalyst were loaded in a quartz reactor and heated at 673K for 6 h, followed by cooling down to room temperature in argon gas. A high purity premixed gas containing 95% argon and 5% hydrogen was used as reducing agent. Traces of oxygen and water vapor were removed by passing the gas through an activated molecular sieve trap kept in an ice bath. The catalyst bed was heated linearly at 5Wmin. A continuous TPR response profile was obtained when the difference in H2 concentrations in the gas streams between reference and sample, resulted fiom the reduction of metal oxide by H2. RESULTS AND DISCUSSION EFFECT OFALKALI METAL CARBONATE
The effect of alkali metal on the formation of gas, oil and carbonized solid was examined in the catalytic gasification of rice straw over nickel catalyst. Various alkali metal carbonates were separately added with nickel catalyst and the results were presented in Table 1. Addition of alkali metal carbonates gave a large influence on conversion to gas and composition of gas. The formation of gas was dominated and increased in the following order; Li< Cs< KSNa. In case of lithium, conversion of oil, char and gas shows similar to without catalyst. This result caused to a solubility of lithium. The solubility of lithium24is the lowest than that of other carbonate.
360
explain the increase of H, and CO, yields and the decrease of the CO yield. However, the CO,/biomass ratios used barely influence the gas generated in CO, gasification. The highest H,/CO ratios are obtained in steam gasification, with values at the beginning of the experiment ranging from 4,using a S/B ratio of 1.54,to 2 with a S/B ratio of 0.53. The H2/C0 ratio is around 1 at the initial time of the pyrolysis experiment, and the lowest values of this ratio correspond to COz gasification. ACKNOWLEDGEMENTS The authors express their gratitude to the Ministerio de Educacibn y Cultura for providing financial support for the study (Project 2FD97-0890) and also to the Fundaci6n Caja de Madrid for a research grant and financial support awarded to L. Garcia.
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10.
Aznar, M.P.; Corella, J.; Delgado, J.; Lahoz J. (1993) Improved Steam Gasification of Lignocellulosic Residues in a Fluidized Bed with Commercial Steam Reforming Catlysts. Ind Eng. Chem. Res., 32, 1- 10. Aznar, M.P.; Caballero, M.A.; Gil, J.; Olivares, A.; Corella, J. (1997) Hydrogen by Biomass Gasification with Steam-Oxygen Mixtures and Steam Reforming and CO-Shift Catalytic Beds Downstream of the Gasifier. In: Making a Businessfrom Biomass (Ed. by R.P. Overend, E. Chornet), pp.859-860.Pergamon, New York, 1997. Rapagna, R.; Jand, N.; Foscolo, P.U. (1 998) Catalytic Gasification of Biomass to Produce Hydrogen Rich Gas.Int. J. Hydrogen Energy, 23,551-557. Turn,S.; Kinoshita, C.; Zhang, Z.; Ishimura, D.; Zhou, J. (1998)An Experimental Investigation of Hydrogen Production from Biomass Gasification. Int. J. Hydrogen Energy, 23,641-648. Garcia, L.; Salvador, M.L.; Arauzo, J.; Bilbao, R. (1999) Catalytic Steam Gasification of Pine Sawdust. Effect of Catalyst WeightlBiomass Flow Rate and Steam/Biomass Ratios on Gas Production and Composition. Energy&Fuels, 13, 851-859. Wang, D.; Czenik, S.; Montank, D.; Mann, M.; Chornet, E. (1997) Biomass to Hydrogen via Fast Pyrolysis and Catalytic Steam Reforming of the Pyrolysis Oil or Its Fractions. Ind. Eng. Chem. Res., 36,1507-15 18. Wang, D.; Czernik, S.; Chornet, E. (1998)Production of Hydrogen from Biomass by Catalytic Steam Reforming of Fast Pyrolysis Oils. Energy&Fuels, 12, 19-24. Minkova, V.; Marinov, S.P.; Zanzi, R.; BjSrnbom,E.; Budinova, T.; Stefanova, M.; Lakov, L. (2000).Thermochemical Treatment of Biomass in a Flow of Steam or in a Mixture of Steam and Carbon Dioxide. Fuel Process. Techn., 62,45-52. Pindoria, R.V.; ,Megaritis, A.; MessenbSck, R.C.; Dugwell, D.R.; Kandiyoti, R. (1998). Comparison of the Pyrolysis and Gasification of Biomass: Effect of Reacting Gas Atmosphere and Pressure on Eucalyptus Wood. Fuel, 77, 12471251. Simell, P.A.; Hakala, N.A.K.; Haario, H.E.; Krause, A.O.1. (1997). Catalytic Decomposition of Gasification Gas Tar with Benzene as the Model Compound. Ind. Eng. Chem. Res., 36,42-51.
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11. Simell, P.A.; Hepola, J.O.; Krause, A.O.I. (1997) Effects of Gasification Gas Components on Tar and Ammonia Decomposition over Hot Gas Cleanup Catalysts. Fuel, 76, I 1 17-1127. 12. Arauzo, J.; Radlein, D.;Piskorz, J.; Scott, D.S.(1994) A New Catalyst for the Catalytic Gasification of Biomass. Energy&Fuels, 8, 1192-1196. 13. Teuner, S. (1987) A New Process to Make 0x0-feed. Hydrocurb. Process., July 52. 14. Baker, E.G.; Mudge, L.K.; Brown, M.D. (1987) Steam Gasification of Biomass with Nickel Secondary Catalysts. Ind. Eng. Chem. Res., 26, 1335-1339. 15. Arauzo, J.; Radlein, D.;Piskorz, J.; Scott, D.S. (1997) Catalytic Pyrogasification of Biomass. Evaluation of Modified Nickel Catalysts. Ind. Eng. Chem. Res., 36, 67-75. 16. Simell, P.; Kurkela, E.; Stahlberg, P. (1994) Formation and Catalytic Decomposition of Tars from Fluidized-Bed Gasification. In: Advances in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater), pp. 265-279. Blackie A&P, London. 17. Kinoshita, C.M.; Wang, Y.; Zhou, J. (1995) Effect of Reformer Conditions on Catalytic Reforming of Biomass Gasification Tars. Ind Eng. Chem. Res., 34, 2949-2954. 18. Garcia, L.; Salvador, M.L.; Arauzo, J.; Bilbao, R. (1998) Influence of Catalyst WeighVBiomass Flow Rate Ratio on Gas Production in the Catalytic Pyrolysis of Pine Sawdust at Low Temperatures. Ind. Eng, Chem. Rex, 37,38 12-3819. 19. Scott, D.S.; Piskorz, J. (1 982) The Flash Pyrolysis of Aspen-Poplar Wood. Can. J. Chem. Eng,, 60,666-674. 20. Scott, D.S.;Piskorz, J.; Radlein, D.(1985) Liquid Products from the Continuous Flash Pyrolysis of Biomass. Ind. Eng. Chem. Process. Des. Dev., 24,581-588. 21. Al-Ubaid, A.; Wolf, E.E. (1988) Steam Reforming of Methane on Reduced NonStoichiometric Nickel Aluminate Catalysts. Appl. Caful.,40, 73-85. 22. Garcia, L.; Salvador, M.L.; Bilbao, R.; Arauzo, J. (1998) Influence of Calcination and Reduction Conditions on the Catalyst Performance in the Pyrolysis Process of Biomass. Energy&Fuels, 12, 139-143.
357
The effect of alkali metal on the catalytic gasification of rice straw over nickel catalysts supported on kieselguhr Seong BO Kim', Seun -WOO Lee', sang-Sung N ~ I ,K ~ U wan Lee' and Cheong Song Choi Catalyst Research Division I , Korea Research Institute of Chemical Technolou, PO.Box 107, Taejon 305-600, Korea Department of Chemical Engineering, Sogang University, Seoul 121 - 742, Korea
'
8
Abstract Rice straw was catalyticallygasified over nickel catalysts supported on kieselguhr. This has been done by varymg the content of alkali carbonate, lithium metal (5-20wt%) and various sodium compounds. In the case whlch alkali metal carbonates were separately added with nickel catalyst, conversion to gas was increased in the following order of Li< Cs< K5Na because of the lowest solubility of lithium. However, alkali metals were used to as co-catalyst by lmpregnation method, gas formation was increased in the following order; Cs< K g a < Li. These results showed same aspects with TPR patterns.
INTRODUCTION Thermochemical conversion of biomass by pyrolysis, gasification and liquefaction of biomass into gas or liquid fuels has been receiving a recent attention as fossil hydrocarbon supplies are dwindled. Biomass like rice straw, wood and etc. are chemically composed of three different components of cellulose, hemicellulose, and lignin. Thus, thermochemical products were very different with feed stocks. Oil and solid obtained through carbonization and liquefaction contains many nitrogen and sulfur compounds which cause environmental problems. In the mean while, gasification produces a clean synthesis gas which contains a high concentration of hydrogen. Therefore, it may be a better method than liquefaction and carbonization. Various catalysts were examined to increase the products obtained from thermochemical reaction of biomass. Appell and co-workers' demonstrated that solid organic materials, including urban refuse, agricultural wastes, wood and bovine manure, can be converted directly to heavy fuel oil in the presence of a catalyst such as sodium carbonate. Since then, a great number of works have been reported on the direct liquefaction of biomass in the presence of suitable c a t a l y ~ t s ~Recent ~~. publication, the addition of Na2C03 greatly increased the reaction rate. Minowa et
358
al?' gave a detailed description on the cellulose liquefaction using Na2C03catalyst in hot compressed water at different temperatures and established a chemical mechanism based on the product distribution. Jean Claude et a1.* reported that obtained oil yield was more than 50% by the pyrolysis of bitumen with KzCO3, Li2C03,Na2C03at 350-4000. Elliott et al.' reported that nickel catalyst showed an excellent activity on the gasification. Also, Lee et. a11I-'' demonstrated that main reaction course to produce gas from rice straw was the formation of gas from oil step and alkali carbonate like Li2C03,Na2C03,K2CO3 activate the formation step. Addition of alkali metals can allow a timely tuning of their catalytic alkali additives, acting as either structural of chemical b e h a ~ i o u r ' ~ -In~ ~particular, . promoters, are widely used to enhance the activity of FeRu ammonia synthesis catalysts'2"8 and the selectivity to heavier hydrocarbons of Ni catalysts in COhydrogenation.12,'3918-23 Furthermore, alkali doping markedly affects the reactivity pattern of supported Ni catalysts in the hydrogenolysis exchange and hydrogenation r e a ~ t i o n s . ' ~ "These ~ * ' ~ effects of alkali promoters on the behavior of heterogeneous catalysts have been associated with an electron enrichment of metal particles (electronic effect) whch influences their interaction paths with reactantlproduct molecules.12*18-22 Besides, alkali compounds are currently added to Ni steam reforming catalysts to inhibit cockmg and sulfur- oisoning phenomena, though the 8 3 mechanistic action of promoters is still undecided. 1 In this study, we focused on the effect of alkali metal over nickel catalyst. Also, effect of alkali metal co-catalyst was carried out in the gasification of rice straw.
EXPERIMENTAL EXPERIMENTAL PROCEDURES The gasification of rice straw was conducted in a conventional stainless steel autoclave (2L) as shown in Figure 1. Autoclave had a stirrer with a magnetic drive system and a gas reservoir. Nitrogen was used as purge gas. Additional nitrogen was added to achieve the initial pressure. Rice straw (50g), catalyst (5g) and distilled water (500ml) were introduced into the autoclave and heated with an electric heater at 300°C for 30min. Then, reaction mixtures were intensively extracted with dichloromethane. After this, the mixtures were allowed to stand for several hours and filtered. The cake on the filter glass was dried at 105 OC for 5hrs. Finally, the solid residue was obtained. The dichloromethane was recovered from the filtrate of extract using a rotary evaporator (Buche RE 111).
p a E chro matag rap hy
N2 g a s
Figure 1. Experimental scheme for the gasification of rice straw
359
Dark brown heavy oil and solvent soluble material were obtained in this manner. The amount of formed gas was measured with a gas meter and the composition of evolved gases was determined by gas chromatography with TCD (silica gel column of 60/80 meshes). Conversion to gas, oil and carbonized solid were calculated by the following equation: Weight of material obtained fiom each phase x 100 Conversion (%) = Weight of drymg biomass The moisture content of rice straw was measured after drymg at 105 "C for 5 hrs and the measured value was 8 %. Ash was analyzed after heating at 600 "C for 5 hrs and its value was 4 % (ASTM D1102-84). PREPARATION OF CATALYSTS
The nickel catalyst (about 50 wt% nickel on kieselguhr) was prepared by an ordinary precipitation method. Sodium carbonate solution was added to a slurry of kieselguhr and nickel nitrate solution at 70 "C and precipitate was obtained. This precipitate was washed with water thoroughly and then was dried at 105"C for 12 hrs, crushed to 60150 mesh, calcined at 350°C for 4 hrs. This was activated with 100% hydrogen at 200, 300 and 350 "C for 4 hrs. These prepared catalysts were stored in nitrogen atmospheric bottle and desiccator. CHARCTERIZATION OF C A T L Y S E REDUCTION (TPR)
TEMPERATURE PROGRAMMED
For TPR studies, 50 mg of the calcined catalyst were loaded in a quartz reactor and heated at 673K for 6 h, followed by cooling down to room temperature in argon gas. A high purity premixed gas containing 95% argon and 5% hydrogen was used as reducing agent. Traces of oxygen and water vapor were removed by passing the gas through an activated molecular sieve trap kept in an ice bath. The catalyst bed was heated linearly at 5Wmin. A continuous TPR response profile was obtained when the difference in H2 concentrations in the gas streams between reference and sample, resulted fiom the reduction of metal oxide by H2. RESULTS AND DISCUSSION EFFECT OFALKALI METAL CARBONATE
The effect of alkali metal on the formation of gas, oil and carbonized solid was examined in the catalytic gasification of rice straw over nickel catalyst. Various alkali metal carbonates were separately added with nickel catalyst and the results were presented in Table 1. Addition of alkali metal carbonates gave a large influence on conversion to gas and composition of gas. The formation of gas was dominated and increased in the following order; Li< Cs< KSNa. In case of lithium, conversion of oil, char and gas shows similar to without catalyst. This result caused to a solubility of lithium. The solubility of lithium24is the lowest than that of other carbonate.
360
Table 1. Effect of various carbonates on the gasification of rice straw.
Various sodium compounds were compared and shown in Table 2. Only Na2C03 showed an excellent conversion. But, other compounds including strong alkali compound represented a lower value. This result suggests that strong basicity of the catalyst is not required in gasification of rice straw. Table 2. Results of various components on the gasification of rice straw.
THE EFFECT OFAMOUNTS OF Na2C03
First, the effect of amounts of Na2C03 on this reaction was investigated. Figure 2 shows that gas yield was gradually increased until 1.0 weight of Na2C03per nickel and carbonized solid was sharply decreased. Figure 3 shows the addition of Na2C03 to nickel on kieselguhr largely increased the formation of carbon monoxide. c, 3 Y
c 0
.A-
:?
(u
>
S 0 U
70 bO 50
40 30 20
LO n Ratio o f Na2COYNickel
Figure 2. The effects of amounts of Na2C03on the gasification of rice straw
361
700
500
1
O
O
1.4
2
0
6oo 400 -
0
E 300 200 -
--*
100 -
0
o
o
o
.2
.4
.6
.a
1
Ratio of Na,CO,INickeI
Figure 3. Composition of formed gases depends on amounts of Na2C03during catalytic gasification of rice straw We also try to understand the role of nickel. Two catalyst systems, Na2C03and Nikieselguhr + Na2C03, were compared in this gasification. As shown in Figure 4, Na2C03 increased the formation of oil and the nickel promoted the formation of gas products. This result suggests that NazC03promotes the oil formation step and nickel increases the gas formation from oil and char.
0
-
.d UI s L . a Q c -3 > 0 V
.Char
A0 60 40
G as
20
0
200 250 300 N i/kkselguhr+ N a2C 0 3
200
250
300
N a2C 0 3
Tern perahrre(oC 1
Figure 4. Effect of reaction temperature and Na2C03on the gasification of rice straw ROLEAS CO-CATALYST OFALKALI METAL
Also, the role of alkali metal as co-catalyst was examined. Various alkali metal nickel catalysts were prepared by impregnation method (5wtOh) using alkali metal carbonates such as Li2C03,NaZCO3,K2C03,Cs2C03and obtained result were shown in Table 3.
362
Table 3. Results of various nickel catalysts prepared by impregnation method using different alkali carbonates.
I
1
catalyst
I
LiNi/kieselmhr NaNi/kieselguhr KNi/kieselguhr CsNfieselrmhr
I
I I
Composition of gas (mmol) Conversion to H2 co CH, gas (wt.%) 45 I 874 I 167 I 6 115 Trace 33 648 137 Trace 34 228 28 I 180 I 101 I 1.2
coz I 1
359 271 298 251
I I
Conversion to gas was increased in the following order; Cs< K-
Figure 2. TPR profiles of alkali metal supported on kieselguhr catalyst.
Also, effect of amounts of lithium loaded on nickel catalyst was studied and conversion to gas had a maximum value at 5 wt.% per nickel as shown in Table 4. Table 4. Effect of amounts of lithium loaded on nickelkieselguhr ~~
~
Composition of gas (mmol)
Amounts of
Lithium(wt%)
Conversion to gas (wt.%)
0 5 10 20
28 45 37 35
H2
CO
CH4
COZ
107 874 412 385
207 167 109 102
Trace 6 1.2 3
158 359 300 288
CONCLUSIONS 1.
2.
Addition of Na2C03 enhanced activity for catalytic gasification of rice straw over nickel catalyst and largely increased the formation of gas. Na2C03 activates the oil formation step and nickel catalyst increases the gas
363
3.
4.
5.
formation fiom oil and char. The formation of gas depended on the nature of alkali metal carbonates and is increased in the following order of Li< Cs< KlNa because of the lowest solubility of lithlum. The formation of gas was increased in the following order; Cs< KlNa< Li when nickel-alkali metal catalyst prepared by impregnation method using alkali metal carbonate such as Li2C03, Na2C03, K2C03, Cs2C03 were used in the gasification of rice straw. The TPR pattern of nickel-alkali metal catalysts were moved to lower temperature fiom Cs to Li and showed same aspects with reactivity.
REFERENCES 1. Appell, H. R.,Fu, Y. C., Friedman, S., Yavorsky, P. M., and Wender, I.,( 1971) U.S. Bureau of Mines, Pittsburgh. 2. Boocock, D. G B., Mackay, D., Mcpherson, M., Nadeau, S.and Thurier, R.,(1979) Can. J. Chem. Eng., 57,98. 3. Boocock, D. G B., Mackay, D., Franco, H. and Lee, P., (1980) Can. J. Chem. Eng., 58, 446. 4. Walton, T. E. and Paudler, w. w., (1981) Fuel, 60,650. 5. Eager, R. L., Mathews, J. F. and Pepper, J. M., (1982) Can. J. Chem. Eng., 60,289. 6. Minowa, T., Fang, Z., Ogi, T. and varhegyi, G., (1997) J. Chem. EngJapan, 30, 186. 7. Minowa, T., Fang, Z. and Ogi, T., (1998) J. of Chem. Eng.Jupun, 31, 131. 8. Jean-Claude, C. and Esteban, C., (1978) Fuel, 57,304. 9. Elliot, D. C., Butner, R. S. and Sealock, L. J. Jr., (1988) Res. Thermochem. Biomass Convers., 696. 10. Lee, S.W., Kim, S.B., Lee, K.W. and Choi, C. S.,(2000) Korean J. Chem. Eng., 174. 11. Martin, G.A., Imelik, B., Naccahe, C., Coudurier, C., Praliaud, H., Meriaudeau, P., Gallezot, P. and Vedrine, J.C.,(Eds.), (1982) Studies in Surface Science and Catalysis. Metal-Support and Metal-Additive Effects, vol. 11, Elsevier, Amsterdam, 3 15. 12. Mross, W.D., (1983) Catul. Rev. Sci. Eng., 25(4), 591. 13. Praliaud, H.,Dalmon, J.A., Mirodatos C. and Martin,G.A., (1986) f. Cutal., 97, 344. 14. Chai, GY. and Falconer, GJ.L., (1985) J. Cutal. 93, 152. 15. Praliaud, H., Primet, M. and Martin, GA., (1983) Sui$ Sci., 17, 107. and Falconer, J.L. (1989) Appl. Cutul. 50,I89 16. Campbell, T.K., 17. Dean, J. A,, (1992) Lange's Handbook of Chemistry.
364
Hot Gas Filtration via a Novel Mobile Granular Filter *Nicolas AbatzoglodP2, Martin Gagnon' and Esteban Chornet', 2* 3: Universitk de Sherbrooke, Dept. of Chemical Engineering, Sherbrooke (Qukbec), Canada; Enerkem Technologies Inc., 4245 Garlock, Sherbrooke(Qukbec), Canada; NREL, Golden, Colorado, USA
'
ABSTRACT: The existing technologies for hot gas-fine particles (< 10 pm) flow separation include, electrostatic precipitator, ceramic, metallic and granular filters. The latter are of interest because of their flexibility and strength at high pressures, high temperatures and in corrosive environments, while maintaining h g h efficiency over time. However, granular filters have also their drawbacks: (a) fixed bed configurations are pressure-dependent in their operation because of the formation of a surface layer at the top of the filtering media; (b) conventional moving beds require a relatively large quantity of filtering media moving over a distance which produces erosion on the collector vessel and pipes. To keep the advantages while minimising the drawbacks of granular filters, a configuration of novel mobile granular bed filter (NMGBF) has been designed, scaled-up and applied at pilot-scale as one of the modules of a fluidised bed gasifier hot-gas-conditioning-train. The filter's operation whch follows Ergun's and Darcy's laws, is characterised by a quasi-fured bed regime of the filtering media, coupled with recirculation through a central section of the filter operated as a fluidised bed. This configuration prevents the build-up of a surface layer and consequently, renders the filtering process less pressure-dependent, more compact, less costly to operate and techcally more reliable over long filtration periods compared with other technologies. This paper reviews the operating conditions of existing technologies in comparison with the NMGBF. It also describes the NMGBF characteristics including the concept, the basic model development and typical real time cold and hot gas filtration results. INTRODUCTION As early as 2000 b.c., the Chinese used fixed granular beds for the filtration of wine. During the 20" century, granular bed filters have been widely used in the industry for cold and hot dry gas filtration.
365
A typical case is the particulate removal from hot synthetic gas producer to be used in gas turbines aiming at the production of electricity through integrated gasification combined cycles. The particles carried over with the synthetic gas constitute a major technological problem since they are responsible for both chemical and mechanical erosion in turbines. It is now widely accepted that outlets particulate concentrations must be in the range of 1 to 20 PPM to meet both the environmental emission standards and the estimated tolerances for gas turbines I*'. At present, there are four major existing technologies that have shown potential to reach these performances in hot gas filtration: electrostatic precipitators (ESP), ceramic, metallic and granular filters.
LITER4 TURE BACKGROUND Hot and dry gas cleaning technologies Curtiss-Wright has studied the application of the ESP to PFBC (pressurised fluidisedbed combustion) and gasification gases with nine cylindrical electrodes providing a collecting area of 39 m2, under a pressure of 570 kPa and a temperature of about 700°C'. The Coal Mining Research Centre (CMRC) and Kawasaki Heavy Industries of Japan have jointly developed a two stage moving granular collector, shown in Figure 1, in connection with a 40 t/d fluidised bed gasifier under operating conditions of 1.8 MPa pressure and 450 "C for as long as 900 hours Porous sintered ceramic candle filters have been tested at 720-1050 H a , at about 800 "C and for a collection Regarding duration of 600 hours at the Grimethorpe PFBC Establishment in UK the metallic felt filters, a system comprising four filter modules containing some 300 bags made of Inconel 601, covering a total filtering area of 430 m2 have been operated at RWTH Aachen Our NMGBF has been tested under atmospheric pressure, at a temperature of 400-500 "C, a flowrate of 20 Nm3/h and an upstream particle load of 3000-5500 mg/Nm3. The NMGBF is depicted in Figure 2. Table 1 is a particle collection efficiency comparison between 6 filter technologies. To complete the overview of these 6 technologies, Table 2 displays their main advantages and their drawbacks
',
'.
'.
'.
Fig.2 NMGBF shown without
Fig. 1 Kawasaki moving granular. bed filter
'
filtering media 366
'.
Table 1 Efficiencies and operating conditions comparison between six filter technologies for hot and dry gas cleaning
’.
No. Filtration technologies 1 2
ESP C a) Porous candle e r a b) Fibre m i
T (“C)
Efficiencies (%)
790-850 776-862
Operating~-pressures (Wa) 540-640 720-1050
200-800
101.3
99.9 +
260-300 0 - 870
360 101.3
99.9 95 - 99.99
0 - 870
101.3
94 - 99.99
95-99.5 99.4-99.7
C
3
4
Metallic G a) Fixedbed r a n u b) NMGBF 1 (This work) a r
The above technologies could find their commercial operation as gasparticles separation for gasifiers, combustors, cement industry, metallurgical ovens, industrial boilers and refineries. Table 2 General characteristics of existing hot and dry gas cleaning technologies Filter types ESP
Filtration principles Electromagnetic force
Ceramic
Inertial impact
Advantages Low pressure drop Low maintenance frequency High gas-particle filtration flowrate Low operating costs Continuous operation
0
0
Effective over time Effective with low particle diameter Reliable Relatively insensitive to gas stream fluctuation No corrosion or rusting of components
367
’* ’. 4,
Drawbacks Requires electricity Efficiency goes down with the increase of the particle resistance Temperature dependent High capital cost Production of ozone during gas ionisation Large filtration space required Ceramics are costly and fragile High maintenance cost Corrosive agents sensitive Large dimensions required Temperature dependent Plugging of the fabric m a y occur due to gas adhesive components
Table 2 continued High maintenance cost Effective over time Large dimensions required Effective with low Temperature dependent particle diameter Plugging of the fabric may Reliable occur due to gas adhesive Relatively insensitive components to gas stream fluctuation Filtration efficiency Corrosive agents Granular Inertial decreases for submicronic insensitive impact particles High working temperature Low pressure drop Wide operating pressure range Low operating cost Many granular can be used-as filtering media Note: In all cases it is considered that particles of 10 microns and over are precollected by cyclones before the main filter. Metallic
Inertial impact
Moreover, specifically for the granular bed category, Table 3 makes a detailed comparison of characteristicsbetween three different filter types.
NMGBF CHARACTERISTICS Even if granular beds have several advantages compared with other technologies, they have also drawbacks: fixed granular beds are pressure-dependent because of the formation of a surface layer at the top of the filtering media, while conventional moving granular beds require large quantity of filtering media moving relatively over long distances producing erosion on the collector vessel and pipes. The NMGBF has been conceived on the basis of keeping all advantages of granular beds, whle minimising these drawbacks. The NMGBF is operated on a slow moving granular bed over short distances. This renders the NMGBF's operation less pressuredependent and keeps erosion to a minimum level; moreover, the filter is more compact, less costly to operate and technically more reliable.
GEOMETRY The NMGBF prototypes are composed of two concentric cylinders, side exits and a conical metallic base as shown in Figure 2 (hot pilot-scale filter) and 3 (cold mock-up bench test). The filtering media is located in the annulus between the 2 concentric cylinders. The gas to be cleaned enters the filter at the bottom through tuyeres becoming a distributor plate. The internal cylinder is divided in two parts. The gap between these two parts allows the filtering media to flow slowly back to the lower part of the internal cylinder; the granular media is then entrained by the gas flowing upward and closes the recirculation loop. A grid is located at each exit to prevent the filtering media to be entrained but allows the gas and the particles to flow freely out the of the filter.
368
Table 3 Comparison between different granular bed filter types Filter types Moving bed
Fixed bed
NMGBF (mock-up at ambient temp.)
*
Advantages Continuous self-cleaning property Captured dust easy disposal Stable high efficiency Adaptability to load changes Simple Low cost
Less pressure dependant Low cost Simple Easy to regenerate (long life cycle)
Drawbacks High operating pressure Collector vessel and pipes erosion due to filtering media long circulation distance Formation of a surface layer Pressure dependent Difficult to regenerate (cake at the surface layer) Efficiency between 84 to 96% for submicronic particles
‘, ’. 43
Examples of operating conditions Filtering media diameter of 2to6mm Downward moving speed of the filtering media of 10 mh Gas flowrate 180 Nm3/hr/filtering section (m2) Particles to be filtered: 2.02 Pm Efficiency 99% and over Gas flowrate 1200 Nm3/hr/filtering section (m2) Max. pressure drop before cleaning or regenerating the media: 60 mbars Optimum filtering height: 35/40 mm Particles to be filtered: submicronic to 10 pm Efficiency 99% Particles to be filtered: submicronic to 10 pm Downward moving speed of the filtering media: 0.025 to 0.04 rdh * Gas flow rate 320 to 800 Nm3/hr/filtering section (m2) Efficiency 99.9% for particles higher than 1 micron
Based on preliminary results, detailed calculation in subsequent paper
The NMGBF has similarities to two filter types: the spouted bed and the fluidised bed filter. Geometrically the NMGBF is a spouted bed with a draft tube in the central spout jet region as shown in Figure 3. The presence of the draft tube helps to control the quantity of solids falling in the internal cylinder thus reducing the flowrate, to ensure the fluidisation conditions, minimises turbulence and pressure drop across the filter. This configuration gives a dilute phase entrained solids in the jet region and a slow moving packed bed in the annulus. This operation prevents from building-up a surface layer and it homogenises the particle concentration in the filtering media. The operating conditions (flowrate and the maximum entrained height) of the NMGBF are similar to those of the fluidised bed 6q’.
369
E x t e r n a l Cylinder I n t e r n a l Cylinder
F i l t e r i n g Media
ids Gop
Falling Solids Cone
Lower I n t e r n o l Cylii? d e r Solids Flushing Pipe
Conical Base
Inlet
Fig. 3 NMGBF mock-up. ANALYSIS AND SIMULATION OF FILTER 'SOPERATION
The gas fed in the filter through the distributor plate has two ways to reach the side exits: either to change direction and pass through the slot of the internal cylinder or to continue vertically through the internal cylinder, reach the sealing of the filter, change direction and flow down through the filtering media. The filtering media collects the particles mainly through the deep bed impact mechanism towards the side exits. If the gas flow along the internal cylinder is high enough solids are fluidised and, depending on the internal cylinder height, they are sprinkling out. In parallel, new solids are falling into the lower part of the internal cylinder to be entrained with the gas. This regime is characterised by a periodic oscillating entrained solids movement of the filtering media in the annulus preventing the build-up of a surface layer and keeping pressure drop low. For a better understanding of the NMGBF behaviour Figures 4, 5 and 6 show a CFD (done on software Fluent) two phase flow simulation with sand as filtering media ( d h 500 pm, ph 2560 kg/m3) in the annulus and air at ambient temperature at a flowrate of 50 Nm3/h (800 Nm3/hr/m2).in steady state. Figure 4 shows the filtering media falling in the lower internal cylinder plugging the gas flow. The pressure rises, the weight of the filtering media in the internal cylinder can not balance the upward gas velocity force, resulting in an upward chaotic expansion and carry-over of the sand (Figure 5). Finally, the sand settles down at the top of the filtering media in the annulus (Figure 6 ) and the process starts again. The frequency of this oscillating movement depends strongly on the gas flowrate and its corresponding superficial gas velocity and the falling solids gap and the properties of the solids. In this way, the filtering media in the annulus is slowly moving downward at the same rate as the sand falls down from the annulus area to the central spout jet region. If particles enter at the bottom with the gas flow, the slow moving filtering media in the annulus would collect them m a d y by inertial impact mechanism. 370
TO WARDS A SEMI-EMPIRICAL MODELLING AND SCALE-UP For a successful filtration process, the NMGBF must fulfill the operating criteria described below:
Entrainment of particles To ensure an entrained fluidised bed operation in the concentric slotted internal cylinder, the gas flow velocity must be high enough to obtain at least the minimum fluidisation velocity of the solids. This can be calculated through equations l a and b (Thonglimp) '.
I "
Fig.4 Plugging period
y. 5 Expansionperioc
the solids.
u mf
=
the solids.
'ig. 6 Falling period of the
solids.
Remf *Pf Dp*Pf
The total distance to travel in the internal cylinder is given by the minimum fluidisation height of the solids and can be approximated through equations 2 (Ergun) and 3 '.
'
Re GaMv = 1.75mf+ 3 V&tnf &
mf
=l-
(I--Emf 150-
2 3 Y 'mf
1
Re
mf
M PSSH*f
(3)
37 1
Avoid bridging in the slotted internal cylinder To ensure regular operation, solids bridging in the gap between the two internal cylinder parts must be avoided. Generally, it occurs when the kinetic pressure due to high gas cross-flow becomes higher or equal to the radial component of pressure due to weight of the filtering bed but also due to the chaotic solid-solid particles interaction forces. This phenomenon is a function of the granular media particles sphericity, size and surface tension and cohesive properties. To decrease the importance of bridging, the gas velocity must be kept low. Besides, the gap between the two internal cylinder parts must be large enough to minimise the importance of the bridging forces by keeping the ratio (weight of filtering media) / (gas kinetic force) as high as possible
'.
Maximum efficncy A sufficient bed length must be kept to insure an appropriate contact time (at least 0.1 to 0.5 s) between the filtering media and the gas-particle flow for an optimum particle collection efficiency 4, '. The above criteria led to the NMGBF mock-up and pilot-scale dimensions reported in Figure 7 and Table 4.
4& LDF
t
--i
I
I
FH
CBA
Fig. 7 Dimensions of the NMGBF.
372
Table 4 NMGBF dimensions in function of operating conditions parameters. Variable
NMGBF mock-up dimensions
NMGBF pilot-scale dimensions
0.3 0.1 0.1 0.65 0.9 0.1 0.1 0.135 60 320-800 320-800 1000-3000
0.575 0.0765 0.05 0.725 0.7 0.15 0.1 0.125 60 40-80 80- 160 2000-5000
Filter Dia. (DF) (m) Internal cylinder Dia. (DC) (m) Exits Dia. (DE) (m) Cylinder height (CH) (m) Filter height (FH) (m) Filtering media height (FMH) (m) Falling solids gap (FSG) (m) Exits height (EH) (m) Conical base angle (CBA) (degree) Standard Gas flowrate (Nm3/hr/m2) Real Gas flowrate (m3/hr/m2) Particles load (PPM) FLOWRATE SPLIT-UP
The flowrate across a granular media is proportional to the resistance of this media according to Darcy's law (eq. 4a and b) and is represented in Figure 8. This system can be analysed using an electrical resistance analogy. In the same way as in resolving an electrical resistance system based on Ohm's law (eq. 5) the total resistance of the filter can be calculated (eq. 6).
AV = RI
1 Rtot = (-+Rla
(5)
1 1 + ) -l Rlb R2+ R3a*R3b R3a+R3b
The total operating pressure drop of the filter can be approximated by equation 7. Knowing that the pressure drop across the two pathways is equal to the total pressure drop, the flowrate through each pathline can be calculated as follows: If Rla = Rlb then:
373
Qla = Q l b =
AP Rla*,uf
*A,
AP
Rlb*pf
*A
If R3a = R3b and R2=0 then:
1 AP * A = AP * A -Q2 = Q3a = Q3b = 2 R3a*pf R3bP f 2 2 In our filtration system, the gas flow splits-up proportionally resistance encountered across the two possible pathways.
(9)
3
the ombined
Ii - t
Fig.8 Filter resistances schematic. PRESSURE DROP PREDICTION Figure 9 reveals the typical NMGBF pressure drop over the variation of the gas superficial velocity approximated by Ergun's law under fixed bed conditions, reported in equations 10a and lob, for an upward gas flowrate ranging from 0 to 50 Nm3/h (0 800 Nm3/hr/m2)through Ottawa Sand as filtering media using the NMGBF prototype dimensions and no particles injection 6* *. The NMGBF's minimum fluidisation velocity is 0.22 d s according to the solids and fluid physical properties using equation la. When fluidisation conditions occur, the constant pressure drop over the variation of the flowrate is calculated through equation (1 1) 6p *.
374
5.00E+02 4.50E+02
e 4.00E+02
-p! 3.50E+02 Q)
5m g
3.00E+02 2.50E+02
(D
2
U
f
2.00E+02 1.50E+02 1.00E+02
e! a 5.00E+Ol O.OOE+OO 0
0.5
1
1.5
2
2.5
Superficial gas velocity (mls)
Fig. 9 NMGBF pressure drop over the superficial gas velocity (eq. 10a, b & 11). TYPICAL RESULTS COLD MOCK-UP Figure 10 reports the filtration efficiency of the particles naturally contained in air at ambient temperature under the NMGBF prototype operating conditions mentioned in Table 4. Figure 10 shows that the gas-particle separation efficiency for submicronic particle is comprised between 54 to 96 %. Submicronic particles removal efficiency is the major limitation of the NMGBF as for granular bed filters. With particle diameters ranging from 1 to 10 pm, the efficiency reaches nearly 96%. 100 -
- -
---
1
-*-
95
90 85
80 -
55
75
5
70
I
65 --
60 5550
+
7
Fig. 10 Filtration of particles contained in air.
375
I
.
_ I."
"
.l...l_lll........
......
""
! !
88 -
,
86 -
\
j i
HOT REAL-TIME Pilot-scale particle collection efficiency has been found to be similar to the cold and dry experiments over 300 hours filtration time. Figure 12 presents a relatively constant pressure drop over time for the hot and dry experiments conducted on the pilot-scale filter under a gas flowrate of 20 Nm3h (80 Nm3/hr/m2),with a particle load of 3000 mg/Nm3 and using Ottawa sand as the filtering media. Deep holes on Figure 12 are air pulses to back-flush the solids plugging the exits. 9 8
.--
7
h
b
6
a
e
5
-
4
; -
3 2
f
&
1 0
-1
00
Tim e
(s.)
Fig. 12 NMGBF pilot scale pressure drop over time.
376
CONCLUSIONS A novel mobile granular bed filter has been conceived to be applied to industrial operations in order to meet emission control standards. It is simply composed of two concentric cylinders, a conical base and side exits. Its geometry is similar to a spouted bed with a concentric draft tube operating as a typical fluidised bed. The filter's operation is characterised by a recirculating between an entrained bed and a quasi-fixed bed regimes. The downward filtering media velocity in the annulus prevents the build-up of a surface layer and renders the operation less pressuredependent over long filtration periods thus allowing a better control efficiency and lower operational costs compared with other technologies. The solids circulation in the internal cylinder under a fluidised dilute phase regime keeps the turbulence and the pressure drop at a minimum level. Thonglimp's equation estimates the minimum particle fluidisation velocity and the particle entrained height and Ergun's law approximates the pressure drop across the filter. Darcy's law predicts the gas flowrate distribution through out the filter. The filter showed excellent efficiency (99.9%) for both the cold- and hot-dry experimental work for particles of more than 1 micron. The filtration efficiency is slightly less than 96% for submicronic particles. The major impact of the NMGBF is the possibility of ensuring stable pressure regimes over long filtration periods thus allowing better control, efficiency and lower operational costs
ACKNOWLEDGEMENT This work was financially supported by the Natural Sciences Engineering Research Council (NSREC) and Enerkem Technologies Inc.. Experimental work was performed at the facilities of the GRTPC (Groupe de Recherche sur les Technologies et ProcCdCs de Conversion) and the department of Chemical Engineering at UniversitC de Sherbrooke. Special acknowledgements are addressed to J. Bureau, A. Mincic and H. Gauvin for technical assistance.
NOMENCLATURE A CFD DP Dfm ESP g Ga Hnlf I
L Mv M NMGBF PFBC PPM
Filter cross-sectional area Computational fluid dynamics. Particle diameter. Filtering media diameter. Electrostatic precipitator. Gravity Galileo number (D:pf2g/pz). Bed height at minimum fluidisation. Current. Distance to travel. Density ratio ((pp-pf)/pf). Mass of solids in the system. Novel mobile granular bed filter. Pressurised fluidised-bed combustion. Particle per million. 377
Flowrate. Resistance. Reynolds number at minimum fluidisation (u,,,dIpp&f). Cross-sectional area of the column (vector in the sense of flow). Temperature. Superficial fluid velocity at minimum fluidisation. Superficial fluid velocity (measured in the empty column). Interstitial fluid velocity. Solids velocity. Permeability of the filtering media. Potential difference. Pressure drop. Fluid voidage. Solids voidage. Bed voidage at minimum fluidisation. Viscosity. Fluid density. Solids density. Filtering media density. Solids sphericity.
Q
R Red S T u m f
UO Vf VS
P AV
AP Ef ES
Emf
Pf Pf Ps PSn
w REFERENCES 1. 2. 3.
4.
5. 6. 7. 8.
Takamatsu T., Maude C. (1991) Coal gasification for ZGCCpower generation, IEA Coal research, IEACW37, ISBN 92-9029-190-7. NOVEM (1995) Development of standard procedures for gas quality in biomass gasifier-producer generation systems, JOU2-CT93-0408, contract number 355200/2130, p.118. Abatzoglou N., Chornet E., Bureau J., A. Mincic (1999) A Mobile Granular Bed Filtration Apparatus for Hot Gas Conditioning, Canadian Patent Application, No 2,268,376 (April 7, 1999), Spanish patent Application, No 9901711 (July 27, 1999). Dumon R., Joffie R. (1984) Dt?poussit?reursindustriels, Mason, ISBN 2-22580065-0, p. 92-107. Enfor project C-258 (1983) A comparative assessment of forest biomass conversion to energyforms, phase 1, volume VII, 151 p. Davidson J. F., Clifi R., Harrison D. (1985) Fluidization, Academic press Inc., ISBN 0-12-205552-7, p. 733. Tsubaki J., Tien C. (avril 1988) Gas filtration in granular moving beds- an experimental study, Canadian Journal of Chemical Engineering, vol. 66, p 271275. Gidaspow D. (1994) Multiphase flow andfluids, Continuum and kinetic theory descriptions, Academic press, ISBN 0-4 12-09600-9,p. 467.
378
Design Of A Moving Bed Granular Filter For Biomass Gasification R. C. Brown, J. Smeenk, and C. Wistrom Center for Sustainable Environmental Technologies, Iowa State University,Ames,IA 5001I USA
ABSTRACT We are developing a moving bed granular filter suitable for dry scrubbing product gas from a fluidized biomass gasifier. The filter is intended to remove particles from producer gas with high efficiency and low pressure drop. Additionally, tar and other contaminants including alkali may be reduced depending on filter operating temperature and media composition. Moving bed granular filters operate on the principle that a flowing bed of particles can effectively scrub particulate contaminant from a gas stream. Although very promising for achieving high filtration efficiencies, the relatively large footprint of the equipment and high throughputs of granular material as filter media are ofien cited as drawbacks to moving bed granular filters. We are exploring mechanical changes in moving bed granular filters to overcome their present drawbacks. In this paper we describe three fluid dynamic design features that improve performance of the filter. Cold flow testing has demonstrated that the new design increases gas flow throughput and decreases gas pressure drop through the filter compared to traditional designs.
INTRODUCTION We are developing a moving bed granular filter suitable for dry scrubbing product gas from a fluidized biomass gasifier. The filter is intended to remove particles from producer gas with high efficiency and low pressure drop. Additionally, tar and other contaminants including alkali may be reduced depending on filter operating temperature and media composition. Moving bed granular filters operate on the principle that a flowing bed of particles can effectively scrub particulate contaminant from a gas stream. Although very promising for achieving high filtration efficiencies, the relatively large footprint of the equipment and high throughputs of granular material as filter media are cited as drawbacks to moving bed granular filters. We are exploring mechanical changes in these filters to ameliorate these problems. A counter-flow arrangement establishes a dust cake at the interface between the gas and granular bed. We hypothesize that the dust cake, if carefully maintained at optimum thickness, will enhance filtration
379
efficiency. However, the counter-flow of gas and granules and the momentum of gas above the dust cake introduce challenges in the fluid dynamic design of the filter. This research is being performed in two phases: evaluation of pressure drop and flow characteristics of cold flow models of the filter and collection efficiency tests with a pilot-scale filter coupled to a 4.5 tonne per day biomass gasifier. In this paper we describe three fluid dynamic design features developed in the cold flow model that improve performance of the filter. BACKGROUND Recent analyses suggest that barrier filters (ceramic candles) and granular bed filters are the most promising approaches to hot-gas clean up for power systems [l, 2, 31. Granular filters are especially attractive for biomass power applications because, unlike ceramic candle filters, they are not susceptible to degradation by alkali. Furthermore, granular filters offer the prospect of simultaneous removal of particulate matter and trace contaminants. Potential disadvantages of granular filters include acquisition of a suitable filter media, disposal or the filter media, and/or renewal of the media. Squires and Pfeffer [4 ] were among the first to consider the use of granular beds for control of fly ash emissions. A louvered panel arrangement held 16-30 mesh sand in a fixed bed. Reported collection efficiencies were as high as 99.8%. Lippert and coworkers [5] reported collection efficiencies of essentially 100% for fixed beds operated at superficial velocities less than 0.4 d s . Significantly, they attributed these outstanding results to the formation of a dust cake at the surface of the beds, a result confirmed by tests with Plexiglas models operated at ambient conditions. It was hypothesized that dust bridges the gaps between individual media granules and the collection mechanism shifts from interception deep within the bed to impaction at the freeboard-bed interface. Both collection efficiency and pressure drop of fixed granular beds increases as contaminant builds up on the bed. Cleaning usually entails reverse flow of gas similar to the process for cleaning bamer filters. Moving bed granular filters overcomes the problem of periodic cleaning. The use of moving beds dates back to the 1940's [6]. Some of the earliest designs employed cross-flow configurations. Granular media flows downward by the action of gravity through a section enclosed by screens or louvers. Gas flows horizontally from a dirty gas plenum through the vertically flowing granules, which intercept the particles, and exits into a clean gas plenum. The Dorfan Impingo filter [7], offered commercially in the 1950's, used 1.3 cm to 3.8 cm pebbles enclosed in 30 cm thick panels. Several decades later, the Combustion Power Company developed a crossflow filter in which the gas flowed radially outward through an annular moving bed of 3 mm to 6 mm pea gravel [8]. Collection efficiencies for submicron particles were low, which lead to a system for electrically augmenting the performance of the filter. Plugging of the screens that enclosed the granular media was often a problem in this design. Combustion Power Company went on to develop a screenless moving bed filter to avoid plugging problems [6]. The resulting design appears to be the first countercurrent flow moving bed granular filter. A central gas pipe injects gas downward into the center of a hopper-shaped granular bed. The gas turns 180" to flow upward through the downward flowing solids. Granular material is fed to the surface of the bed through a complex of eight, gravity-fed pipes. Collection efficiency was 99% for
380
particles greater than 4 pm diameter and exceeded 93% for smaller particles. Some of the literature published on this filter suggests that most of the dust capture occurs in a zone very close to the injection point of gas into the bed. Delft University of Technology camed out a series of experiments in a counterflow moving bed filter similar to that for Combustion Power Company but employed a simpler feed system for the granular media [9]. These tests indicated acceptable collection efficiencies only at relatively low filter velocities. Based on observations by other researchers that formation of a dust cake is important to efficient dust collection for fixed bed granular filters, we have developed a new concept for a moving bed granular filter that makes use of this property. The goal is to establish a quasi-steady dust cake that is continuously or periodically renewed on the upstream side of the dust cake and swept away on the downstream side. In the proposed filter, granular material moving downward by gravity spills out of a centrally located dipleg to form an interfacial region where most of the gas cleaning occurs and the dust cake is formed. The lower edge of the filter cake is dispersed by the downward flow of granular material while the upper interface is covered by a fresh layer of granular material cascading from the dipleg above the interface. In this fashion, the interface establishes a dust cake of quasi-steady thickness, which is controlled to give high collection efficiency and acceptable pressure drop.
EXPERIMENTAL METHODS AND PROCEDURES Experiments were performed to establish appropriate hydrodynamic behavior of the moving bed filter. Tests were performed with a plexiglass prototype using soda-lime glass beads as the granular material and air as the gas stream. All tests were conducted at room temperature (25°C) and atmospheric pressure.
FILTER DESIGN The filter includes three innovations: a tangential gas inlet, a flow straightening section, and a screened gas disengagement section. As illustrated in Fig. 1, the gas enters the filter through a tangential inlet, which imparts a cyclonic motion to the gas flow. Inside the filter, the gas swirls downward towards the interface between the gas and granular bed. By imparting cyclonic flow, the momentum of the gas is preserved, reducing pressure drop normally associated with sudden expansion into a filter. However, bed granules and dust cake on the surface of the bed would be disturbed unless the radial component of the gas flow is redirected in the axial direction before the gas reaches the bed surface. A flow straightening section, consisting of evenly spaced fins projecting radially about the circumference of the annular space above the surface of the bed, accomplishes this redirection. The flow-straightening section also evenly distributes the gas flow over the surface of the bed, which is important to the efficient utilization of the filter media. Gas cleaning is hypothesized to occur primarily at this interface. The accumulation of dust particles on the granules and in the voids between granules forms a thin dust cake, which aids in the capture of dust particles in the gas flow. The granular bed flows downward by gravity while the gas flows upward through the bed. This counter-flow of gas and granules means that the dirty gas engages the
381
Gas Disengagement
Gas
Granular Material
-----
Fig. 1 . Schematic of moving bed granular filter
382
bed where the granules are the dirtiest while clean gas disengages the bed where the granules are the cleanest, which is important to high particulate collection efficiency. However, achieving counter-flow requires careful design of the filter’s gas engagement and disengagement sections. Some moving bed filters employ gas spargers or louvered panels to engage upward flowing gas with downward moving granules. However, since we hypothesize the formation of dust cake to be important for efficient dust collection, the entering gas flow must not disrupt the interfacial area between bed and gas. This is accomplished by letting granules flow out of a solids downcomer to spread out into a large, conical interface. Dirty gas flows downward through this interface, depositing dust, and then turns to flow upward through the downward flowing granular bed. The gas disengagement region also requires a special configuration to allow high gas flows through the filter. The upward flowing gas induces a drag on the granules that causes the bed to expand and eventually fluidize, an undesirable behavior that limits gas throughput through the filter. We have designed a gas disengagement section consisting of a small diameter feeder tube conveying granular material to a larger diameter downcomer. At low gas velocities, the granules from the feeder tube spread out into a conical pile much like the one in the engagement section. However, at high gas velocities, these particles expand upward against an annular porous plate or screen that prevents their continued expansion. The screen allows gas to exit the filter while retaining granular material. Gas does not enter the feeder tube by virtue of the large gas flow resistance in that direction. The filter was constructed from 6.35 mm thick polycarbonate plexiglass to give a clear view of bed fluid dynamics. The body of the filter is 19.8 cm dia., the downcomer is 14.2 cm dia., and the feeder tube is 10.2 cm dia. Not illustrated in Fig. 1 is a feed hopper above the filter and a discharge barrel at the bottom of the filter. The feed hopper provides fresh granular material to the top of the filter. The discharge barrel, which is sealed against the atmosphere, accepts dust-laden granular material from the bottom of the filter. Air at 20°C simulates the gas to be cleaned in the filter. The granular bed consists of washed 4-mm diameter soda-lime glass beads assumed to be perfectly spherical. The granular bed flow rate is controlled with a slide gate at the bottom of the filter. Bed depth, cross-sectional flow area, granular size and density, and gas travel distance were kept constant in these tests.
INSTRUMENTATION A rotameter, corrected for temperature and pressure, was used to measure volumetric flow rate of air up to 2.8 m3/min. Pressure drop was recorded at each superficial gas velocity using three Magnahelic differential pressure gauges (5, 12.7, and 25.4 mm HzO). Three differential pressure gauges were used to increase recording accuracy at different gas flow rates. Triplicate measurements were performed for superficial gas velocities between 0 - 3.1 d s in increments of 0.15 d s . At each gas flow rate, the granular bed material was allowed to flow for five minutes to reach steady state before the gas flow rate, differential pressure, and visual observations of the moving bed were recorded. In some experiments, titanium oxide smoke was injected into the air flow to visualize gas entering the filter and to obtain a qualitative indication of filtration efficiency. No quantitative measurements of collection efficiency were performed in the present experiments.
383
RESULTS AND DISCUSSION
The effects of three design features on performance of the moving bed granular filter were evaluated. These features include the screened disengagement section; the tangential inlet; and flow straightening fins above the gas engagement section.
EFFECT OF SCREENED DISENGAGEMENT SECTION We made qualitative observations on the behavior of granular material moving through the downcomer as a function of superficial gas velocity, which is defined as the volumetric flow rate divided by the cross sectional area of the annular region formed between the filter body and the downcomer. The filter was tested both with and without the retaining screen installed above the disengagement section. Observations on the behavior of the granular bed in the absence and presence of the screen are recorded in Tables 1 and 2. Without the screen, granules at the surface of the bed in the disengagement section become agitated with increasing superficial velocity and eventually elutriate from the filter. We estimate the maximum operating velocity to be 0.9 d s without a retaining screen. When a screen is installed above the disengagement section, the granules expand against the screen but are not able to
Table 1 Performance of downcomer without retaining screen in disengagement section Superficial gas Observations of downcomer flow velocity ( d s ) 0 - 0.3 Smooth downward granular bed flow with no fluidization. 0.3 - 0.6
Smooth downward granular bed flow. Bed material begins to expand upward.
0.6 - 0.9
Smooth downward granular bed flow. Full expansion of the bed with the upper region becoming fully fluidized.
BO.9
Material in top half of downcomer fully fluidized; granular material entrained and elutriated through the gas exit.
Table 2 Performance of downcomer with retaining screen in disengagement section Superficial gas Observations of downcomer flow velocity ( d s ) 0 - 0.5 Smooth downward granular flow with no fluidization. 0.5 - 0.8
Smooth downward granular flow. Bed beginning to expand against screen with slight fluidization on top.
0.8 - 1.5
Smooth downward granular flow. Full expansion against the screen with no fluidization present.
~1.5
Top half of granular bed in downcomer is stationary and fully expanded against screen while bottom half becomes fluidized.
3 84
fluidize. We estimate the maximum operating velocity to be 1.5 m / s . The screen increased gas throughput by 67% compared to the filter operated without the screen.
EFFECT O F TANGENTIAL INLET Both a conventional gas inlet entering perpendicularly to the body of the filter and a tangential inlet that imparts a cyclonic swirl to gas entering the filter body were investigated. Both configurations employed retaining screens in the disengagement sections and 4 mm dia. glass beads in the granular bed. In tests with the tangential inlet, flow-straightening fins were employed just above the gas engagement section to prevent scouring of the interfacial region of the bed, as subsequently described. Figure 2 plots pressure drop as a function of superficial velocity in the filter for the two inlet configurations. Pressure loss was significantly less for the tangential inlet than for the perpendicular inlet. At the maximum operable superficial gas velocity of 1.5 m/s the pressure loss for the tangential inlet was 44% less than for the perpendicular inlet. The momentum of the gas is better conserved using the tangential gas inlet. When the gas enters through the tangential gas inlet, the momentum of the horizontally flowing gas is efficiently converted to cyclonic flow in the body of the filter. On the other hand, when the gas is injected perpendicularly into the filter, the gas loses momentum as the flow is rearranged. 700 600
500 E E
2
400
g $
300
u)
Lf
200 100
0 0
0.2
0.4
0.6
0.8
1
1.2
1.4
Superficial1 velocity ( m k )
Figure 2: Pressure drop comparison of perpendicular and cyclonic inlets
385
1.6
EFFECT OF FLOW STRAIGHTENING FINS Although the tangential inlet greatly reduces pressure drop in the filter, the cyclonic flow scours the surface of the granular bed, disrupting the formation of dust cake deposited by entering gas. Fins were installed in the annular space formed between the downcomer and filter shell in an effort to straighten gas streamlines flowing downward toward the engagement section of the granular bed. Qualitative observations were made of the bed surface as a function of superficial gas velocity for perpendicular and tangential gas inlets and recorded in Tables 3 and 4, respectively. These results illustrate that fins reduced scouring of the interfacial area at gas velocities above 0.9 m / s . Figure 3 is a rendition of gas stream lines for the tangential Table 3 Behavior of granules at bed interface with no fins in gas engagement region Gas-solid contacting region observations Superficial gas Velocity ( d s ) Gas engagement interface is smooth with no scouring. 0 - 0.9 0.9 - 1.5
Interface becomes scoured on side opposite of gas inlet.
Table 4 Behavior of granules at bed interface with fins in gas engagement region Superficial gas Gas-solid contacting region observations Velocity ( d s ) 0 - 1.5 Gas engagement interface is smooth with no scouring. Gas is evenly distributed with little scouring.
>1.5
U
U
Fig 3. The role of fins in preventing bed surface scouring when using tangential inlet (a) flow lines in absence of fins (b) flow lines in presence of tins
386
inlet with and without fins based on visual observations with titanium oxide smoke and the behavior of granules at the bed surface. The fins were effective in converting the cyclonic flow into linear flow that entered the bed at high incidence angles.
CONCLUSIONS The moving bed granular filter presented in this paper is capable of high gas throughput with low-pressure drop while evenly distributing the gas over the gas-solid contacting region. Three design features achieve these characteristics: a tangential gas inlet, a flow straightening section upstream of the gas engagement region, and the use of a retaining screen in the gas disengagement section. Based on the results of these cold flow tests, we have designed and constructed a hot-flow moving bed granular filter to be tested with a 4.5 tonne per day biomass gasifier in the near future.
REFERENCES 1. Mustonen, J. P. et al. (1991) Technical and economical analysis of advanced particle filters for PFBC applications. In: Proceedings International Conference on Fluidized Bed Combustion,Vol. 1, pp. 475-480, Montreal, Canada. 2. Wilson, K. B. & Haas, J. C. (1994) Granular -bed and ceramic candle filter comparison for advanced power systems. In: Proceedings Eleventh International Pittsburgh Coal Conference, Vol, 1, pp. 131- 136, Pittsburgh, PA. 3. Staubly, R. K. et al. (1994) Overview of METC’s gas stream cleanup program. In: Proceedings Eleventh International Pittsburgh Coal Conference,Pittsburgh, PA. 4. Squires, A.M. & Pfeffer, R. (1970) Panel bed filters for simultaneous removal of fly ash and sulfur dioxide: 1. Introduction. J. Air Pollut. Control Assoc. 20, 534-538. 5. Lipert, T. E. et al. (1981) Testing and verification of granular bed filters for removal of particulates and alkalies. High Temperature, High Pressure Particulate and Alkali Control in Coal Combustion Process Streams. In: Proceedings U.S. DOE Contractors’ Meeting, COW-810249, pp. 471-489. 6. Hall, H. J. & J. C. Munday (1946) Purification ofgases. U.S. Patent 2,411,208. 7. Dorfan M. I. (1952) Method and apparatus for suppressing steam and dust rising from coke being quenched. U.S. Patent 2,604,187. 8. Saxena, S. C. et al. (1985) Particulate removal from high temperature, high pressure combustion gases, Prog. Energy Combust. Sci. 11, 193-251. 9. Zevenhoven, C. A. P., et al. (1992) Thefiltration of PFBC combustion gas in a granular bedfilter. Filtration and Separation 29,239-244.
387
Redox Process for the Production of Clean Hydrogen from Biomass S. Biollaz, M. Sturzenegger, S. Stucki General Energy Research, Paul Scherrer Institut, Villigen, Switzerland
ABSTRACT: The concept of using a metal oxide based thermochemical cycle for producing hydrogen from biomass is discussed. The cycle is based on the partial reduction of iron oxide using producer gas fiom a wood gasifier and subsequentreoxidation of the original oxide by water vapour. The eficiency of the process depends strongly on the thermochemical solid state properties of the metal oxide phases as well as the thermal and dynamic integration of the process with the air driven biomass gasifier. Design parameters for the development of materials and processes are discussed and research needs identified.
INTRODUCTION The central problem in efficiently using wood for the decentralised production of heat and power has been the gas clean-up process. Tars and particles need to be removed efficiently to prevent depositions or excessive wear in the power conversion devices (gas engines, gas turbines).The problems and costs associated with efficient gas cleanup, especially for the conversion or removal of tars, have not been solved so far in a satisfactory way and have hindered the wide-spread introduction of gasification as a means of utilising fuel wood with high exergy efficiency. Fuel cells (FC) have the advantage of providing clean, near zero emission conversion of reformed fuels to electricity, with high conversion efficiency in relatively small units. FC demonstration plants operated with natural gas are typically in the power range up to 250 kW electricity (1, 2). The high efficiency in small units makes fuel cells ideally suited for decentralised utilisation of fuel wood in the future, provided the fuel can be processed to be compatible with the input requirements of the fuel cell system. The specific requirements of fuel cell power converters with respect to levels of tars and particles in the feed gas have not yet been investigated in detail. There are good chances that high temperature fuel cell systems such as the solid oxide fuel cell (SOFC) operating at similar temperature as the gasification process might be able to cope with tars if condensation is avoided by integrating thermally the fuel cell and the gasifier. The coupling of the low temperature fuel cell systems, in particular the proton exchange membrane fuel cell (PEMFC), with wood as primary fuel needs high quality fuel processing. The PEMFC requires hydrogen fuel with levels of CO not exceeding 20 to
388
50 ppm. For such an application new ways of gas processing of the raw gas from a wood gasifier are needed. The present paper discusses the option of using a metal oxide redox cycle for gas conversion to clean hydrogen, with separate gas streams avoiding the contamination of the converted gas. A process using cyclic oxidation and reduction of sponge iron has been examined by Friednch et al. for the production of hydrogen for fuel cells (3-6). The sponge-iron process has been shown to reach an overall efficiency of 50% for the production of hydrogen from producer gas. With a cold gas efficiency of the gasifier of 70% this translates into an ef'ficiency of 35% for hydrogen from wood. With an efficiency of the fuel cell of 50% (operating on hydrogen) the overall efficiency of the system comes down to less than 20% which clearly does not justify the investment into an expensive fie1 processor and power converter system. A metal oxide process of the type proposed by Friedrich will only be feasible if the overall efficiency for hydrogen production from wood can be increased to values above 50%. Besides reachmg minimum limits for efficiency the system as a whole must be economically competitive with the benchmark technology which is fixed bed wood gasification with appropriate wet gas clean-up for operating a gas engine. There are two aspects whch need to be considered in the development of a redox system for hydrogen production from wood: 1) materials: in order to overcome some inherent limitations of the process based on iron oxides, the oxide materials need to be modified by alloymg with other metal oxides such that the thermodynamics of the chemical conversion is more favourable for reaching high efficiencies. At the same time the materials need to be stable for a sufficient number of redox cycles. 2) system integration: The gasifier and the redox process need to be coupled thermally in such a way that losses and condensation can be avoided. At the same time the integrated gasifier with redox filter needs to be simple in design in order to meet the economic goal of low investment costs. We see cost saving potential in the design of the gasifier itself as it is expected that a simple grate-type gasifier will be sufficient to produce a gas suitable for processing in the filter.
The present paper reviews the potential of using the redox properties of modified iron oxide for efficient production of hydrogen in decentralised units. Fig. 1 shows, in a logarithmic scale, the range of present and future hydrogen applications as well as the power range of available gasifier technologies. The range of 100 to 1000 kW (shaded area) seems to be a reasonable target, as it suits best the criteria of decentralised fuel availability, hydrogen market and heat utilisation. Heat utilisation is a prerequisite for an economic production of electricity or a chemical commodity from wood. Heat utilisation becomes expensive, however, if large heat distribution systems are involved. Therefore, we see the most promising applications in the range indicated.
389
Energy H,
Chemical H,
Fig. 1 Present and future hydrogen applications with target for decentralised fuel availability, hydrogen market and heat utilisation.
THE REDOX PROCESS USING IRON BASES OXIDE MATERIALS The redox-process is based on a thermochemical cycle making use of the following set of chemical reactions (for simplicity the reaction is formulated with pure iron oxides as the solid reactants): CO + H2 + C,H, + Fe304 + C02 + H20 + FeO FeO + H20 + Fe304 + H2 Both these reactions take place at temperatures between 800 OC and 900 "C , The iron oxide acts as a solid state oxygen source for the oxidation of the combustible part of the producer gas (reaction 1) and as an oxygen receptor in the subsequent water splitting reaction. It is of central importance that the bulk of the oxide material can be cycled between the two oxidation states without breaking apart. The close relationship between the crystallographic structure of Fe304and FeO is an important prerequisite to fulfil this requirement. In t h i s context it is important to keep the reaction conditions such that reduction to metallic iron does not occur. Figure 2 illustrates the thermodynamic limitations for the reduction of iron oxide by means of producer gas at 1100 K. It shows that the reduction of Fe304to FeO with CO only proceeds as long as the CO2/CO ratio is smaller than 18. Correspondingly, reduction only occurs until the H20/H2 ratio in the gas reaches a value of 19. This means that about 95% of the CO and H2, respectively, can be converted to COz and H20, respectively. An analysis of the thermodynamic data (7) revealed that the amount of remaining combustible components strongly depends on the specific character of the chosen oxide. Based on the thermodynamic data for a more accurate stoichiometry for the reduced iron oxide, i.e. Feo.9470,a CO2/COratio of only 2.2 will be established above the oxide, a ratio far less favourable for an efficient conversion of the producer gas. T h s deviation 390
from ideal stoichiometry for FeO with its effect on the thermochemical equilibria might have been the reason for the limited overall process efficiency achieved in previous attempts to use the iron oxide process for hydrogen production (5). On the other hand, alloying iron oxide with other transition metal oxides ylelds mixed iron oxides with new and in some cases more preferable thermochemical properties. From Figure 2 we can estimate that, e.g., the manganese-iron mixed oxide MnFe2O4can be reduced to the ideal solid solution (&.33Feo.67)0 by H20-Hz mixtures with ratios as hgh as 184. Thus, thermodynamics suggests an almost complete conversion of the producer gas when optimised oxides are employed. One should not neglect, however, the flip side of the coin: A h g h H20/H2ratio in equilibrium with the mixed oxide redox system means that for the hydrogen producing back-reaction 2 large amounts of water vapour will be
0 0.01
1
0.1
10
100
500
P(CO2) P(W
Fig.2 Existence of pure and mixed iron oxides as a function of the partial pressure ratios H20/H2and CO2/CO,respectively. The equilibrium composition of the producer gas after conversion is indicated with P1 (pure iron oxide) and P2 (manganese iron oxide). needed to oxidise the reduced oxide and to produce the same amount of HZ. The challenge in the development of materials for the redox process will be the fine-tuning of the thermochemical properties of mixed oxides for optimised performance in the reduction as well as the oxidation mode of the cycle. This optimisation will have to minimise overall thermal losses by finding tradeoffs between heat loss in the provision of water vapour and fuel losses in incomplete conversion of the raw gas. To illustrate the feasibility of the proposed closed loop process, pure iron oxide Fe304was subjected to a producer gas from a 100 kW fluidised bed gasifier (reaction 1) and subsequently exposed to water vapour (reaction 2). For the reduction, Fe304 powder was placed in alumina crucibles and heated in a tubular furnace at 900 "C and 1000 OC, respectively while being exposed to a flow of producer gas for at least 25 minutes. The characterisation of the product by powder X-ray diffractometry and thermal analysis revealed that Fel-,O indeed is the main product. Metallic iron and carbonaceous materials were identified as by- products. Generally, samples prepared at 391
900 "C were contaminated with higher amounts of carbonaceous material than those prepared at 1000 "C. The amount of iron showed the opposite tendency. We assume that at higher temperatures deposits of carbonaceous materials preferably reacted with the oxidic bed reducing it locally to Fe. Although such a reaction increases the conversion of combustible components, it has to be avoided. We expect that complete reduction of the oxidic bed to the metal will impair the material's stability with respect to cycling. An important aspect of the development of a material for the redox process is its catalytic activity for converting tars with low coking tendency.
co, co, 10
I 20
30
Time [min]
Fig. 3 Evolution of the gas phase during the reaction of the reduced iron oxide in a N2H20 gas stream (p(H20)=123 mbar) at 1073 K. The circles (0)represent CH4, diamonds (+) represent H2.
The capability of the reduced oxide to produce H2 from H20 (reaction 2) was proven by exposing the oxide at 1073 K to a N2-H20gas stream with a water vapour pressure of 123 mbar. Figure 3 shows the evolution of the gaseous species during the exposure. H2 was identified as the main product. It is formed as soon as the water vapour is added to the gas stream (indicated by arrow). H2 production is typically completed after 7 to 12 minutes, a time range that was also observed for synthetic Fe,.,O (8). At first glance surprising is the evolution of methane (CH4)during the early stage of the measurements. It must be pointed out here that the intensity of the methane signal is not to scale and has not been evaluated quantitatively. We ascribe the occurrence of CH4 to the deposition of carbonaceous materials during the reduction. The observed formation of methane shows that the complete conversion of carbonaceous materials in the reduction step of the redox material (reaction 1) is not only important with respect to cycling stability, but as well with respect to product quality, i.e. purity of the produced hydrogenand with that its possible applications as a fuel for fuel cells or other. If it can be confirmed that the main impurities are in fact methane only and no detectable CO, as these preliminary experiments show, then the gas can be used as a fuel for PEM fuel cells.
3 92
SYSTEM INTEGRATION Economically, the redox process only makes sense if the expenses for the redox converter can be balanced by savings in the gasifier and primary gas clean-up system, i.e. if the system as a whole does not cost more than a conventional fixed bed downdraft gasifier with wet or dry gas clean-up (benchmark). It is expected that the redox
- ...
Fig. 4 Flow-chart of Integrated Gasifier with Redox Filter in batch-wise operation.
Fig.5 Flow-chart of Integrated Gasifier with Redox Filter in continuos operation. filter material will be reactive to convert tars to C02 and H 2 0 and that for this reason the primary gas quality produced by the gasification system is of minor importance. This means that gasification may be canied out using a simple design of the gasifier, such as a sub-stoichometrically operated grate hrnace. State-of-the-art fixed bed gasifiers are thermally optimised such that the outlet temperature is in the range of 400 "C, i.e. considerably lower than the adiabatic temperature of the gasification process itself. Conventional fixed bed gasificers must 393
therefore be modified to reach the needed temperature of 800 to 900 "C for operation with the redox gas processing unit. In order to achieve high thermal efficiency and to avoid the condensation of tars, the redox system needs to be integrated thermally into the gasification system. Practically this could be done by i.e. containing the redox filter and the gasifier in one furnace. Two different concepts of operation of the redox filter are conceivable: One is the batch-wise operated system as it has been discussed in (6) (see Figure 4). Two fixed bed reactors containing pelletised oxide are operated alternatively in the oxidation and in the reduction mode. The flow of the hot media is switched in intervals the duration of whch depends on the redox capacity of the oxides, the kinetics and the size of the fixed bed reactors. The other concept is a continuous system consisting of a twin fluidised bed with two reaction zones (see Figure 4) between which the fluidised metal oxide pellets are circulating (see Figure 5). Both systems have their advantages and drawbacks with respect to design, operation and requirements to the redox material. Design details are currently specified.
CONCLUSION The concept of using a thermochemical cycle involving the reduction and oxidation of metal oxides is a promising alternative to produce high quality fuel fiom biomass via air gasification. The feasibility of the process will depend strongly on the overall efficiency which can be reached. Of central importance for reaching attractive efficiencies is the development of a suitable oxide material. A thermochemical analysis of candidate metal oxides which, from their structural properties are suitable for cycling, reveals that without stabilisation of the FeO stoichiometry the achievable conversion of producer gas is not high enough. Stabilisation can be achieved by alloying the oxides with other transition metals. The redox process takes place at relatively high temperatures which implies thermal integration of process steps in order to minimise losses. The advantages and disadvantages of fixed bed or fluidised bed operation of the oxide will have to be evaluated.
REFERENCES 1. Pokojski, M. (1999) The first demonstration of the 250-kW polymer electrolyte
fuel cell for stationary application (Berlin). Journal of Power Sources, 86, 140-144. 2. George R.A. (1999) Status of tubular SOFC field unit demonstration. Journal of Power Sources, $6, 134-139. 3. M. Selan, J. Lehrhofer, K. Friedrich, K. Kordesch, G. Simader (1996) Sponge Iron: Economic, Ecological, Technical and Process-Specific Aspects. Journal of Power Sources, 61, pp. 247-53 4. V. Hacker, G. Faleschini, H. Fuchs, R. Fankhauser, G. Simader, M. Ghaemi, B. Spreitz, K. Friedrich (1 998) Usage of Biomass Gas for Fuel Cell by the SIR Process. JournaI ofpower Sources, 71, pp. 226-30
394
5. K. Friednch, V. Hacker, H. Fuchs, R. Fankhauser (1998) Solid Biomass Gasification for Fuel Cells Gas Purification - Sponge Iron Reactor. Summary. Techmsche Universitat, GrdAustria 6 . Hacker V., Fankhauser R., Faleschini G., Fuchs H., Friedrich K., Muhr M., Kordesch K. (1999) Hydrogen production by steam-iron process. Journal of Power Sources, 86,531-535. 7. Barin I. (1995) Thermochemical Data of Pure Substances, 3rd edn.. VCH Weinheim. 8. Ehrensberger, K. (1 995) Zweistufiger Metalloxidzyklus zur chemischen Speichenrng von Sonnenenergie. PhD thesis, University of Zurich.
395
Hydrogen production from biomass by low temperature catalytic gasification Tomoaki Minowa and Zhen Fang National Institute for Resources and Environment, Onogawa 16-3, Tsukuba, Ibaraki 305-8569 Japan
ABSTRACT: Cellulose, a major component of woody biomass, was reacted in the hot-compressed water of 200 - 350 "C and 4 - 22 MPa using a reduced nickel catalyst. The product distribution was analyzed to elucidate the overall reaction mechanism on the low temperature catalytic gasification. Cellulose decomposed to water-soluble products and gases with increasing in the reaction time, and then, the water-soluble products decreased to form gases after no cellulose remained in the reactor. The water-soluble products are considered as intermediates. The obtained gases mainly consisted of COz, H2 and CH4. At low gas yield, C 0 2 and H2 were obtained in excess of equilibrium, and then, H2was consumed to form CH4 with increasing in the gas yield. The reaction rate of methanation is slower than that of steam gasification, and the low temperature catalytic gasification can be expected as a method of hydrogen production from wet biomass.
INTRODUCTION At present, biomass is world wide recognized as a rich source of renewable energy. For the traditional conversion processes like combustion, pyrolysis and gasification, the feedstock needs to be dried to a relatively h g h degree. However a large part of the available streams of biomass and organic waste have a h g h moisture content, whch makes them difficult to treat. It is therefore important to develop a new conversion technology suitable for wet biomass and wastes. Elliott et al. reported wet biomass to be directly gasified into methane rich fuel gas by low temperature catalytic gasification using a metal catalyst in hot-compressed water of 350-400 OC and 20-30 MPa in around 1990.''3 They have developed a continuous-flow reactor process, named the Thermochemical Environmental Energy System (TEES). On the other hand, we found that hydrogen rich gas could be obtained during the low temperature catalytic ga~ification.~ During the gasification, three major reactions, steam reforming (Eq. (l)), water-gas shift (Eq. (2)) and methanation (Eq.' (3)), can occur, and our finding suggests that methanation reaction proceeds slowly.
396
+
CHxOy+ (1-y) H2O CO + (x/2+1-y) H2 CO + H20 COZ + Hz CO + 3 Hz CH4 + HzO C02+4H2+CH4+2H20
+
or
(1)
(2) (3)
In this paper, cellulose was reacted in hot-compressed water at different reaction temperatures and time using a reduced nickel catalyst. The aim of the study is to elucidate the overall reaction mechanism on the low temperature catalytic gasification and to get an insight into the hydrogen production.
EXPERIMENTAL, MATERIALS A micro-crystalline cellulose (E. Merk) was used as the starting material. A commercial catalyst (NI-3288, Engel-Hard) was used as the nickel catalyst. It was crushed to 60-200 mesh, and reduced using hydrogen gas at 350 "C for 4 hour before reactions.
REACTION The reaction was performed in a conventional autoclave (120 cm3 capacity and 18 cm3 of head-space) with a magnetic stirrer. Five g of cellulose, 2 g of the nickel catalyst and 30 ml of water were charged in the autoclave. Nitrogen gas was used to purge the residual air in the autoclave, and it was added to a pressure of 3 MPa to avoid vaporization of water during the reaction. Then, the reaction was started to heat the autoclave using an electric furnace. When the temperature in the autoclave was raised to the desired temperature of 200-350 "C, the autoclave was immediately cooled down to room temperature using an electric fan. At 350 "C of the reaction temperature, the temperature in the autoclave was held during 0-1 h before the cooling down. Since the heating rate, as a function of time, was well reproduced in the experiments, the reaction temperature can be considered as function of the reaction time.
SEPARATION After cooling down to room temperature, the product gas was removed to a sampling bag. Its volume was measured using a gas meter (Shmagawa-sheilu, W-NK-SB), and its composition was determined with gas chromatography (Shimadzu, GC-12A with TCD detector or GC-9A with FID detector). Then, the autoclave was opened, and reaction mixture was removed for separation to aqueous phase, oil and residue, respectively. The aqueous phase was separated by washing the reaction mixture with water and by filtration. Then, acetone was used to separate tarry product from water-insoluble fraction by washing and filtration. The oil was obtained from the acetone solution by 397
evaporating acetone at 70 "C. The acetone-insoluble fraction remaining on the filter paper was b e d at 70 "C to obtain the residue which includes recovered catalyst. ANALYSIS The elemental composition of the oil and residue was determined by an elemental analyzer (AMCO,NA-1500). The carbon amount in the aqueous phase was measured by a TOC meter (Yanaco, TOC-8L). Cellulose remained after the reaction was measured as glucose unit preserved in the residue. That is, the residue was hydrolyzed to glucose, and the glucose amount was measured by the phenol-sulfuric acid method. The amount of cellulose was calculated by the following equation. cellulose (g) = residue (8) x glucose (g/g) obtained from the hydrolysis of residue glucose (g/g) obtained from the hydrolysis of cellulose
(4)
The solid reaction products, which have no glucose-unit, in the residue (hereafter referred to as the char) were calculated by difference: char (g) = residue (g) - remained cellulose (g) - catalyst (g)
(5)
The hydrolysis procedure was referred to in the literature.'-' The residue or cellulose, about 0.1 g, was hydrolyzed with 7 ml of 71 vol% sulfuric acid at 30 "C for 1 h, and then, it was diluted to 125 ml (to 4 vol% sulfuric acid) and boiled for 1 h for the secondary hydrolysis. The obtained solution of 50-100 pl was diluted to 2 ml, and 50 pI of 80 wt% phenol and 5 ml of 100 vol% sulfuric acid were added for coloring. The solution was cooled down to room temperature using a water bath for more than 20 min, and then, an adsorption at about 487 nm was measured by an adsorption meter (Shimadzu, W-16OA). The total sugar amount in the aqueous phase was measured in a similar manner.
RESULTS AND DISCUSSION PRODUCTS DISTRIBUTION The amount of water-soluble products could not be measured correctly, because a part of the water-soluble products can be vaporized during the drymg process of the aqueous phase, which is needed to get the products. In addition, the amount of char could not be measured correctly at same cases, since the residue included much amount of the recovered catalyst. On the other hand, carbon amount in each fraction could be measured, therefore, the product distribution is discussed on a carbon basis. Figure 1 shows the product distribution. Below 260 "C, almost cellulose remained after the experiments, and it decomposed quickly between 260 and 300 "C. Over 300 "C, no cellulose remained in the autoclave after the reaction. Between 260 and 300 "C,gases and water-soluble products were obtained; little oil and little char were obtained. The yield of the watersoluble products had the highest value at 300 "C, and it decreased with temperature 398
above 3 10 "C. On the other hand, the increase in gas yield continued at over 310 "C, although no cellulose remained in the autoclave at such temperatures. Thus, the water-soluble products were intermediates for gas formation. +Total
+Oil
+Cellulose +Aqueous
+Char +Sugar
120
.............
80
8 Is
E
..............
60
. s F
f
3.
100I
j
+Gas
..............
40
20
0 200
250 300 350 Reaction temperature I degreeC
0.5h I h hold hold
Fig. 1 Product distribution at dfferent reaction temperatures cellulose water-soluble products 3 gases (6)
+
In our previous work, hydrolysis is played an mportant role at the first step of cellulose decomposition in hot-compressed water without catalyst.* In this study, the sugar, which is the hydrolyzed products of cellulose, was detected at low reaction temperature of 260 and 280 "C, and its concentration was about 40% in the watersoluble products. This means that hydrolysis also plays an important role for gasification with the nickel catalyst, and the obtained sugar can decomposed quickly to non-sugar products. GAS PRODUCTS The obtained gases mainly consisted of carbon dioxide, hydrogen and methane. Hydrocarbons of Cz+were detected, while small carbon monoxide was determined. This suggests the water-gas shift reaction (Eq. (2)) proceeds quickly. Figure 2 shows the mole fractions of each gas in obtained gas at different gas yield. The higher gas yield shows the higher proceeding of the gasification reaction. The mole fractions in equilibrium are also shown on right. At low gas yield, carbon dioxide was obtained in the high mole fraction, and it approached quickly to the equilibrium mole fraction of 0.49 with proceeding the reaction. Hydrogen was produced in higher mole fraction beyond the equilibrium mole fraction of 0.05 in the wide range of the gas yield. Ths suggests that the methanation reaction (Eq. (3)) proceeds slowly. At higher gas yield 399
than 0.6, the mole fraction of methane tended to increase, while that of hydrogen tended to decrease slightly. At the gas yield as close as 1, the equilibrium mole fraction is expected to obtained.
1 4 H 2 +C02 '
-0-CH4
-Et-CO
I
+C2+ I
t
Equilibrium 0.49, C02 0.46, CH4
0.05, H2
0
0.2
0.4
0.8
0.6
1
Gas yield I mollmol on carbon basis
Fig. 2 Mole fraction in obtained gas vs. gas yield To c o n f i i the methanation reaction and the reverse reaction, separate experiments with hydrogen, carbon dioxide and methane as the reactants were performed. Table 1 shows the results. Run 1 was a blank experiment, and small amount of H2, C02 and CH4 were detected due to contamination. Run 2 was performed with H2 and C 0 2 as reactants, and CH4 was produced after the reaction. It means that methanation reaction through H2 and C 0 2 can occur. On the other hand, Run 3was performed with CH4 as the reactant, and neither H2 nor C02was obtained (blank level). This means that the reverse reaction of methanation (steam reforming of CH,) cannot occur under these conditions. From these results, the steam reforming reaction (Eq. (1)) and the water-gas shift reaction (Eq. (2)) occur at fust to produce C 0 2 and H2, and then the methanation reaction proceeds to form CH4 during the low temperature catalytic gasification: cellulose / decomposed products
+ COz + + CH, H2
400
(7)
Table I Experiments with gases as reactants
Feedstock
Product gases/ mmol
Run 1 N2, 3 MPa
Run 2 NZ,3 MPa Hz, 2 MPa (about 86 mmol) COz, 1 MPa (about 45 mmol)
Run 3 Nz, 3 MPa CH4, 2 MPa (about 88 mmol)
2.9 0.3 0.2 0.0
53.4 40.6 6.1 0.0
2.7 0.5 87.3 0.0
I
co TOTAL REACTION MODEL
In our previous studies, water-soluble products are also considered as intermediates for the oillchar production under catalyst free conditions:* cellulose
+ water-soluble products + oil + gas (COz) + char (sugar + non-sugar aqueous products)
(8)
In this study, a significant amount of catalyst (40 wt% to cellulose) was used, and little oil and little char were obtained. On the other hand, oil and char were obtained . ~ reaction in the considerable yields at the lower catalyst loading of 5-10 ~ t % The pathway of polymerization, oil and char production, may compete the gasification pathway. To draw the reaction model, separate experiments were performed with the decomposition products of cellulose. Cellulose was reacted in 300 "C of hotcompressed water with 5 wt% of Na2C03(Run 4) to obtain the water-soluble products and the oily products, whch were easily separated by decantation. Then both were gasified with the nickel catalyst. As shown in Table 2, the water-soluble products could be gasified to Hz, COz and CH4,while the oily, polymerized products could not be gasified. T h s result supports the competition of gasification pathway to polymerization pathway. The following total reaction model can be drawn.
/
Gases (C02 + H2) -b
C&
gasification hydrolysis Cellulose -bWater
soluble products b Non-sugar aqueous products Sugars degradation
\+
polymerisation
Oil
401
Char
Table 2 Reactivity of the decomposition products Run 4 Run 5 Run 6 Feedstock Cellulose Oily products Aqueous products Yield I % on carbon basis 7.0 47.6 Gas 15.0 Oil 53sa 32.8 2.1 17.3 1.5 Char Aqueous phase 3 1Sb 16.6 24.5 74.1 75.7 Total a: oily product (feedstock for Run 5) b: aqueous products (feedstock for Run 6) conditions; Run 4, water 30 ml, Na2C030.25 g, 300 "C, 0.5 h holding Run 5, water 30 ml, Ni cat. 1 g, 350 "C, 0 h holding Run 6, Ni cat. 1 g, 350 "C, 0 h holding.
1. acetaldehyde 2. propylene aldehyde 3. acetone 4. butanedione 5. acetic acid 6 . furfural 7. propionic acid 8. iso-propanol 9. ethanol 10. 2-butanol 11. cyclohexanol
1 Before gasification
L
0
d I
r I
.
i I
d I
d I
i
-
d
d
-
I
-
I
-
I
d
-
I
d
i
d
N I
I
0
&
N
0
<
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I
I
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A 0
d
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1
1
1
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-
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1
i
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"
1
0
o
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1
i 1
0
d l
0
l
0
d 1
i 1
Fig. 3 Analysis of aqueous phase before and after the gasification
402
In this model, the nickel catalyst catalyzes the degradation pathway and the methanation reaction. The water-soluble products are considered as important intermediates. Therefore, aqueous phase was analyzed with GC-MS and GC-FID (a DB-wax column). Figure 3 shows the chromatogram of before and after gasification of aqueous phase (Run 6). Before gasification, aldehydes, ketones, organic acids and fiufural were detected, while these compounds disappeared, and alcohols were obtained after gasification. Aldehydes, ketones and organic acids are considered as direct reacting species for the gasification, and more detail reaction scheme is planned to be clarified in a future work.
CONCLUSION
Cellulose, a major component of woody biomass, was gasified in hot compressed water different reaction temperatures of 200-350 "C using a reduced nickel catalyst. The reaction mechanism is discussed based on the product distribution. The watersoluble products are considered as intermediates, and the obtained hydrogen is consumed by the methanation reaction. The analysis of the water-soluble products shows that aldehydes, ketones and organic acids are considered as direct reacting species for the gasification. The simplified reaction model of Eq. (9) is proposed. The reaction rate of methanation is slower than that of steam gasification, and the low temperature catalytic gasification can be expected as a method of hydrogen production from wet biomass.
ACKNOWLEDGMENT This study was funded by the New Sunshine Project of Agency of Industrial Science and Technology. The Agency of Industrial Science and Technology supported the fellowship of Dr. Zhen Fang. The authors thank deeply Dr. Shin-ya Yokoyama of National Institute for Resources and Environment and Dr Tsutomu Suzuki of Kitami Industrial University for their valuable comments. We thank also Ms. Yukiko Fukuda for her assistance with the experimental work. REFERENCES Elliott, D. C., Butner, R. S. and Sealock, L. J., Jr. (1988) Low temperature gasification of high-moisture biomass. Res. Thermochem. Biomass Convers., 696-710. Sealock, L. J., Jr. and Elliott, D. C. (1991) Method for the catalytic conversion of lignocellulosic materials. U. S. Patent 5019135; Chem. Abstr., 115,947248. Elliott, D. C, et al. (1993, 1994) Chemical processing in high-pressure aqueous environments 1-4. Ind. Eng. Chem. Res., 32, 1535-1541, 1542-1548, 33, 558565, and 566-574. Minowa, T., Ogi, T. and Yokoyama, S. (1995) Hydrogen production from wet cellulose by low temperature gasification using a reduced nickel catalyst. Chem. Lett., 937-938. 403
5.
6.
Saeman, J. F., Bubl, L. and Harris, E. E. (1945) Quantitative saccharification of wood and cellulose. Ind. Eng. Chem.,Analytical Edition, 17,35-37. Mok, W. S., Antal, M. J., Jr. and Vfirhegyi, G. (1992) Production and parasitic pathways in dilute acid-catalyzed hydrolysis of cellulose. Ind. Eng. Chem. Rex, 31,94-100.
7. 8.
Japan Industrial Standard (1976) Testing method for lignin in wood for pulp. JIS P8008. Minowa, T., Zhen, F., Ogi, T. and Varhegyi, G. (1998) Decomposition of cellulose and glucose in hot-compressed water under catalyst-fiee condition. J. Chem. Eng. Jpn., 31, 131-134.
404
Production of Substitute Natural Gas by Biomass Hydrogasification M. Mozaffarian, R.W.R. Zwart Netherlands Energy Research Foundation, ECN, Westerduinweg 3, P.O.Box I , I755 ZG Petten, The Netherlands
ABSTRACT: Hydrogen, generated from renewable sources, is llkely to play a major role in the future energy supply. The storage and transport of hydrogen can take place in its free form (Hz), or chemically bound, e.g. as methane. However, the storage and transport of hydrogen in its fiee form is more complex, and probably would require more energy than the storage and transport of hydrogen in chemical form. An additional important advantage of the indirect use of hydrogen as energy carrier is, that in the future renewable energy supply, parts of the existing large-scale energy *astructure could still be used. Production of Substitute Natural Gas (SNG) by biomass hydrogasification has been assessed as a process for chemical storage of hydrogen. Thermodynamic analysis has shown the feasibility of this process. The product gas of the process has a Wobbe-index, a mole percentage methane, and a calorific value quite comparable to the quality of the Dutch natural gas. With a hydrogen content below 10 mol%, the produced SNG can be transported through the existing gas net without any additional adjustment. The integrated system has an energetic efficiency of 81% (LHV). In the long term, the required hydrogen for this process can be produced by water electrolysis, with electricity from renewable sources. In the short term, hydrogen may be obtained from hydrogen-rich gases available as by-product from industrial processes. Results of thermodynamic analysis of the process and experimental work, application potentials of the process in The Netherlands, and plans for future development are presented.
INTRODUCTION Hydrogen, generated from renewable sources, is likely to play a major role as an energy carrier in the future energy supply. Due to the finiteness of fossil energy sources, and the global environmental damage caused by them, the world has to switch gradually to other primary energy sources. In the long term, only biomass and other renewable sources such as water, wind, and sun will be available. Most of these energy sources, however, have a fluctuating character, resulting in dissimilarities between energy availability and energy demand. Discrepancies between demand and supply of energy can be solved by temporary storage of the surplus of energy as hydrogen, through water electrolysis.
405
The storage and transport of hydrogen can take place in its fkee form (Hz), or chemically bound as, e.g. methane or methanol. However, the storage and transport of hydrogen in its fkee form is more complex, and would probably require more energy than the storage and transport of hydrogen in chemical form. The liquefaction of hydrogen, for instance, takes place at -253OC.Such low temperatures can be obtained by successively compressing the gas and expanding it through a throttle valve or expansion turbine. About 30% of the thermal energy of hydrogen (LHV) is necessary as electricity for compression [I]. An additional advantage of the indirect use of hydrogen as an energy carrier is that the existing large-scale energy infrastructure for distribution and use can be utilised. Several routes for chemical storage of hydrogen have been studied [I], fiom which SNG production by biomass hydrogasification has been identified as the option with the highest energetic efficiency. As illustrated in figure 1, between the present fossil fuel-based energy supply system and the future hydrogen economy, there would be a long transition phase, during which both fossil and renewable sources of energy would be applied simultaneously. During this phase hydrogen, produced from renewable sources, might be introduced to the energy market by the biomass hydrogasification process. The use of the existing gas infrastructure for transportation of the produced SNG makes a gradual transition to a hydrogen economy possible. The process of SNG production by biomass hydrogasification has to be developed practically.
End uSer
Present situation
Transition situation
Future situation.
Fig. 1 SNG production by biomass hydrogasification. Gasification of carbon-containing feedstocks in a hydrogen atmosphere, is called hydrogasification. Hydrogasification of coal has been investigated since the 1930s in Germany, Great Britain and The United States [2]. Generally, the conversion increases
406
with increasing pressure, temperature and residence time. Carbon conversions over 80%, with a selectivity of 90% for methane and ethane, have been obtained in hydrogasification of brown coal, in a 240 tons per day plant in Germany. In the early 1980s, Steinberg et al. [2] carried out flash hydrogasification experiments with wood in an entrained-flow reactor. At pressures between 14-34 bar and temperatures between 800-1000°C, carbon conversions were over 90%. At very short residence times ( 4 s) conversion to methane dominated, whle at longer residence times the conversion to methane decreased, and the conversion to carbon monoxide increased. Several processes have been developed for production of methane-rich gases from coal, biomass, or organic solid wastes [3,4,5,6,7,8]. The required hydrogen in these processes is produced within the process [9], e.g. by gasification of residual char from the hydrogasifier (charcoal gasification case in figure 1). The use of an external hydrogen source is new [lo], and gives the possibility to apply the hydrogasification process not only for upgrading of biomass and organic wastes to a methane-rich gas, but also as a process for chemical storage of hydrogen. A simplified flowsheet for SNG production by biomass hydrogasification is shown in Figure 2.
Biomass hydrogasifier
gas clean-up
final m metha-
nation
-
water
SNG
removal
m
I H2 Fig. 2 Simplified flowsheet for SNG production by biomass hydrogasification. Hydrogen and pre-treated biomass are fed to the hydrogasification reactor. The produced gas passes a high-temperature gas clean-up section for removal of contaminants, followed by a final methanation step for the conversion of residual CO. Removal of H 2 0 from the product gas of the methanation step, results in SNG as the final product. MODELLING WORK
The complete process for SNG production from biomass and hydrogen by hydrogasification has been modelled in the ASPENPLUS simulation package. Figure 3 shows a simplified diagram of the developed ASPENPLUS model. Biomass (poplar wood) is dried in a steam dryer to the desired moisture content, after whch it enters the hydrogasification reactor, together with a hydrogen stream. The reactor is operated at 800°C and 30 bar. A C02 stream is used to pressurise the biomass to the hydrogasification operating pressure. Both the hydrogen and C02 streams are assumed to enter the system at atmospheric pressure, and are compressed (within the system) to the hydrogasifier operating pressure.
407
HRSG
-
Fig. 3 ASPENPLUsmodel for SNG production by biomass hydrogasification.
The product gas of the hydrogasifier is cooled from 800°C to the inlet temperature of the methanation section (about 4OOOC). It is assumed, that within the temperature range of 400-800°C, a high-temperature gas clean-up system can be used, in order to remove solid residues as well as gaseous impurities (HzS, HCI, HF, N H 3 ) from the product gas. The clean gas is then fed to the methanation section. The methanation is based on the ICI high-temperature once-through process, using a series of inter-cooled reactors operating at successively lower exit temperature [ 113. The product gas of the methanation section contains mainly CH4, H2, H20, and CO2. Removing H20 from this stream results in SNG as the final product, which leaves the system at high pressure. The heat released from the hydrogasifier product gas, and the heat generated in the methanation reactors, are used to generate superheated steam (40 bar and 540°C), which enters a steam turbine. A fraction of partly expanded steam is used to dry the biomass, while the remaining part of the steam is used for power generation. Steady-state integral system calculations have been performed to determine the overall mass and energy balances. Both the properties of the SNG produced by hydrogasification and the standard Dutch natural gas (Groningen quality) are presented in table 1. As can be seen, the quality of the produced SNG (calorific value and Wobbe-index) is comparable with Groningen natural gas. The Wobbe-index (MJMm3) is defined as the ratio of the gross calorific value to the square root of the relative density of a gas:
HHV
w=Jm with HHV (High Heating Value in MJ/Nm3), pg and pair (gas and air density in kg/Nm3). Wobbe-index is a measure of the amount of energy delivered to a burner via 408
an injector. The energy input is a linear function of W. Two gases of differing composition but having the same Wobbe-index will deliver the same amount of energy for any given injector under the same injector pressure. hydrogen content of the SNG is below 10 mol%. At these hydrogen concentrations, the produced SNG might be transported through the existing natural gas infrastructure without any adjustment [ 121. Table 1 Properties of standard Dutch natural gas and SNG produced by hydrogasification. Gas composition CH4 HZ COZ c2+
N2 Molecular weight LHV LHV Wobbe-index
[mol%] [moI%] [mol%] [mol%] [mol%] [kghol] [MJkgl [MJhol] [MJ/Nm3]
SNG
NG
81.6 8.7 8.5
81.3 0.0 0.9 3.5 14.3 18.6 38.0 708.3 44.2
0.8 17.4 39.0 676.1 43.9
The efficiency of SNG production is defined as the ratio of the thermal energy leaving the system as SNG, and the thermal energy entering the system as biomass and hydrogen. With this definition, and the input/output data presented in table 2, the SNG production efficiency is calculated as 8 1% (LHV).
Table 2 Energy balance for SNG production by hydrogasification. In [MW] 147.3 171.6 11.7
Biomass HZ Electricity SNG Char Total
330.6
out [MW] 11.7 257.7 20.0 289.4
The dry gas composition of the product gas leaving the hydrogasifier is given in table 3. For comparison, also gas compositions produced by direct oxygen-blown gasification (IGT Process) and indirect gasification (Battelle Process) are given [ 131. Compared with other biomass gasification processes for SNG production, biomass hydrogasification has the following advantages: (1) The gas produced by hydrogasification has a very high CH4 concentration and a low CO concentration, compared to the gas produced by other biomass gasification processes. Therefore, only a relatively small methanation step will be required to upgrade the produced gas to SNG. As methanation is a highly exothermic process, a small methanation step for hydrogasification means, that less heat will be generated in that step. In other words, a higher percentage of the thermal input would be converted to SNG.
409
(2) The required heat in the gasifier is not delivered by combustion of biomass, which is the case in the direct oxygen-blown gasification. Therefore, less carbon dioxide is produced in the gasifier. Besides, an expensive oxygen plant is not required. (3) Both the gasification reactions and the heat production by the exothermic reactions take place in the same reactor (the hydrogasifier), rather than in two separate reactors, which is the case in the indirect gasification processes, such as the Battelle Process. (4) The hydrogasification process is more compact, and has fewer components than competing processes. (5) The presence of excess hydrogen in the hydrogasifier, especially in combination with high operating pressures and temperatures, leads potentially to a reduction in tar formation. Table 3 Dry gas composition [mol%] for different gasification processes.
Hydrogasifier
CH4
50.2
c,’
co co2
H2
4.3 10.1 34.8
IGT 02-blown gasifier 17.8 1.5 16.0 39.4 25.3
Battelle gasifier 17.8 6.2 46.5 14.6 14.9
EXPERIMENTAL, WORK As part of the technical assessment of the biomass hydrogasification concept, an
experimental programme was defined in 1999 and performed at Deutsche Montan Technologie (DMT), Germany. Willow wood, and char produced from willow wood, were used as feedstock. The proximate and ultimate analyses of both feedstocks are listed in table 4.The following types of experiments were carried out: (1) Lab-scale experiments in a pressurised thennobalance facility, operating up to 50 bar and 1000°C under isothermal as well as non-isothermal conditions. (2) Bench-scale experiments in a pressurised fluidised bed facility, operating up to 5 bar and 800°C.
Table 4 Proximate and ultimate analyses offeedstocks. Proximate analysis water ash volatile Ultimate analysis C H N S
willow wood
willow wood char
[wt% ar] [wt% ar] [wt% ar]
10.0 3.5 64.8
6.0 5.1
[wt% ar] [wt% ar]
45.5 5.7 0.6 co.1
[wt% ar] [wt% ar)
410
88.0 1.o
0.8 <0.1
Several experiments have been carried out in the DMT pressurised thennobalance facility [ 14,15,18], from which some of the results will be presented below. The main parts of the thermobalance are a stainless steel reactor and a microbalance. The reactor, in which the sample is placed, can be heated electrically either to a constant temperature for isothermal runs or with a constant heating rate for non-isothermal runs. First, the sample is weighted and placed within a basket, which is used as sample holder. The reactor is then pressurised and heated up to the requested pressure and temperature in isothermal runs. After that, the basket, which is connected with the microbalance via a chain, is introduced into the hot reactor by moving down the chain. For chain moving a motor-driven system is used. Reaction of the sample starts with an uncontrolled heat up of the sample, after which it continues at isothermal condition. During the tests the reaction gas is flown through the reactor from bottom to top. The gas stream, including the reaction products, leaves the reactor at the top and is cooled in a heat exchanger where liquid products are condensed and sampled. Thereafter, a part of the gas stream is continuously fed to a gas chromatograph for online analysis of gaseous products. A gas meter measures the total gas stream. If reaction occurs too fast for on-line GC-analysis, gas samples can be taken from the gas stream during the run by an automatic gas sampler for later GC-analysis. Typical amount of feed material was about 500 mg willow wood and 350 mg char. The particle size of the feed material used was less than 0.7 mm. As gasifjmg agent hydrogen, carbon dioxide, steam, or helium was used. The tests were run over a period of 3 hours or longer. Only in case of the steam gasification of char the runs were stopped after about 80 minutes, since the gasification reactions finished after about 40-50 minutes. In most other cases, the running time was prolonged (>3 hours) until a high degree of constancy in weight signal was obtained. At the end of the experiments, the concentration of the main gas components produced was always lower than the detection limit of the gas chromatograph ( e g the detection limit of methane is 40 ppm). Gasification is a two step process, consisting of pyrolysis of the fuel (through primary as well as secondary reactions), and gasification of the char, originating from pyrolysis. The overall reaction of pyrolysis consists of the conversion of the fuel into char, tar, and gases, especially aliphatic hydrocarbons, CO, C02, and water. The gasification reactions of char are slower than the pyrolysis reactions. During the pyrolysis step of biomass hydrogasification, whch took place in the first minutes of the thermobalance experiments, a high percentage of biomass was converted, from which a high fraction to methane and ethane, especially at high pressures. Figure 4 and figure 5 present the biomass conversion, and the release rate of the main carbon-containing components of the product gas at 850°C and two different pressures (1.5 bar and 30 bar). In both cases, a biomass conversion of about 80 wt% (as received) was achieved within a minute. The figures show a shift from carbon monoxide and carbon dioxide to methane, by increasing the operating pressure. This can be explained by a combination of methanation and reversed shift reactions: CO + 3H2 C) CH4 + H20 COz + Hz C) CO + H20 In the presence of excess hydrogen, the rate of methane formation through the methanation reaction increases by increasing the operating pressure, while at the same time carbon dioxide will react with hydrogen to produce carbon monoxide through the reversed shift reaction. The same trend is observed for the release rate of ethylene and 41 1
ethane, i.e. a shift from ethylene to ethane by increasing the operating pressure. T b s can be explained by hydrogenation of ethylene:
The reaction will shift to ethane formation by increasing the operating pressure, especially in presence of excess hydrogen. At 850°C and 30 bar, beside 100% formation of the pyrolysis products: Cz&, CO, and COz, more than 95% of ethane, and about 80% of methane were formed within the first 10 minutes. After that only methane, and in much less extent ethane, continued to be formed, through the hydrogasification of char. Figure 6 presents the mean concentration of the main carbon-containing components in the product gas as a function of pressure. It can be seen how the concentration of CO, COz, and C2H4 decreases with pressure, while the concentration of CH4 and C2H6 increases with pressure. The effect of operating pressure on final conversion of biomass and carbon, for different temperatures, is presented in figure 7. Both biomass and carbon conversions 100
80
T
60
7
40 20
0
1
2
4
3
k
3
Eg
E
- C2H4 %-
* C2H6
bmconv
.P n
- 0 5
time [min]
Fig. 4 Biomass conversion and release rate of main carbon-containing components in the product gas at 850°C and 1.5 bar. 100
C2H6
0.2
0
0,o
0
1
2
4
3
5
time [min]
Fig. 5 Biomass conversion and release rate of main carbon-containing components in the product gas at 850°C and 30 bar.
412
increase with pressure and temperature. At high operating pressures of 30-50 bar, biomass conversions of 92-94 wt% (ar), and carbon conversions of 86-88 wt?? were achieved. It should, however, be mentioned that these conversions were achieved after long test runs of about 4 hours. Figure 8 and figure 9 present the conversion of biomass and carbon, as a h c t i o n of time, for different pressures and temperatures. The test runs begin with rapid biomass conversions of 80 wt% (ar), and carbon conversions of 60 wt%, followed by hydrogasification of the remained char, with a slow reaction rate. The reactivity of char gasification was studied under three different atmospheres (hydrogen, carbon dioxide, and steam), at 800°C and two operating pressures (5 bar and 30 bar). The required char was prepared by pyrolysing willow wood in a tube kiln under nitrogen atmosphere at a temperature of 800°C, within a period of about 8 hours. The char was cooled down under inert conditions. 70 60
10 0 5
10
15
20
25
30
pressure [bar]
95
-
413
35
40
45
50
Figure 10 presents the weight loss of the char samples under different atmospheres and pressures. It can be seen, that char gasification with steam resulted in the highest conversion rate, followed by C 0 2 and H2 gasification. Based on this observation, a higher char conversion is expected during the hydrogasification of biomass in a pressurised fluidised bed reactor, because in addition to hydrogen, also the pyrolysis products, C02 and H2O can take part in the conversion of char, according to the following reactions: C + 2H2 f) CH4
c -k c02 2co f)
C + H20 t)CO + H2
90
60 50
0
60
120
180 time [min]
240
300
360
Fig. 8 Biomass and carbon conversions at 850°C as functions of time for different operating pressures.
100
-
-ss
90 80
'e
70 -A- Biompn, cony. WO'C
4 C a h n fonv. B O O T +Carbon fonv. 650°C
60 50
0
60
120
180
240
300
360
time [min]
Fig.9 Biomass and carbon conversions at 30 bar as functions of time for different operating temperatures.
414
100%
80% -0-5 badH2 ci
t"
-0-30 bar/HP
+5 badC02
40%
c 0
t30 bar/C02
20% 0%
-A-
,
4 0
5 badsteam
t 3 0 barldeam
60
120
180
240
300
360
time [min]
Fig. I0 Weight losing curve of char (prepared from willow wood) at 800°C, different atmospheres and pressures. Three experiments were carried out in a bench-scale pressurised fluidised bed gasifier (D=8 cm, H=130 cm), operating at 5 bar and 8OO0C [16,17,18]. The experimental time was divided into 2 phases, each of 120 minutes. During the first phase, about 1 kg of the willow wood, together with feeding gaslgasifjmg agent were fed into the gasifier, while in the second phase no biomass was fed any more. In order to improve the mass and heat transfer, the fluidised bed reactor was filled with approximately 5 kg of bed material (molochite), corresponding to a mean fluidised bed height of 90 cm. Willow wood with a particle size of 0.7-2.5 mm was used. As feeding gadgasifying agent hydrogen, nitrogen, or carbon dioxide was applied, with purities above 99.99 ~01%. Comparing the pressurised fluidised bed gaslfication tests in H2 and in N2 (pyrolysis) atmosphere, the specific CO formation (m3STPkgfeed mf) of H2 gasification is 54% higher than in the case of pyrolysis with N2. T h s can be explained by the reversed shf?reaction. This reaction results also in water formation, which can lead to steam gasification. On the other hand, the specific CH4 formation is 75% hgher in the case of H2 gasification. Besides, the carbon content of the product gas related to the carbon content of the feedstock in the case of H2 gasification is, with 70%, higher than in case of pyrolysis, with only 59%. This might be explained by a combined gasification of char with hydrogen and with steam, the latter produced during the reversed shft reaction, as mentioned above. Regarding the PFB gasification of willow wood with C02, the specific CO formation is 60% higher than in the case of H2 gasification, and 146% hgher than in the case of pyrolysis with N2. The carbon content of the product gas is with 94% the highest of all PFB test runs. These two points show clearly that the Boudouard reaction has taken place. A part of the CO formed has probably reacted with H20 through the shift reaction. With this reaction it can be explained that the specific H2 formation in the case of C02 gasification is 25% higher than in the case of pyrolysis with N2. In summary, the experimental programme has shown the feasibility of biomass hydrogasification as the most important step within the total process of SNG
415
production by gasification of biomass in a hydrogen atmosphere, with respect to the following aspects: (1) Production of a gas, very rich in methane at the same process conditions (p,T), applied within the previous modelling work. (2) Conversion of a sufficient amount of biomass to gaseous product, withm a reasonable residence time of the biomass feedstock. The remaining char could be used within the process or to generate steam. The investigation of the autothermal operation of the hydrogasifier requires a detailed examination of the process heat balance. Concerning the thermobalance and PFB tests, neither the input heat of the electric heating system, nor the heat losses, are known. However, based on the experimental results, it seems that the methanation reaction is a major heat source in the hydrogasifier, followed by other hydrogenation reactions. The product gas in the thermobalance tests, as well as in the PFB tests is highly diluted with hydrogen. This is not desired in practice, due to the following reasons: (1) In case of a diluted gas, hydrogen should partly be recirculated, resulting in higher costs and process complexity. (2) Presence of too much hydrogen in the hydrogasifier might result in a process, whlch cannot be operated autothermally any more.
Therefore, this aspect has to be taken into account when designing a biomass hydrogasifier.
ECONOMIC ASPECTS According to the Dutch energy policy, presented at the end of 1995 [19], the contribution of renewable energy to total energy consumption should increase from about 1% in 1995 to 10% (288 PJh) in 2020. With a potential of 75 PJh (avoided fossil fuel use), biomass is, with a contribution of 26%, the most important renewable source to meet the 2020 goal [20]. In an attempt to investigate the potential of the biomass hydrogasification process for the Dutch energy supply in 2020, rough estimations have been made for the SNG production costs [21]. It is concluded, that the hydrogen price plays a dominant role in the SNG production costs, followed (to a much less extent) by the price of biomass. By utilisation of cheap hydrogen, the cost of SNG production would be of the same order as the price of standard natural gas [21]. In other words, the large-scale production of substitute natural gas by hydrogasification might be economically feasible, when in the future, large amounts of cheap electrolytic hydrogen would become available for the energy market. However, in some specific cases, the process might also be applicable in the present and short-term situation. T h s is especially the case, when hydrogen, or a hydrogen-rich gas, is available as by-product from industrial processes. For The Netherlands, the hydrogasification process is also an interesting option to import renewable energy from parts of Europe where hydropower or solar energy, and biomass are available to a large extent as natural resources. Another option is to trade in (green) certificates, by converting biomass with hydrogen to SNG, at the location where biomass is produced. This is in accordance with the Kyoto agreements,
416
concerning the recoption of the role of emission trading and other economic flexibility mechanisms.
FUTURE WORK ECN has defined an RD&D programme for SNG production by biomass hydrogasification process, with the objective to prepare a future demonstration in The Netherlands. The first phase of this programme comprises an assessment of t e c h c a l and economic prospects of the concept. For final evaluation of this phase, it is important to integrate the experimental results within the developed ASPENPLUS model. Kinetic data obtained from the experiments should be translated into a practical reactor system, in which gasification will have to take place. As a result of exothermic reactions (in particular the methanation), taking place in the hydrogasifier, the gasification process has the potential to be operated autothermally. Modelling studies should show at which process conditions this might be possible. Also, from the modelling studies it should follow, whether the process can be operated in 'once-through' mode, or, due to dilution of the product gas from the gasifier, recirculation of hydrogen will be necessary. Beside attention to the gasification section, also other system components have to be studied in more detail. Important system components are: feedstock pre-treatment (drymg and sizing of biomass, compression of hydrogen stream), membrane separation of hydrogen (if necessary), gas clean-up, methanation, and the conditioning section, whch is required to bring the product gas from the methanation step to the natural gas quality. CONCLUSIONS (1) Based on the thermodynamic and flowsheeting analysis, this work has shown the feasibility of SNG production by biomass hydrogasification. Addition of an external hydrogen stream to the hydrogasification reactor results in a final product gas with a quality comparable to that of standard Dutch natural gas. With a hydrogen content below 10 mol%, the produced SNG can be transported through the existing natural gas infrastructure without any adjustment. The energetic efficiency of SNG production from biomass and hydrogen by hydrogasification is calculated as 8 1 % (LHV). (2) The experimental programme has shown the feasibility of biomass hydrogasifcation as the most important step of the total process, with respect to the: a) production of a methane-rich gas, at the same process conditions (p,T), applied in the modelling work; b) conversion of a sufficient amount of biomass to a gaseous product, within a reasonable residence time of biomass feedstock. (3) For final evaluation of the biomass hydrogasification concept (proof of principle), it is im ortant to integrate the experimental results, withm the developed ASPEN' model. (4) Large-scale production of SNG by biomass hydrogasification might be economically feasible when, in the future, large amounts of cheap electrolytic hydrogen will become available for the energy market. However, in some specific cases the process might also be applicable in the present and short-term situation.
P
417
This is especially the case, when hydrogen or hydrogen-rich gases are available as by-product from industrial processes. For The Netherlands, biomass hydrogasification may also be applied to import renewable energy, or to trade in (green) certificates.
ACKNOWLEDGEMENTS Financial support from the Netherlands agency for energy and the environment (Novem), and Gasunie, the Dutch gas supply company, are gratefully acknowledged.
REFERENCES 1. M. Mozaffarian, M. Bracht: Indirect hydrogen: Evaluation of indirect hydrogen energy systems for a sustainable energy supply. ECN-CX--98-110, Netherlands Energy Research Foundation (ECN), Petten, August 1998. 2. M. Steinberg, E.W. Grohse, and Y.A. Tung: Feasibility study for the coprocessing of fossil fuels with biomass by the Hydrocarb Process. BNL-46058, DE 9 1011971, Brookhaven National Laboratory, 1991. 3. H.J. Blaskowski: Production of pipeline gas fiom coal. U.S. patent 4,410,336, October 1983. 4. D.S. Scott: Hydrogasification of biomass to produce high yields of methane. U.S. patent 4,822,935, April 1989. 5. H.F. Feldmann: Process for converting solid wastes to pipeline gas. U.S. patent 3,733,187, May 1973. 6. H.F. Feldmann: A process of converting solid waste into a combustible product gas. patent 1408694, October 1975. 7. H.F. Feldmann: Process and apparatus for converting solid wastes to pipeline gas. U.S. patent 4,005,994, February 1977. 8. H.F. Feldmann: Apparatus for the production of methane containing gas by hydrogasification. U.S. patent 4,152,122, May 1979. 9. M. Mozaffarian: Hydrogen conversion in synthetic natural gas by biomass Netherlands Energy Research Foundation hydropyrolysis. ECN-CX-98-148, (ECN), Petten, December 1998. 10. Werkwijze voor de conversie van waterstof in synthetisch aardgas. Octrooiaanvrage Nr. 1010288, ECN, October 1998. 11. Catalyst Handbook. second edition, edited by Twigg M V , ISBN 1874545359,1996. 12. J.P.A. Okken: Waterstof energie-toepassingen; een compilatie van mogelijke technieken in de toekomstige Nederlandse energiehuishouding. ECN-C--92-065, Netherlands Energy Research Foundation (ECN), Petten, October 1992. 13. A.V. Bridgwater, G.D. Evans: An assessment of thennochemical systems for processing biomass and refise. ETSU B/T1/00207/REP, 1993. 14. M. Kaiser: Hydrogasification of willow wood in a pressurised thermobalance, DMT (Essen), July 1999. 15. M. Kaiser, M. Mayer: Gasification of willow wood in a pressurised thermobalance, Final-report DMT (Essen), November 1999. 16. M. Kaiser, B. Piepiorka: Hydrogaszjkation of willow wood in the bench scale pressurisedfluidised bed gasijer, DMT (Essen), July 1999.
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17. M. Kaiser, B. Piepiorka: Inert gaslfication of willow wood in the bench scale pressurisedfluidised bed gasifier, DMT (Essen), December 1999. 18. A. Streier, M. Kaiser, M. Mayer, and B. Piepiorka: CO, gasification of biomass, DMT (Essen), February 2000. 19. Derde energienota. Dutch Ministry of Economic Affairs, SDU publishers, The Hague, 1995. 20. Duurzame Energie in Opmars, Actieprogramma 1997-2000. Dutch Ministry of Economic Affairs, The Hague, March 1997. 21. R. van der Woude: Met hydrovergassing 10% groen gas (in 2020). ECN-7.2863GR3, Netherlands Energy Research Foundation (ECN), Petten, November 1998.
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A Study of Carbon Dioxide Mitigation Effect by Biomass Energy Plantation for Electricity and Methanol Shin-ya Yokoyama’ and Kiyotaka Tahara2 National Institute for Resources and Environment, 16-3, Onogawa, Tsukuba, Ibaraki 305-8569, Japan Department of Industrial Chemistry, Keisei University, 3-3-1, Kichijoji Teramachi, Musashino, Tokyo
’
ABSTRACT:A study for life cycle assessment of bioenergy production was made on the basis of energy production. Biomass such as fast growing trees is planted for energy production for electricity and methanol. The plantation area and biomass productivity are assumed to be 160,000 ha and 10 dry tons per ha per annum. For power generation by biomass combustion, it is assumed that the efficiency of power generation is 22% and the t h e m 1 efficiency for methanol synthesis from synthetic gas through gasification is 53%. In the case of power generation, about 466,000 ton-C (on carbon basis) can be avoided as an alternative of coal [ 11. In the case of methanol synthesis, about 379,000 ton-C (on carbon basis) can be avoided as an alternative of natural gas. Same analysis has been made under TASK 25 of IEA Bioenergy and information exchange will be needed together with quantitative comparison [2]. BIOMASS POWER GENERATION Biomass energy plantation is a conception to apply biomass more actively in COz mitigation. It is an idea to repeat planting tree and harvesting the mature tree, then burning the tree and converting the heat into electric power. In the electricity generation using fossil fuels such as coal, oil and natural gas, the emission of CO2 is unavoidable. If the energy of the renewable biomass is used to substitute fossil fuels, the use of the latter can be diminished. That is, the use of biomass replaces fossil fuels for energy. Though C02 is once emitted to the atmosphere when biomass is burned, it is fixed again as biomass as long as an equal amount of trees are planted susutainably. Because of this, biomass is called to be carbon neutral. In the following, we have tried to calculate the input energy necessary for plantation and the energy produced by the biomass, and have quantitatively investigated a mitigation effect of COZ emission from the viewpoint of LCA (Life Cycle Assessment) based on the energy in power generation using biomass instead of
420
coal.
PLANTATI 0N Here we thmk about a cycle of 6 years of harvesting and planting in fast-growing trees llke eucalyptus and poplar. Table 1 shows the features of trees for plantation. On the scale of plantation, an area of 160,000 ha in a 40 km-square is realistic. The growth of biomass is set to be 10 tondyear in dry weight and the carbon to be fixed is about 5 tons. The heating value of the biomass is assumed to be 4,500 kcal/kg. This value is a half of 20 tons, the maximum growth speed of eucalyptus so far reported in the tropics and the subtropics, and is a value possible to be achieved technically. Besides, the growth rate of sugarcane in Brazil and in the Central America is reported to be 50 tonslyear per hectare in dry weight, so the above value is considered to be proper.
Table 1 Properties of biomass. species
eziculypnis or hybrid poplar
productivity [t-dryihdy] heating value [kcai/kg] carbon content [ -1 rotation of harvesting [y]
10 4500 0.5 6
Table 2 shows the energy input, that is, diesel oil, natural gas, and electricity, for plantation, reported by Turhollow et a1 [3]. If carbon dioxide emission from the consumption of diesel, natural gas, and electricity is available, total carbon dioxide emission from the plantation can be obtained. In case of Japan, carbon dioxide emission from diesel oil, natural gas, and electricity is 71.46kg-C02/GJ, 51.56 kgC02/GJ,and 121.64 kg-C02/GJ,respectively.
Table 2 Annual energy input for sustainable plantation system. average ener:y per hectare(GJ) estab lis hment fertilizer
diesel oil
natural pas
electricity
total
0.14
-
-
0.14
0.34
2.81
0.28
3.33
0 29 0.17
0 10
0 02
0 41
-
0.17
7.3 1
-
-
7 31
2.40
-
-
2.40
10.55
2.90
0.30
13.76
~
chemcals machinery ~~
~
~
~~
harvest ing ~
hauling total energy
42 1
POWER GENERATION BY BIOMASS
Let us think about the case of power generation by wood using an area of 40 kmsquare for plantation. Table 3 gives the outline of this system. If one-sixth of the total 160,000 ha is cut each year, the wood is 1,600,000 tons, and the heat generated is 7.2 x lo9 k W y e a r based on the assumption of heat generation of 4,500 kcallkg in biomass. If this heat is used in power generation, a power about 1.84 x 10gkWyear is generated provided that the efficiency of power generation is 22% and the operational rate of a power plant is 60%. This is equivalent to a power plant of 350,000 kW. The efficiency is assumed to be lower than that of coal-fired power generation, it will be improved by the advanced technology such as IGCC in future. Table 3 The Outline of the biomass energy plantation system. ~
plantation area [ha]
1.6 X lo5
Net efficiency of power genaration [“A] operational rate of power generation [%] havesting area per annum [ha] wood production area per annurn[t/y] heating value of biomass per annum [Mcal/y] electricity supply per annum [ k W y ] plant size [kW] electricity for wood chipping [kWh/y]
22 60 2.67X 10‘ 1.6 X lo6 7.2 x 109 1.84X 109 3.49 x 105 1.33 X 10’
net electricity supply per annum [ k W y ]
1 . 8 4 109 ~
EFFECT OF C02 MITIGATION Let us consider the effect of C02 reduction in power generation with biomass produced by sustainable biomass plantation. Table 4 summarizes the results. The C 0 2 emission in plantation is a value estimated according to the assumption described previously. Surely the construction of a power plant is necessary for biomass power generation, and the energy required in the building of a power plant and the accompanying C02 emission should be calculated. On this calculation, a power plant for biomass is supposed to be similar to that for coal power plant. That is, with a power plant of 1,000,000 kW for coal thermal power as standard, the one for biomass is in proportion to 0.7 power of the scale. The C02 generated in the material supply necessary for the building of a power plant of 1,000,000 kW has been calculated by Uchiyama et a1 [4]. In the power plant building, steel, aluminum and concrete are needed and C02 is emitted during the manufacture of these materials.
422
Table 4 C02 mitigation effect by biomass power plant olantation area rhal ~~
1.6 X 1 O j
~
C 0 2 emission b y plantation [t-C02/y] 1.70X lo5 C 0 2 emission by power plant construction [t- COziy] 2.56 X lo3 _--_----__--_-_-______________________ total C 0 2 emission [t-COz/y] 1.73 X 10’ C02 reduction by replacement with biomass power generation [t-COz/y]
1.88 x 106
net COz-reduction [t-C02/y]
1.71 X lo6
net carbon reduction [t-C/y] net carbon reduction per hectare [t-C/ha/y]
4.66X 10j 2.94
In thls paper, all materials needed for the construction of power plant are domestically provided. The C02 emission accompanied with the material transportation is thus counted after departing from the harbor abroad. COz emission accompanied with the production of steel, aluminum, and concrete is 1.180 kg-C02/kg, 2.035 kg-C02, and 0.099 kg-C02, respectively. Besides, the lifetime of a power plant is supposed to be 30 years. As shown in Table 4, the C02 emission during the building of a power plant is 2,560 tons/year. The C02 emitted in the building of a power plant is about 1.7% of that emitted in plantation. Therefore, the C02 emission during the building of a power plant is only several percent of that emitted in plantation including the emission during the mining, transporting and refinement of the respective materials in their producing place. So how is the net effect of C02 mitigation in biomass power generation? It is calculated by subtracting the C02 emission in plantation and in the building of a power plant from the C02 emission produced when the electric power generated in a biomass power plant is replaced with a coal thermal power plant. As shown in the bottom column of Table 4, an energy plantation of 160,000 Myear results in a mitigation of 1,710,000 tons of C02, or 466,000 tons of carbon. This means a reduction of about 3 tons of carbon per hectare a year. It indicates that about 3 tons of carbon can be substituted for coal by energy plantation using biomass with a fixation ability of about 5 tons carbodhdyear. In other words, about 466,000 t-C (on carbon basis) can be avoided per annum as an alternative of coal as long as biomass is continuously used in power generation through continual planting.
METHANOL SYNTHESIS At the present moment, almost of methanol has been produced from natural gas through steam reforming followed by synthesis from hydrogen and carbon monoxide. Methanol is a promising fuel in future, because it is easily handled as liquid fuel, compatible with the existing infrastructure of motor vehicles, and wide application for direct burning and fuels for fuel cells. Furthermore, it is widely used for bulk chemicals. If methanol can be synthesized by hydrogen and carbon monoxide produced by biomass gasification, the consumption of natural gas can be reduced and thus C02 emission can be also reduced. Figure 1 shows the schematic flow of methanol
423
production from biomass. In this case, electricity is supposed to be supplied by the natural gas power plant. The biomass properties and plantation size are the same as the case of power generation. As for the COz emission from respective energy and materials, the same value is applied to this methanol production system. biomass
blomasr
transport
plantatlon
storage
use
Fig.1 Schematic flow of methanol production from biomass. Although many processes are reported for biomass gasification, BCL (BattelleColumbus Laboratory) Process is applied for this analysis [5]. According to the report, the system is fast-fluidized bed and indirectly heated. The temperature is 863 C and the operating pressure is 0.101 MPa. As for product gas characteristics, the yield of (kmoVton-dry feed) is 45.8, and HHV ( M J / N d raw gas) is 15.19. Cold gas efficiency is 80.1%. Table 5 The outline of methanol production system.
plantation area [ha] wood Droduction per m u m r t h l conversion to methanol r%l
1.6 X 10j 1.6X lo6
production of methanol [t/d]
ca 2500
j2.a~~
The COz is emitted from plantation, methanol plant construction, methanol transport, natural gas power plant and the total COz emitted is 2.88 x lo5 tons per annum The effect of COz mitigation by replacing natural gas with biomass is estimated to be 1.68 x 10 tons per annum. Therefore, their difference is the net mitigation of COz, which is 1.39 x lo6 ton per annum.This means a reduction of about 2.4 tons of carbon per hectare a year. It indicates that about 2.4 tons of carbon can be substituted for natural gas by energy plantation using biomass with a fixation ability of about 5 tons carbon per hectare per annum. In other words, about 379,000 t-C (on carbon basis) can be avoided as an alternative of natural gas per annum.
424
C02 mitigation effect by methanol production from biomass
Table 6
1.6 X 10'
plantation area [ha]
C02 emission by plantation [t-COz/y] 1.7X lo5 C02 emission by methanol plant construction [t-COz/y] 1.59X 10" emission by methanol transport [t-CO?,/y] 3.05 x 104 COz emission from natural gas power plant [t-C02/y] 6.82 X 10" C02 emission by natural gas production [t-COz/y] 3.63 x 103 .............................................................................................. total C02 emission [t-C02/y]
2.88 X lo5
C02 reduction b y biomass replacement [t-C02/yl
1.68 X lo6 1.39X lo6
net CO:! reduction [t-COz/yl ~~
~
~~~
~
~
~~
net carbon reduction [t-C/y] net carbon reduction per hectare[t-C/ha/y]
~~
3.79X 10' 2.37
REFERENCES 1.
2. 3. 4. 5.
Tahara K., Kojima T., Inaba A., Ogi T. and Yokoyama S. (1998), Reduction of C02 emission by biomass power generation with sustainable afforestation Evaluation by LCA -,J. Jpn. Inst. Energy., 77(5), 403-409 IEA Bioenergy, (www.ioanneum.ac.at/iea-bioenermtask 25) Turholow A. F. and Perlack R. D., (1992) Emissions of C02 from energy crop production, Biomass and Bioenergy, 1(3), 129-135 Uchiyama Y. and Yamamoto H., (1902) CRIEPZ Report, Y91005 Johansson T. B., Kelly H., Reddy A. K. N. and Williams R. H., (eds) (1993), Renewable Energy, Island Press
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A small-scale stratified downdraft gasifier coupled to a gas engine for combined heat and power production M. Barrio, M. Fossum', J.E. Hustad Norwegian University of Science and Technology, Department of Thermal Energy and Hydro Power, 7491 Trondheim, Norway 'Sintef Energy Research, Department of Thermal Energy, 7465 Trondheim, Norway
ABSTRACT: A small scale (30 kW thermal input) stratified downdraft gasifier is erected in the laboratory at the Norwegian University of Science and Technology (NTNU).The gasifier will be coupled to a gas engine to produce heat and power from biomass fuels. The paper describes the gasifier in detail (500 mm height and 100 mm diameter). One of the singularities of the gasifier design is that it allows for variation of the point of air injection and that air preheating is also possible. The gas engine was originally a diesel engine but it has been modified for producer gas and/or natural gas operation. These changes mainly affect the compression ratio and the fuel injection system. The paper describes the gas engine and explains the modifications. Experiments have been performed of gasification of wood pellets. The feeding rate was about 5 kg/h, giving an effect of 30 kW. The amount of air supplied to the reactor has been varied in the experiments, in addition to the location of the supply. The fuel gas composition has been measured with a GC.The amount of product gas obtained is about 12.5 Nm3/h and has a heating value of 4.9 MJ/Nm3. From these data, the power produced by the gas engine is expected to be about 5 kW,. The gas engine will be operated with mixtures of synthesis gas and natural gas and detailed measurements of cylinder pressure, compression ratio and heat released by the engine are planned in addition to emission measurements of CO, unburned hydrocarbons and NO,. The dependency of the results on the ratio of synthesis gadnatural gas will further be evaluated. INTRODUCTION The objective of this project is to produce heat and power at small scale from biomass fuels. Biomass gasification represents a competitive alternative to direct combustion for optimisation of the electricity production. Fixed bed gasifiers are known to produce a gas with low tar and particulate content, and therefore suited for small scale power production'*2. The possibility of combining producer gas and natural gas is also of particular interest and specially for Norway with large resources of natural gas. The plant consists of a stratified downdraft gasifier, a cleaning system and a gas engine. The thermal input to the gasifier is 30 kW, supplied by 5 kg/h of wood pellets. The design of the gasifier is particularly flexible. It allows the use of different gasification agents, different location of the gasification agent supply and almost every
426
part can be replaced. At the first phase of the project, the gasification agent will be air. This small scale reactor is suitable for laboratory work but one of its drawbacks is that the heat losses are very high compared to the amount of energy involved in the process. As a first approach, a deep bed filter will be used for gas cleaning, with sand as the filter media. Later, a more advanced high temperature filtration method will be employed3. The gas engine was originally a diesel engine that has been modified for synthesis gas or/and natural gas. A relevant feature of this project is the measurement system that has been implemented in the engine. It will allow measurement of the cylinder pressure and the crank angle, providing thus relevant information for the understanding of engine operation with producer gas. This paper describes the gasifier and the engine. The first experiences with the gasifier are also presented as well as the modifications required by the diesel engine in order to operate with producer gas and/or natural gas.
SMALL SCALE STRATIFIED DOWNDRAFT GASIFICATION Downdraft gasification generally produces a low particulate and low tar gas and is therefore suited for power generation in small scale applications. There exist mainly two designs for downdraft gasifiers: the Imbert gasifier and the stratified downdraft gasifier (or open core gasifier). The Imbert gasifier is usually a cylindrical reactor with a constriction near the bottom. The gasification agent is injected just above the constriction, creating a high temperature zone. The stratified downdraft gasifier has no constriction and both feedstock and gasification agent are fed from the top of the gasifier. The air flows through the fuel bed supported by the grate at the bottom of the gasifier. Although much more popular, the Imbert gasifier presents certain disadvantages like the fuel limitations due to the constriction and the difficulties at scale up due to the radial air supply. The stratified downdraft gasifier is easier to construct and the different zones are easier to access. This gasifier has in principle better scale-up properties than Imbert gasifier~~,’*~. In the stratified downdraft gasifier four zones can be distinguished. At the top of the bed the pellets are heated and dried. Below this zone the temperatures start to be higher and the pellets release their volatile matter (devolatilization, or pyrolysis). The gaseous products from devolatilization are partially burnt with the existing air. This phenomena is called flaming pyrolysis5 and it is the source of heat for the drying and pyrolysis already mentioned as well as for the subsequent char gasification. The temperature in this zone can reach 1000-1100 “C and since it is a thin zone it has been called the “pyrolysis front” in this investigation. In the third zone, the char reduction zone, the hot gases formed in the flaming pyrolysis zone (mainly COz, HzO) react with the remaining char in absence of oxygen at around 800-900 “C. The char is converted into the product gas mainly by the following endothermic reactions: Boudouard reaction: C + COz t)2CO Water gas reaction C + H 2 0 H H2 + CO
AHR = + 160,9 kJ/mol AHR= + 118,4 kJ/mol
The hot gases and the char coming from the flaming pyrolysis zone provide the energy required. As these reactions proceed the temperature sinks progressively until it becomes so low (700 “C) that the reaction rates are insignificant. This means that the
427
extent of the char reduction zone is dependent on the amount of ener y entering the reduction zone and consequently also on the heat losses from the reactor . The bottom zone, or inert char zone, consists then of char that has not reacted because the temperatures are too low- and ash. This zone is nevertheless very important since it acts as a reservoir of charcoal that can absorb heat when conditions of operation will change and moreover acts as a particle filter. The length and position of the zones described above depend on numerous parameters that interact with each other: pyrolysis rate, gasification rate, rate of asNinert char removal, temperature profile, heat available for reaction, fuel feeding rate, air feeding rate, heat losses, etc. It is still a challenge to understand how the zone's distribution inside a gasifier affects the quality of the product gas: gas amount, composition, tar content and particle content. This challenge has been faced through experimental work and through modelling'. Di Blasi' has recently presented a dynamic model for stratified downdraft gasifier where finite rate kinetics for the chemical reactions are included: pyrolysis, gas-phase combustion, gas-phase water shift and heterogeneous char reactions. Another important matter is the stability of the gasifier operation; the zones can move, as observed by several researcher^"^. This fact might affect the product gas quality and this is unwanted while operating the engine. The raw pellets travel downwards inside the reactor at a rate that can be obtained from the feeding rate, assuming a constant bed height:
!
Vpellcrs
- feeding rate(kg I h ) * pellets bulk density( m3I k g ) reactor diameter(m2)
(1)
Once the pellets have crossed the pyrolysis zone, they continue to travel downwards at the same speed (vpellers)if one assumes that the bulk density of the pyrolised pellets is the same than the one of the raw pellets':
The mechanical strength of the fuel is relevant to maintain certain porosity of the bed to prevent plugging6. In the char reduction zone, the pyrolysed pellets are partially transformed to gases through the gasification reactions. The remaining pellets (inert char) and the ashes leave the bed through the grate. This can be expressed as:
The pyrolysis front moves towards the raw pellets, as a flame front, as already suggested by Reed et al.''. Flaming pyrolysis is a fast reaction at around 1000 "C and the length of this zone is relatively small and constant. Therefore, the zones will not move if the speed of this pyrolysis front matches the velocity of the bed moving down, in other words, if the pyrolysis rate is equal to the char consumption rate. I The dimensions of the pyrolysed pellets are approximately 90%of those of the raw pellets.
428
’
front
where
- ’pyrolysis - ‘char consumed -
absolute velocity of the pyrolysis front vpyrolysis velocity of the pyrolysis front relative to the moving bed vfron,
If the pyrolysis rate is higher than the char consumption, then the pyrolysis front moves up (vfronr>O)and the length of the char reduction zone increases’. If this is the case, the flaming pyrolysis front will climb indefinitely until it reaches the top of the bed and the drying zone almost disappears. Although such operation mode is stable, the radiation from the top of such bed is a significant heat loss’. This operation mode is called “Top stabilisation mode” and has been modelled by Di Blasig.
DESCRIPTION OF THE STRATIFIED DOWNDRAFT GASIFIER The gasifier and the surrounding systems are illustrated in Fig. 1 and Fig. 2. This section gives a detailed description of each element.
FEEDING SYSTEM The feeding system consists of a hopper and of two pneumatic sliding valves. This system provides a pseudo-continuous operation of the reactor. Since the two valves never open simultaneously it is a tight system. A screw feeder was tested earlier but it crutched the pellets. The feeding can be done manually or automatically. Several tests have been conducted to calibrate the feeding system, so both the weight and bulk volume of each charge of pellets is estimated. Pellets have quite uniform properties regarding size, composition and mechanical strength and are therefore better suited for small scale applications.
AIR SUPPLY The pressurized air network at the Department’s workshop (6 bar) supplies the air for the installation. A filter removes the eventual dust and moisture from the air and thereafter the air pressure is reduced to 3 bar. The air is then stored in a manifold that distributes it to 7 independent channels. The amount of air through each channel is measured with a rotameter and controlled with a valve. In addition, an air preheater has been installed before the manifold. Each air channel supplies air to a circular distributor around the gasifier. From each circular distributor five pipes evenly distributed supply air to the reactor. Such system guarantees an even air distribution, for any combination of channels in use. The location of the air levels is shown in Fig. 3, together with the position of temperature and pressure measurements. Contrary to the usual design where the gasifier works under atmospheric pressure, the air pressure at the reactor is slightly above atmospheric pressure, depending on the pressure losses downstream the reactor. With this design, the amount of air entering the gasifier is well controlled although there is an increased risk for leakages.
429
Pellets
Pneumatic valves
Air distribution
circular pipe
Window Air distribution circular pipe Air tubes
Grate Solids separation chamber
Gasout
crank Ash collector
Fig. I Sketch of the stratified downdraft gasifier
430
llets Pmduct Gas Outla
Ash Collcnor
Gas Sampling
Lim
Fig. 2 Sketch of the gasification unit
Fig. 3 Location of air levels (Ai), temperature (Ti) and'pressure(Pi) measurements.
43 1
REACTOR The reactor itself is a cylinder of 100 mm diameter and 500 mm height, made of refractory ceramic. It has 35 holes for air supply, 7 for temperature measurement and 4 for pressure measurement. The reactor was originally designed with a glass window, so the zones could be observed. It was possible to observe the zones through the window but it was necessary to permanently cover it due to leakages. A steel cylinder of 250 mm surrounds the reactor, with glass wool filling the space between both cylinders as insulator. The steel cylinder is also covered with approx. 7 cm of insulation. Between the feeding system and the reactor there is a view port that is used to ignite the bed and acts as well as a safety pressure release valve. The small dimensions of the reactor pose problems like for example that there is no agitator to avoid bridging, that a bed height indicator could be a disturbance and the same with the thermocouples. But on the other hand, such small diameter almost guarantees radial uniform temperatures, and uniform air distribution. The bottom section of the gasifier accelerates the gas flow so that the char and ash particles will more likely deposit at the ash collector while the gas exits the reactor. The grate is placed between the reactor and the bottom section. It is a perforated plate of 10 mm thickness, with a crank so it can be shaken manually. As experienced by other researchers, the grate is problematic. Its design has been changed several times either because the char losses were too high or because it blocked and it has been the most challenging element of the installation. Its correct operation and a controlled ash removal is vital for the stable operation of the gasifier". When the grate blocks, the amount of char removed is zero and this also affects the velocity of the pyrolysis front, as shown earlier.
MEASUREMENT EQUIPMENT The temperature inside the reactor is measured with seven thermocouples (Type K). To avoid channelling, the thermocouples are placed only 5 mm into the bed. The temperature inside the manifold is also measured, as well as at several points at the outlet pipe. The pressure in the manifold and inside the reactor is also measured with the help of pressure transducers. In order to take gas samples, a sample line has been built (Fig. 4). It consists of a steel condenser and three glass bottles in an ice bath, a moisture filter (silica gel), a pump and a gas flowmeter. A sample of 0.5 ml is taken from the gas sample line with a syringe and inserted into a gas chromatograph from SRZ Instruments equipped with a TCD detector and a Supelco column (Carboxen 1000). ; Product gas
Ice Bath Water Outlet
Ice Bath
Fig. 4 Gas sampling line.
432
SAFETY CONSIDERATIONS To allow an emergency stop, an independent pipe supplies nitrogen to the reactor. If necessary, the air flow can be rapidly stopped and substituted by nitrogen. Since there is a moderate risk for leakage, the gasifier is placed inside a room with suction system and all the control and measurement equipment is placed outside the room. CO detectors are also installed.
THE ENGINE In general both diesel and Otto engines can be converted to operate on gaseous fuels. To ensure high efficiency and low emissions it is recommended that the engine is modified according to the specific combustion properties of the actual gaseous fuel, which may be very different compared to both diesel and gasoline. The diesel engine is more attractive for conversion to producer gas as the reduction in power and efficiency is less compared to an Otto engine. This is due to the higher compression ratio of the diesel cycle and also the operation conditions with high excess air ratios which reduce the difference in the volumetric energy content of dieseYair mixtures and producer gadair mixtures. The major engine modifications for a diesel engine include reduction of the compression ratio and installation of an ignition system. The ignition system can either be a spark plug system or a system using diesel fuel in a prechamber as an ignition source for the gas. Direct injection diesel engines are more suited for producer conversion than prechamber engines due to the less heat loss to the cylinder walls, which affect the ignition of the lean producer gas. Diesel engines fuelled on producer gas are normally operated at a self aspirated mode. Contaminants in the producer gas, especially particles, can cause damages to a turbocharger. The producer gas is mixed with the intake combustion air and distributed to each cylinder by the intake manifold. For small scale integrated gasification and gas engine system the suction from the gas engine is used to feed air into the gasifier. Overheated exhaust manifolds have been reported from installations running on producer gas'*. This is most likely due to the low burning velocity of producer gaslair mixtures compared to diesel or natural gas and thus a complete burnout of the gas mixture may occur in the exhaust gas channel and into the exhaust manifold. The high temperature problem can be solved by cooling of the exhaust manifold. The engine used in this project is a 3-cylinder naturally aspirated direct injected diesel engine (Zetor 24901). The table below shows the original specifications of the diesel engine: Table I Original engine specifications. Concept Number of cylinders Displaced volume (dm3) Bore (mm) S troke(mm) Compression ratio Power @ 2200 rpm (kW) Torque @ 1500 rpm (Nm) Specific fuel consumption (gkwh)
433
3 2.7 102 110
17
34.2 150
245 + 10%
The engine was built in 1980, and was originally design for standby stationary power production so instead of an airblown radiator for cooling it is equipped with a water jacket heat exchanger, common in marine applications. The heat exchanger is cooled with cold water being then possible to quantify the amount of heat rejected from the engine. The engine’s exhaust manifold is also cooled to avoid problems with overheating, as mentioned before. The initial phase of experiments will include a mapping of the engine characteristics using natural gas as a fuel. These data will be the basis for comparison and discussion of the results for different gas mixtures. The next step will be to use synthetic producer gas as a fuel as the most extreme low value gaseous fuel. Further experiments will include mixtures of synthetic producer gas and natural gas.
COMPRESSION RATIO, IGNITION SYSTEM AND FUELLING SYSTEM The original compression ratio is too high for producer gas operation. Our intention is to operate the engine with a variety of gases, from natural gas to producer gas and thus the compression ratio is reduced down to 1l:l.I3 The engine has separate cylinder heads, what makes the modifications easier. The fuel injectors are replaced by spark plugs of the type used in small motorcycle engines. The small spark plugs are used to minimise the thermal mass of the spark plug, thus reducing the cooling effect and the possibility of flame quenching in the ignition zone. Each spark plug is connected to a high voltage coil that provides sufficient ignition energy. The diesel pump is removed, and an optical pickup for the electronic ignition system is mounted in its shaft. The rotational speed of this shaft is ?4of the engine speed, which is needed for the ignition system. The ignition timing can be accurately set with this system. A standard fuel feeding system is mounted and modified in order to be able to operate natural gas, producer gas and mixtures of gases.
MEASUREMENT EQUIPMENT A pressure sensor from Optrand Inc. specially designed for cylinder pressure measurements has been installed in one of the cylinders. In addition, a crank angle decoder has been installed which gives a step signal twice per degree. These pulses trigger the pressure measurements and with this system the cylinder pressure versus the crank angle is recorded 720 times per revolution. The decoder also generates a one step signal for every other revolution in order to identify the compression stroke of the 4-stroke cycle of the engine. The time dependent volume of the cylinder is given explicit from the crank angle and the geometry of the cylinder, connecting rod length and crank radius. Thus, indicated values for power, efficiency, mean effective pressure and rate of heat release can be found. The air flow to the engine is calculated from the pressure drop over a “honeycomb” viscous duct meter. The heat balance for the engine is calculated from measurements of the flow of water through the heat exchanger and the according inlet and exit water temperatures. In addition the temperature of the fuel gadair mixture and the exhaust gas temperatures are measured using thermocouples. Standard operation will include measurements of the oxygen concentration in the exhaust gas which combined with the measured inlet air flow can be used to estimate 434
the excess air ratio. More detailed analysis of the exhaust gas will include measurements of CO, C02, UHC and NO, concentrations. All output signals from temperature sensors and flow controllers are sampled with a commercial acquisition system connected to a PC and data are presented on the screen using a LabView set-up. For the measurements of the cylinder pressure a high speed data acquisition system is installed which also is connected to the PC. To measure the power generated by the engine, a hydraulic brake has been installed. A load cell attached to the brake measures the torque.
GAS CLEANING As a first approach deep bed filtration in planned, using sand as the filtration media. The pressure loss across the filter will be observed and when necessary, the sand will be renewed. Eventually, a pump will be installed after the filter to avoid high overpressure inside the gasifier. The engine will have a suction effect later on. Later, a more complex filtration system at high temperature will be introduced3.
PRELIMINARY GASIFICATION EXPERIMENTS Several experiments have been conducted varying the amount of air supplied and its distribution and varying the preheating conditions. These parameters affect: (1) (2) (3) (4)
(5) (6) (7)
FueYair ratio Temperature profile inside the gasifier Position of the zones Stability of the gasification process Product gas quality Product gas composition Amount of product gas.
It has been found through the experiments that the reactor needs considerable preheating. If the reactor is cold, it absorbs a large amount of the heat produced and the gasification process can hardly be established. All experiments include thus approx. 12 hours of preheating at around 200 "C. After preheating, a small amount of charcoal (small pieces) was sent through the ignition port and thereafter some pieces glowing charcoal. When the temperature starts to rise, the ignition port is closed and the reactor is filled with pellets. The air flow is adjusted to the desired amount and the experiment starts. The pellets have a diameter of 8 mm and a length that varies between 5 and 15 mm. The reactor does not have a bed height indicator. Nevertheless, an alternative simple method has been found to measure and control the fuel level. The bed height is controlled by observing the temperature at the top of the gasifier (Tl). When the gasifier is filled with pellets, a new charge of pellets only affects T1, that decreases suddenly below 100 "C. When the temperature T1 stabilises at around 180 "C depending on the preheating temperature-, the level is below T1 and a new charge is fed. If other temperatures are affected by the new charge of pellets, then the bed height is lower. In most experiments, the reactor has been kept filled with pellets. In this way, one possible parameter, bed height, is eliminated and also the heat produced in the pyrolysis zone can be absorbed by the raw pellets as well as by the char bed and not radiated to the top of the reactor.
435
Fig. 5 shows the temperature record from one of the experiments. The bed height has been kept around T1. The figure shows how the pyrolysis front moves up. This has been observed in all the experiments although the speed of the front has varied. It has been found that several operational parameters like the velocity of the front, the feeding rate and the air excess ratio depend on the temperature of the char reduction zone (Fig. 6). 1200
,
I
1000
800
400
200
0 10
0
20
40
30
50
60
90
80
70
100
time (min)
Fig. 5 Temperature record of one experiment. 25 -
1000
Q
20
V front (mmlmin)
0 lambda '10
V pellets(mm/min)
XAmounlol air (Nm3h)
- AV
pyrolysis (mm/min)
X
-.990 . 9ao
XT pyrolysis ("C)
-- 970
ul
15
-
-.
a
9 y1
-- 950
m E
>"
960
10
_.
5 -
- 940
X I
-
X
-
e ,x"
.-930 .-920
*T - - 910
0,
T
T gasification
900
('C)
Fig. 6 Influence of the temperature of the char reduction zone on the gasifier operation.
The fuellair ratio is not considered a parameter in this discussion. The pellets were supplied so the bed height was constant. This means that the fueYair ratio is not a parameter since it is not the operator who has control over it but the gasifier, as already shown by other a~thors'~.
436
The gasifier has been operated with several air distributions: (1) 100% supply from level A1 (“traditional” open core gasifier) (2) 80% supply from level A1 and 20% supply from level A4 (3) 20% supply from level A1 and 80% supply from level A4. The air excess ratio obtained for distributions (1) and (2) is about 0.4-0.45 while the distribution (3) has given a lower ratio (0.3). Distribution (3) also shows some differences on temperature profile like higher char gasification temperature. Fig. 7 shows the temperature profiles during two experiments: one with distribution (1) and another with distribution (3) but both with approximately the same amount of air supplied. Profiles la, l b and l c correspond to the same experiment, but with the pyrolysis front at different locations. The figure shows that the heat created during the flaming pyrolysis is used in a more efficient way with air distribution (3). The temperature profiles also show that the temperature below the pyrolysis front is always above 800 “C,i.e. there is no inert char zone. This can mean that the length of the char reduction zone is too short and therefore the gasification reaction is uncompleted. 500
-----
450
-3 400
-
air inlet
350
-
300
E
E
2 250 .-0 0
I
200
150
100
50
0
4
>
!
#
r
8
Fig. 7 Temperature profile inside the reactor.
The behaviour of the gasifier suggests that the reactor will operate satisfactory in “Top stabilisation mode”, partially because of the low moisture content of the pellets. This mode of operation was tested in one of the experiments, as shown in Fig. 8. Top stabilisation mode was reached after 200 minutes of operation. One can observe the large variations in the temperature at the top of the bed, as a result from the semi-continuous feeding. Pyrolysis of fresh pellets alternates with char
437
combustion at the top of the bed and this periodic behaviour affects the temperatures along the reactor, disturbing therefore the stable gas production. The height of the bed is extremely difficult to control with this mode of operation. In addition, the large and frequent temperature variations can damage the thermocouples. 1200
1000
800
400
200
I
04 0
120
80
180
240
!(min)
Fig. 8 Top stabilisation mode. Temperature record.
Therefore, the conclusion from this experiment is that top-stabilisation mode, although feasible, is not suitable for our reactor. Under this mode of operation, the semicontinuous feeding becomes an inconvenience. Further tests have been conducted having the height of the pellet’s bed above the point where the air is injected. This mode of operation has shown to be stable and more experiments will be conducted soon. The approximate gas composition of the product gas is shown in Table 2. The amount of gas produced is calculated from the Nz content of the product gas and the amount of air provided.
Table 2 Product gas composition, heating value and amount produced. Product gas composition Nz (% VO~.) 46.8 co (% vol.) 20.6 coz (% vol.) 10.2 02 (% vol.) 1.5 cH4 (% vol.) 0.0 J 20.9 Heating value (MJMm3) 4.9 12.6 Amount of product gas (Nm3/h)
CONCLUSIONS (1) A small scale stratified downdraft gasifier is erected at our laboratory. The thermal input is 30 kW,supplied by 5 k g h of wood pellets. 438
(2) A gas engine is modified so it can operate with the product gas and with mixtures of product gas and natural gas. The modifications affect the compression ratio, the ignition system and the fuelling system. (3) The preliminary experiments show that the pyrolysis front moves up; the velocity of the front depends on the char reduction zone temperature. In order to achieve stability, this temperature should be as high as possible. The air distribution can be altered so the air supplied is more efficiently use for heat production. The amount of char removed also affects the stability of the process. (4) Top-stabilisation mode, although feasible, is not suitable for our reactor. Under this mode of operation, the semi-continuous feeding becomes a problem. It is nevertheless possible to reach stable operation by keeping the bed height above the air injection point. (5) The air distribution seems to affect the air excess ratio. (6) Based on the preliminary experiments, the gasifier produces about 12 Nm3/h of product gas. This gas has a calorific value of approx. 5 MJ/Nm3 and contains approx. 20 %vol. CO and 20 %vol. Hz.
ACKNOWLEDMENTS This project is financed by the Norwegian Research Council. The authors want to thank the enthusiastic cooperation of Ole Birger Svendsgaard, Torsten Goehler and Uk Chantaka.
REFERENCES 1. Reed, T.B. & Gaur, S. (1999). A survey of biomass gasification 2000. The Biomass Energy Foundation, Inc., 1810 Smith Rd., Golden, CO. 80401. Air gasification of 2. Garcia-Bacaicoa, P., Bilbao, R. & Usbn, C. (1995). lignocellulosic biomass for power generation in Spain: Commercial plants, Proceedings of the 2nd Biomass Conference of the Americas, USA, pp. 676-694. 3. Rimes, H. & Sgnju, O.K. Evaluation of granular bed filtration for high temperature applications. This conference. 4. Xu, M., Gu, Z.Z., Sun, L., Guo, D.Y. & Han, T. (1997). Research on straw waste gasification and application in straw pulp mill, Developments in Thermochemical Biomass Conversion, Blackie Academic & Professional, 1997, pp 892-899. 5. Reed, T.B. & Das, A. (1988). Handbook of biomass downdraft gasifier engine systems, The Biomass Energy Foundation Press, 1810 Smith Rd., Golden, CO. 80401. 6. Schenk, E.P. , van Doorn, J. & Kiel, J.H.A. (1997). Biomass gasification research in fixed bed and fluidised bed reactors, Gasification and Pyrolysis of Biomass, Stuttgart, April 1997. 7. Milligan, J.B. , Evans, G.D. & Bridgwater, A.V. (1993). Results from a transparent open-core downdraft gasifier, Advances in Thermochemical Biomass Conversion, Blackie Academic & Professional, 1993, pp. 175-185. 8. Reed, T.B. & Levie, B. (1984). A simplified model of the stratified downdraft gasifier, International Bio-Energy Directory and Handbook, 1984, pp.379-390. 9. Di Blasi, C. (2000). Dynamic behaviour of stratified downdraft gasifiers, Chemical Engineering Science, Vol. 55, No. 15, pp. 2931-2947. 10 Reed, T.B. & Markson, M. (1985). Biomass gasification reaction velocities, Proceedings of FPRS industrial wood energy forum ‘83, No.7, Vol. 2, pp. 355-365.
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11. Gabroski, M.S. & Brogan, T.R. (1985). Development of a downdraft modular skid mounted biomass/waste gasification system. 12. Hansen, U. et al. (1996). Heat and power from small scale biomass plants in rural regions. A: typical application case study in Mecklenburg-Vorpommem. Proceedings of the 9" European Bioenergy Conference, pp. 13 18- 1323. 13. ImechE (1996). Using Natural gas in engines, Seminar Publication. 14. Garcia-Bacaicoa, P.,Bilbao, R., Arauzo, J. & Salvador, M.L. (1994). Scale-up of downdraft moving bed gasifiers(25-300 kgh) - design, experimental aspects and results, Bioresource Technology 48 (1994), pp. 229-235.
440
Small Scale Biomass Gasification: Development of a Gas Cleaning System for Power Generation
',
M. Walker G. Jackson and G.V.C. Peacocke3 I Biomass Engineering Ltd. and Shawton Engineering Ltd., Unit I Junction Lane, Sankey Valley Industrial Estate, Newton-le-Willows, Warrington, WA128DN,UK UK Service and Maintenance, Unit 1, Junction Lane, Sankey Valley Industrial Estate, Newton-le-Willows,Warrington, WA12,UK Conversion And Resource Evaluation Ltd., 9 Myrtle House, 5 Cassowary Road, Birmingham, B2O INE, UK
ABSTRACT: Biomass Engineering Ltd., a new company, formed from the Biomass research division of Shawton Engineering Ltd. has been working on small-scale downdraft gasification for 4 years. Two units have been operated with a net electrical output of 35 and 70 kWe using hardwoods [teak, mahogany], short rotation coppice willow and poplar as feedstocks. A complete 75 kWe gasification system, with wood preparation and spark ignition engine, has been sold to provide the electricity for a building in Northern Ireland. During testing of the gasification system on SRC willow, analyses of the product gases, wastewater and chars were made. Measurement of organics and particulates in the producer gas was performed independently by CRE, with organics content less than 15 mg/Nm3 [ 15 ppm] measured in the raw gas. Several hundred hours operation on willow and poplar have been obtained, with continuous test runs of up to 8 hours coupled to a spark ignition engine. The gas cleaning system has been developed and continually unproved during the past year. Future work will look at building on the previous four years experience with a self-cleaning gas filtration system and use of a series 1 Jenbacher engine [ 146 kWe] for an onsite gasification system for Shawton Engineering Ltd.. INTRODUCTION Shawton Engineering Ltd. started work in downdraft gasification in 1995, as there was a perceived market for small-scale biomass gasification systems. The initial designs for the gasifiers were provided by Marick International Ltd., whose owner joined Shawton as project manager for the design, operation and construction of the units. Limited but significant funding support for the project was obtained from the DTI and the two downdraft gasifiers were constructed in 1996 and operated from 1997 onwards at 44 1
Shawton Engineering. At the Biomass Engineering Ltd. premises, a specific building and engine room was constructed to allow for testing of the gasifier and gas cleaning systems, with an external engine room housing the gas engines as required. The design capacities of the two gasifiers were 35 and 75 kWe. In late 1998, Shawton Engineering Ltd. transferred it’s biomass gasification activities to a stand alone company - Biomass Engineering Ltd, and tendered to Ballymena Borough Council [BBC] [in Northern Ireland] for a contract to install a biomass powered CHP plant. This was for the generation of net 55 kWe [75 kWe installed to allow for parasitic losses], complete with gas cleaning, engine and wastewater treatment. Biomass Engineering Ltd. was successful in the bid to install the system for BBC. The gasifier was to be located in the Ballymena ECOS Millennium Park, which has the aim of demonstrating various renewable energy technologies including solar and wind. By the beginning of 1999, all connections with Marick International Ltd and its owner had been severed, due to numerous difficulties in operating the gasifiers and the inability to achieve a gas quality suitable for use in a gas engine. Two Perkins Elmer engines [ 1000 and 2000 series engines] had to be removed and sent to a local agent for complete overhauling, due to excessive tar and particulate deposition in the engines. In early 1999, Conversion And Resource Evaluation Ltd. were contracted to provide services to Shawton Engineering Ltd. to overcome the operational and gas cleaning problems. In the course of 1999, through into early 2000, changes were made to the operation of the gasifier and gas cleaning system to produce a gas acceptable for use in a spark ignition [SI] engine. This has been achieved with several hundred hours of gasifier operation on willow, poplar and hardwoods. Run times were typically 4-8 hours, although extended run times are possible, depending on the specific testing objectives. The complete 75 kWe gasification has been delivered to BBC in March 2000, with successfulre-commissioning in May 2000. The objective of this paper is to summarise the past 4 years work at Biomass Engineering Ltd., the problems encountered, how they have been solved and the results from the successful operation of the gasifier + gas engine system. Prospects for future work and development of the gasification system are presented below. REQUIREMENTS FOR THE BBC PROJECT Modifications were made to the system to meet the requirements of the Ballymena Borough Council project, the general specification of the system being: 1. Biomass feed preparation [comminution, drymg and conveying to the gasifier], 2. Operation on willow, 3. Gasifier and Engine operation for a minimum of 10 hours, 5-6 days a week, 4. Gas engine to generate 55 kWe net [75 kWe gross] for the use of the building, 5. No heat recovery, 6. Minimal maintenance. The existing 75 kWe gasifier in operation at Shawton Engineering Ltd. would be supplied for this project, as described below.
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75 kWe GASIFICATION SYSTEM As noted above, two downdraft gasification systems were constructed and operated by Shawton Engineering Ltd. Both systems proved to be flawed in their modes of operation and to some extent design and operational difficulties were encountered up until early 1999. The work has primarily focussed on the 75 kWe gasifier, as the smaller gasifier was mothballed for the duration of 1999. The basis of the original gas processing system consisted of the following sequence of unit operations: 1. Conventional cyclone to remove particulates Impingement filter to remove more particulates 3. Optional heat recovery exchanger 4. Wet/dry scrubber 5. Condensate separator 6 . Gas buffer tank 7 . Water repellent candle filters. 8. Gases transfer to the engine. 2.
Other variants of the gas clean up system have been tried including shell and tube condensers, rotary particle separation in a gas centrifbge and other various filters and demisters. These various unit operations were recommended and tested by Marick International however, success was extremely limited and is summarised in Table 1. Other difficulties occurred in the biomass feeding system, which was completely removed and simplified to allow for batch feeding for the end user application. Difficulties were regularly encountered in the gasifier with a heavily tar laden gas. This was due to poor feedstock specification and prolonged start-up at high turndowns in the gasifier leading to the production of significant quantities of tar. After CARE Ltd. was contracted to provide technical services, the specification of the feedstock was improved and a local source of willow and poplar was harvested in early 1999 by Shawton Engineering Ltd. This was then cut to the required size and dried typically to moisture content of 15-20% wt?h[dry basis]. Approximately 10 t of material was recovered for use. By improving the feedstock sue and moisture content, coupled with changes to the gasifier operation, substantial progress was made in producing a very low tar content gas as discussed in below. Engine start-up times were reduced fiom 90 minutes to 10 minutes fiom initial ignition of the gasifier to run solely on producer gas. Continuing developments and improvements were made to the gas cleaning system over the course of 1999 and substantial operational experience gained. Automatic char and ash removal systems were fitted and the improvements are summarisedbelow. OPERATION OF THE 75 kWe - IMPROVED SYSTEM FOR BBC
The resultant system for the 75 kWe system is illustrated in Figure 1 in terms of the basic unit operations used. Approximately 600 kg of prepared biomass is stored in the batch hopper on the gasifier. Air is drawn into the gasifier by means of a gas fan. During gasification, the grate is automatically riddled using a pressure sensor to control the pressure drop through the gasification zone. The riddled char and ash are removed by a slow speed auger and transferred to a storage bin. The hot gases exit the gasifier 443
md particulates are removed in the cyclone. Fine char and ash are collected in an 8 day storage bin on the bottom of the cyclone.
Table 1. Unit Operations tested by Shawton Engineering Ltd. Unit Oueration Overall Evaluation usedtested to early 1999 System not gas Semi-continuous - tight - and prone to leakage feeding system Poor collection of char and ash Char and Ash needs manual removal Removal
Result
I IRemoved.
Simple lid with latch mechanism fitted. Automated char and ash removal from gasifier to storage bin Char knock out pot Limited particle collection. High Removed fouling of packmg with tar Removed. Cyclone resized Cyclone oversized for the Cyclone and smaller geometry cyclone required duty fitted Receiver replaced with Poor arrangement to facilitate Char collection removable storage bin with easy removal of char receiver on cyclone auick release clamm Ineffective due to incorrect Removed Venturi scrubber assumptions about gas flow. I Wash tank not effective. Air IWasher column modified and Wash tank blower to cool recirculated water baffles put in tank. External air blown heat exchanger ineffective for duty required. installed Exit gas too warm [>80°C] Demister removed and Gas temperature too high and Demister on wash demister blocked with tars readily replaced with appropriately tank exit sized mesh for duty required Removed l Poor water condensation Air blast cooler Suitable for removal of residual Retained, but fitted with Cooling coil moisture additional air cooling Not suitable for producer gas and Removed Rotary particle its contaminants separator [RPS] Blocked with tar and particulates Removed Candle filters
I I
I I
I
The gases are then saturated in an internally baffled wash tank, followed by a pressure drop de-mister on the exit of the tank. The water in the tank is cooled externally by recirculation through an air-cooled exchanger. The cooled gases are then passed through a condensing coil to remove more water vapour and entrained droplets. The condensate is drained back to the wash tank through a dip leg. The gases then pass through a charcoal bed to remove trace low molecular weight organics and fine particulates. The remaining gases are then passed through the gas fan and into a gas buffer tank before passing through a two-stage final fabric filter before the engine [5 pm followed by 2 pm filter]. Various analyses have been conducted during 1999, which are presented below.
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WashColurm, Tank
Biom
Gas buEr tank
-&Ash Receiver Figure 1. Final gasifier configuration for the 75 kWe system
RESULTS AND ANALYSES During 1999, various analyses and characterisation of the products were made, in particular, analyses of the chars, wastewater and gases. One crucial assessment was the level of organics and particulates in the raw producer gas from the gasifier. The feedstock specification for the gasifier was a minimum diameter of 10 mm, typical length 80-1OOmm with minimal fines. The moisture content was in the range of 1520wt% [dry basis] during operation in 1999.
Char Analysis The char and ashes recovered from the process have been analysed for CHN [0 by difference] and ash and 8 samples analysed from char removed from the base of the gasifier. The results are given in Table 2. The char and ash recovered ffom the cyclone have not been analysed at this stage. The larger char particles are used in the charcoal filter [see Figure 11. The recovered char as expected is high in ash, reflecting a h g h degree of devolatilisation, typically greater than 90%, assuming an average ash content of 1.5 wt% for the willow. Table 2. Char analysis from willow C 6dev
82.2 1.8
H 1.0 0.1
0 0.92
--
N 0.5 0.1
Ash 15.3 2.38
Producer gas
A gas analysis average [four samples over 2 hours] for the gasifier is given in Table 3. The feedstock moisture content average was 17wt% [dry basis]. These results are 445
typical for an air-blown atmospheric pressure downdraft gasifier (see 1 for example). The batch gas samples were taken from the discharge side of the gas fan using a manual gas sample pump, collecting the samples in compressed gas cylinders. The gas samples were then analysed by gas chromatography by an independent organisation. Table 3. Average Dry Gas analysis for from Willow [vol%]
Methane [CH4] Carbon Dioxide [CO,] Ethylene [C2H4] Ethane [CzHa] H Y W w [H21 Propylene [C3H6] Propane [C3Hs] Carbon Monoxide [CO] n-Butane [C4HIO] Nitrogen + Argon [by difference]
1.9 14.0 0.3 0.1 17.6 0.1
Higher Heating Value [HHV dry gas] Lower Heating Value [LHV dry gas]
5.6 MJ/Nm3 5.2 MJ/Nm3
Wastewater All wet gas-cleaning systems generate a condensate, which is contaminated with organic and inorganic pollutants. The concentration of the pollutants in the condensate is always significant even for gasifiers that produce less than 250 mg/Nm3of tars. Typical wastewater compositions are given in Table 4 (2). The recycled water in the wash column, recirculated for 2 days is also given in Table 4.
Wood moisture Raw gas tar
Gasifier type Co-current 6 mmm3 200-600 % [drybasis]
Co-current 33 250-760
Biomass Eng. Ltd. 15-20 13.3
DOC dissolved organic carbon BOD Biological Oxygen Demand The concentration of pollutant strongly depends on the gasifier and the gas cleaning performance as well as on the amount of wastewater generated.
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Wastewater is produced in the process. In the BBC system, the wastewater from the process is recirculated over the by-product char. This simple but effective system was adopted from the work of Shell Recherche at Rouen and a similar approach has been adopted by other organisations working with small-scale gasifiers. The char will be replaced regularly, dried and returned to the gasifier to minimise wastes. Detailed wastewater and product analysis will be carried out when the unit has been recommissioned for BBC in Northern Ireland and to comply with Integrated Pollution Control [IPC] authorisation for the system. Organics and particulates in the Producer gas
As part of the monitoring program, operated by the DTI as part of the IEA Gasification Group Activity on the review of "tart' sampling protocols, the Coal Research Establishment [CRE] were commissioned to test the 75 kWe system when operating on willow and running the gas engine. The monitoring took place in December 1999 and followed the draft protocols for small-scale downdraft gasifiers (3, 4). The sampling method in summary is: 1. Filtration of the producer gas through a fibreglass filter element, 2. Passing of the gases through an ice bath to remove water, 3. Passing of the cooled gases through a cooled impinger system of 3 vessels [sutrounded by dry ice/acetone mixture], each containing dichloromethane to adsorb organic components, 4. Measurement of the pressure, temperature and flowrate of the gas.
The collection system is weighed before, after the gas sample is taken, and organics are determined by solvent evaporation and distillation. Solids are measured by solvent washing the filter and then drymg. Full details for the basis of the procedure are described by Abatzoglou et al. (3). Due to a sampling error, only data for the raw gas after the cyclone was considered representative and a second monitoring activity will take place when the gasifier is operational in Northern Ireland. The organics level in the gas after the cyclone was 11.1 mg/Nm3 and the particulates were 47 mg/Nm3. Organics associated with the char were -1.3 mg/Nm3, as shown in Table 5. The gas loading in terms of the chemicals analysed for by CRE are summarised in Table 5. As can be seen, the key chemical moieties are BTXs, naphthalene, dibenzofuran and fluorene, comprising 76 wt% of the identified chemicals. This would suggest that there is minimal bypassing of pyrolysis products through the char bed in the gasifier. The level of organics in the gases is very low and clearly meets the proposed levels for "tars" and particulates suggested by Stassen and Koele (12) and summarised by Biihler (5) of 100 mg/Nm3 for "tars" and 50 ms/Nm3 for particulates for engine applications. The range of components found in the gas are typical of those found in downdraft gasification systems (6). Analysis of the process gas after the cyclone indicates that there are no measurable amounts of phenols or methoxyphenols present suggesting minimal bypassing of pyrolysis products through the gasification zone.
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Table 5. Analysis of the raw producer gas and char after the cyclone [mg/Nm3]
Benzene Toluene Xylenes Naphthalene 2-Methylnaphthalene 1-Methylnaphthalene Biphenyl Acenaphthene Dibenzofuran Fluorene Phenanthrene Anthracene Fluoranthene Pyrene Benzo(a)anthracene Chrysene Benzo(b)fluoranthene Benzo(k)fluoranthene Benzo(e)pyrene Benzo(a)pyrene Indeno(1,2,3-cd)pyrene Dibenz(a,h)anthracene Benzo(g,h,i)perylene Total
Organics in Raw Gas 0.47 0.99 1.53 0.83 0.47 0.13 0.52 0.18 3.73 0.99 0.45 0.09 0.13 0.16 0.02 0.11 0.04 0.02 0.02 0.02 0.09 0.02 0.07 11.08 mg/Nm3
Organics in particulates <0.02 0.13 <0.02 <0.02 <0.02 <0.02 <0.02 <0.02 <0.02 <0.02 0.09 0.02 0.25 0.31 <0.02 <0.02 <0.02 <0.02 <0.02 <0.02 0.09 0.02 0.07 -1.3 mg/Nm3
IPC AUTHORISATION AND THE ENVIRONMENT In the UK, gasification and pyrolysis processes (7, 8) are classified as prescribed processes and must comply with IPC [Integrated Pollution Control] (9, lo), BATNEEC [Best Available Technology Not Entailing Excessive Cost] and BPEO [Best Practicable Environmental Option] (1 1). A significant lead-time must be allowed for an application to install a plant and ensure compliance with relevant legislation. The costs of an application and the annual renewal fee for the licence to operate need to be included in the detailed costings of a project, and may be a significant cost component over the life of the project. The IPC system has been superseded by IPPC [Integrated Pollution Prevention Control] from 1 August 2000, and is applicable within the EU.
ENGINE OPERATION The engine used for the gas testing was a Perluns Elmer Series 2000 engine, with a rating of 140 kWe on liquefied petroleum gas [LPG]. As producer gas has a calorific value [CV] considerably less than that of LPG [86.4 MJ/Nm3compared to producer gas CV of 5.2 MJ/Nm3], the engine will be derated. The differences in LHV only reveal part of the problem. The net CV of the aidfuel mixture is the crucial factor and for 448
LPG, the mixture CV is 3.1 MJ/Nm3for LPG/air and -2.4 MJ/Nm3 for producer gaslair [depending on the amount of excess air to minimise CO emissions], therefore the engine deration will be at least 20% without turbo charging or external gas compression. These values are applicable under ideal circumstances. In reality, the deration will be much higher, typically 3540% as excess air will be used with producer gas to reduce CO emissions and the fueVair mixture temperature will typically be above ambient in normal gas engines. The issues of engine deration have been recently summarised by Stassen and Koele (12) and highlighted that non-stoichiometric combustion of the producer gas, variation in inlet pressure and temperature all contribute to deration of the engine. For the engine used in this project, deration is estimated at 45%. The gas engine was inspected regularly during the testing programme and no build up of deposits were observed in the inlet manifold or fuel/air mixture system.
DISCUSSION The testing of a variety of unit operations to clean the producer gas has led to the development of a significant operational resource with regards to the next stage of refining and simplifjmg the gas cleaning system. With the improvements in gasifier operation leading to a very low "organics" gas, further developments to the gas cleaning system will look primarily at reducing the particulate loading in the raw gas and the use of a self-cleaning filter system to remove the particulates. This will then be followed by heat recovery from the gases and condensate removal. By using such a substantially streamlined gas cleanup system, the installed capital cost is expected to drop significantly. System design will allow for power or CHP options. The careful specification of the feedstock, as has been shown in numerous other downdraft systems is essential to generate a very low organics containing gas. To this end, this will remain an important issue for clean biomass fuels to allow the investment costs for the gas cleaning system to be reduced. The use of a carefully specified feedstock dramatically improved operation and allowed for rapid gasifier and engine start-up. One key issue, whch did cause some delays in operating the engine for prolonged periods, was the difficulty in obtaining on site technical support for the engine from the original supplier. The engine uses a sophisticated gadair monitoring system and t h ~ s has been prone to minor failures, which have required software intervention to rectify.
CONCLUSIONS Biomass Engineering Ltd. will build upon their experiences of the past 4 years to develop a low cost, simplified gasification + engine system. The initial development has led to the production of a high quality gas suitable for engine use. The initial investment costs have been high, but the wealth of experience from the past four years has led to substantial improvements in design and the reduction of investment costs. Difficulties have been encountered in getting the right level of technical support for the gas engine. Future work will use a Series 1 Jenbacher gas engine, capable of 146 kWe on producer gas.
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FUTURE WORK AND DEVELOPMENT
Based on the operational experience gained on testing a range of unit operations, the next systems to be constructed will have a variety of improvements, fixther reducing the number of unit operations and hence costs. As the operation of the gasifier allows the production of a very low organics content in the producer gas, various developments have been made to minimise the gas cleaning system and hence installed costs. The next unit, which will be an onsite demonstration and test unit of 150 kWe will use local residues to generate electricity for the use of Shawton Engineering Ltd. in their works. Char and ash removal will be automatic and the gas cleaning system will be selfcleaning with automated ash removal to a storage bin. Other projects for multiple units are under discussion. Further work will assess the optimal quantity of by-product char required to remove all traces of organic contaminants in the excess water generated in the process. This will be done in conjunction with Aston University. ACKNOWLEDGEMENTS
Biomass Engineering Ltd. would like to thank the Dept of Trade and Industry [DTI] through the Energy Technology Support Unit [ETSU] for providing assistance with the monitoring for organics and particulates in the producer gas, as tested by CRE. REFERENCES
1. Mukunda H.S., Dasappa S., Paul P.J., Rajan N.K.S., Shrinivasa U., Sridhar, G. and Sridhar H.V. (1997) Fixed bed gasification for electricity generation. In: Gasification and Pyrolysis - State of the art and firture prospects, (Ed. by M. Kaltschmitt and A.V. Bridgwater), pp. 105-116. CPL Scientific Press Ltd, Newbury. 2. Hasler P., Morf P., Buehler R. and Nussbaumer, T. (1998) Gas Cleaning and Waste Water Treatment for Small Scale Biomass Gasifiers. Swiss Federal Office of Energy and Swiss Federal Office for Education and Science, Zurich. 3. Abatzoglou N., Barker N., Hasler P. and Knoef H. (2000) The development of a draft protocol for the sampling and analysis of particulate and organic contaminants in the gas from small biomass gasifiers. Biomass & Bioenergy, 18, 1,5-17. 4. Knoef H.A.M. and Koele H.J. (2000) Survey of tar measurement protocols. Biomass & Bioenergy, 18, 1, 55-59. 5. Buhler R. (1997) Fixed bed gasification for electricity generation application in Europe. In: Gasification and Pyrolysis - State of the art and future prospects, (Ed. by M. Kaltschmitt and A.V. Bridgwater), pp. 117-128. CPL Scientific Press Ltd, Newbury. 6 . Jayamurthy M., Dasappa S., Paul P.J., Sridhar G., Sridhar H.V., Mukunda H.S., Rajan N.K.S., Brage C., Liliedahl T. and Sjostrom K. (1997) Tar characterisation in new generation agro-residue gasifiers - cyclone and downdraft open top twin air entry systems. In: GasiJcation and Pyrolysis - State of the art and future prospects, (Ed. by M. Kaltschmitt and A.V. Bridgwater), pp. 235-248. CPL Scientific Press Ltd, Newbury.
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7. The Environmental Protection Regulations (Prescribed Process and Substances) Regulations 1991, Statutory Instrument 1991/472, ISBN 0 11 013472 9, UK. 8. The Environmental Protection Regulations (Prescribed Process and Substances) Regulations 1991, Statutory Instrument 19911507, ISBN 0 11 013507 5, UK. 9. Processes subject to Integrated Pollution Control, Chief inspectors guidance Note Series 2 (S2), Fuel production and Combustion Sector. S2 1.07 Carbonisation and Associated Processes: Smokeless Fuel, Activated Carbon and Carbon Black Manufacture, London HMSO, September 1995, ISBN 0-1 1-753177-4, UK. 10. Processes subject to Integrated Pollution Control, Chief inspectors guidance Note Series 2 (S2), Fuel production and Combustion Sector. S2 1.08 Gasification Process: Gasification of Solid and :Liquid Feedstocks, London HMSO, November 1995, ISBN 0-11-753202-9,UK. 11. Environmental Protection Act, 1990, [reprintedwith corrections 19981, Chapter 43, The Stationery Office Ltd., ISBN 0-10-544390-5, UK. 12. Stassen H. and Koele H.J. The Use of LCV-gas from biomass gasifiers in internal combustion engines. In: Gasification and Pyrolysis - State of the art and future prospects, (Ed. by M. Kaltschrmtt and A.V. Bridgwater), pp. 269-281. CPL Scientific Press Ltd, Newbury.
45 1
Gas mixing in a pilot scale (500 kWth) air blown circulating fluidised bed biomass gasifier S.R.A. Kersten', R.H.W. Moonen', W. Prins*,W.P.M. van Swaaij* 'Netherlands Energy Research Foundation, Westerduinweg 3, PO Box 1, 1755 ZG Petten, The Netherlands *Facultyof Chemical Engineering, University of Twente, PO Box 217, 7500 AE Enschede, The Netherlands
ABSTRACT: To study the gas mixing capacity of circulating fluidised bed (CFB)biomass gasifiers, radial and axial gas concentration profiles have been measured and interpreted in a hot pilot scale biomass gasifier (100 kglhrfiel). The presented data of the pilot scale gasifier are unique and provide new insight in the radial gas mixing capacity of circulating fluidised bed gasifiers. Considerableradial gas phase concentration gradients have been observed in the pilot plant gasifier, with a difference between wall and centre concentrationsup to a factor 3. It must be concluded that the radial gas mixing is far from ideal. On basis of these pilot plant data and the standard dispersion model (to describe gas mixing) it can be concluded that the radial PCclet number of the dilute region is in the order of 1000. Such a value excludes the radial mixing of gases almost entirely INTRODUCTION In many western countries biomass is supposed to play an important role in the future energy production system. The Ministry of Economic Affairs of the Netherlands for instance has formulated the policy intention that by 2020, some 10% of the grand national energy generation (which amounts to 2800 PJ,hlyr) has to come from renewables, and 3.5% (100PJrh) in particular from biomass and sustainable waste This provides a major impetus for the application of biomass as a fuel for energy generation. Circulating fluidised bed biomass gasification is one of the possible conversion methods to allow the fuelling of biomass into modem energy conversion equipment. At present the most important performance indicators in biomass gasification technology are:
'.
(1) Thermal efficiency. Typical values are usually around 60 % at pilot scale. Pilot plant data and modelling indicate that the carbon conversion is the most important parameter in thermal efficiency optimisation for the ECN gasifier '. (2) Heating value of the product gas. In case of electricity production a lower limit is set by the gas turbine or gas engine.
452
(3) Impurities in the primary product gas. Tars, HCl, NH3, particulates and metals are important to decide on the type and size of down-stream equipment (-as clean-up section, gas turbine or engine). In case of power generation with a gas turbine (for instance in a combined cycle), the tar content of the gas offered to the turbine is subjected to strict constraints (<< 100 mglnm3) because of blocking problems in the burners and possible deposition on the rotor blades. However, tar contents in the product gas of pilot and demonstration plants exceed the maximum amount permitted, up to a factor 100 (ECN gasifier, Varnamo gasifier) 4*5*6. Decreasing the amount of tar in the product gas by adjusting the gasifier lay-out/operationwill lead to a reduction in capital and operational cost of the gas clean-up section. The efficiency of the gasification process in terms of conversion degree (carbon) and the final tar content in the product gas is strongly related to the flow structure and the degree of reactant mixing in the bed. For instance, during the devolatilisation process the tar yield is approximately 50 % on mass basis 7*8*9*10.A tar content in fhe product gas of 7.5 g/nm3 at typical gasification conditions (T=850 O C 1=0.30)(ECN gasifier), results in a typical tar conversion of 97 % 4*5. Table 1 illustrates that at such high degrees of conversion, the actual conversion is very sensitive to the flow structure and gas mixing state. A first order reaction rate is assumed to describe tar cracking as a first approximation. Table 1. Conversion for a first order reaction versus k.t Plug-flow
Plug-flow + axial dispersion P e p 10
Parabolic velocity profile peePep.0
Parabolic velocity profile Pe,3=300 Pq-
CSTR
0.632 0.950 0.999
0.603 0.914 0.992
0.557 0.887 0.990
0.598 0.899 0.99 1
0.500 0.750 0.875
kt [-I
1 3 7
Not only the tar cracking will be affected by the flow pattern and mixing, but also the carbon conversion is expected to be influenced by the local gas concentration differences. At CFB gasification conditions (T=850 O C - ;1=0.30- z=l [s]), combustion is the dominant mechanism for carbon conversion, because carbon gasification reactions are too slow at the prevailing conditions "*'*.Hence, good contact of oxygen and carbon is required. In this paper the gas mixing in a CFB biomass gasifier is discussed on basis of radial and axial gas concentration profiles measurement in the pilot plant of ECN (100 kglhr fuel). The standard dispersion model is applied to interpret the experimental results.
THEORY AND PREVIOUS WORK Radial gas mixing in turbulent single-phase gas flow is generally considered to be quite good. Reported radial PCclet numbers are in the order of magnitude of 200 I3*l4, which means that radial gas concentration gradients are smoothed out over relative short characteristic reactor length. The effect of addition of a solid phase on the radial gas mixing has not been described unequivocally in the literature. In Table 2 an overview is given of previous work on radial gas mixing in gas-solid risers. The contradictory findings regarding the effect of the presence of solids on radial gas mixing in a riser,
453
R P
76 130
1710
2600
Bader 24
Werther l 3
2600
Amos l4
Gay5n
[%I
****
,
[380-7101 [O - 1101
10-11
[l - 31
[0 - 701
5
[2 - 71
[25 - 1001
[98- 1771
****
****
[O - 1201
p = specific densit: of the used solid phas dp = average dim1 ter of the used solids Gs = applied solids flux & = solids volume fraction
1560 2450
2
250
2600
Adams 23
62 71
[20-1501 220
1500 1100
I
I
’
~
I Used model IFindings ****
-
w
I
1
$,
~~
D, increases with increasing solids volume fraction **** [40- 1801 radial disp. D, increases With velocitv in single phase flow (Pe constant). D, inversely proportional to Gs. D, inversely proportional to [83 - 2851 [26 - 651 radial disp. Gs. [200-3001 [20 - 601 radial disp.1 Single phase flow Pe, is core annulus constant. No influence of Gs on -~~D.. ‘Little effect of G S on D, T300 - 4001 [15 - 301 I radial disp. Per is proportional to Gs. Gs [150 - 2601 [30 - 401 I radial disp. / I not much different than single core hase flow. is proportional to Gs. [500-2000] [2 - 61 radial disp. 1 However Dris highest for single phase flow Pee = radial Pixlet number based on tl :velocity and reported values of D, [-I D, = reported values of the radial dispersion coefficient [m2/s]
I
Table 2. Literature references regarding radial gas mixing in risers Reference I Solids charact. (1) Solids charact. (2) Reported Pee & D,
can probably be attributed to differences in: i) equipment and operating conditions, ii) physical properties of the solids, iii) solids flux and solids hold-up and iv) the models used to interpret the data. However, concerning the influence of a solid phase on the radial distribution of the axial gas velocity, the reported findings are matching 1s*16917. The overall conclusion is that the radial gas velocity distribution shifts from flat (turbulent) to parabolic, when the solids fraction is increased. The work of Van Breugel et al. can be regarded as a pioneering paper in this field Is. On basis of results presented it is argued here that solids reduce the turbulent intensity of the gas phase. This hypothesis can be confirmed with a simple engineering model for developed two-phase flow 18. By applying this model to a dilute riser system, the velocity profiles indeed become sharper due to solids and it was found that the most important parameters are: solids fraction (see Figure 1 ) and solids diameter. Another important reported fact concerning the hydrodynamics is that the slip velocity can be much higher than the terminal velocit of the particles under consideration. This can be attributed to the presence of clusters If20 Following the supposition that the presence of solids reduces the gas turbulence and assuming that gas turbulence is the dominant mechanism for gas mixing, increasing the solids flux is expected to decrease the radial dispersion. Another mechanism that influences gas mixing might be the radial movement of solids. At high solids volume fractions the solids movements towards and from the core will entrain important amounts of gas phase23. This indicates that addition of solids can also increase radial gas mixing. In conclusion, depending on whether the damping of turbulence or radial movement of solids will be dominant, addition of solids will either decrease or increase the mixing intensity. Many authors pointed out that although the axial dispersion coefficient is much higher than the radial dispersion coefficient in a CFB, the back-mixing effect is still negligible because of the large axial convective flow 13*24*27,28. A widely used model to describe mixing in a given flow field is the standard dispersion model. It is represented by the following equation for axis-symmetrical cylinder co-ordinates:
.
rl
Or in dimensionless form by:
-- 1 Peg
)”ag
_ .
a2c+a[[ac -+---+-I a2c]ac PeE ag2 ae Pe, ae2 ae
- 1
1
Pe,
d2c)
ae2
This equation can be solved analytically for some limiting cases (see Table 3).
455
Negligible axial dispersion, constant radial Constant axial and radial dispersion coefficients dispersion coefficient
C
C Jo(an') -= 1+ 0 . 5 c
cm
n=l
e
(0.5-9, )BE'
Jo(AnI2
4,
In the models, described by the equations above, the complete dimensionless radius is used. Other researchers only use the standard dispersion model for the core region, which interacts with a hypothetical recycle or by-pass flow (annulus) 14. For such a model the solution becomes:
C
e-A2Eg'
in which A, are the roots of:
Now D includes an extra fit parameter (W) i.e. the by-pass or recycle flow rate. Expressions for the coefficients A, q, B, D and E are given in table 4.
$IT
Table 4. Expressions for the coefficients A,q,B,D and E I
A=- 1 Pee
I
~
I
4 = -+ 4 D ( f , W , pee B = Pee
D = -2a2v W E=--1 pee
For the analytical solutions the underlying boundary conditions are:
ac @ = 0,8= O)+ point source injection,O = 0 (14 > 0)and 8 = 1-+ ae = 0 f * = (ve) + c = o,g' = oo (ve) + c = c,
456
0
0.2
0.4
0.6
0.8
1
rm 1-1
Figure I . Radial gas velocity profiles Data from Van Breugel 15: "squares, Gs-390 kg/(m2.s)", "triangles, Gs=O kg/(m2.s)". Solid lines are computed with the engineering model for two-phase flow (Nieuwland et al. I*) .
Concerning the radial solids distribution, Matsen claimed that below 0.5 vol. % solids hold-up the radial solids profile is uniform 29. Other researchers found clear
radial distributions, although for risers operating in the turbulent regime the steepness of the radial distributions is less 3033'*32*33. Actual data of gas concentration profiles in a pilot riser reactor are reported by Jiang et al.34. They studied ozone decomposition in a catalytic circulating fluidised bed. By applying an UV detection technique they were able to measure strong radial and axial ozone concentration gradients in their experimental set-up. It was concluded that the ozone concentration profiles are consistent with the trend exhibited by the voidage distribution in both axial and radial direction.
EXPERIMENTAL Pilot scale gas$er All hot experiments are carried out in the pilot scale CFB gasifier (see Figure 2 ) at the Energy Research Foundation in Pemn, the Netherlands. The riser is six meters in length and has an inner diameter of 0.2 meter. Until now, accumulated operating experience amounts to approximately 500 hours of gasification runs. Normally a test lasts for eight hours. However, the longest non-stop test of 45 hours has proven that the facility can operate continuously without accumulation problems. Start-up is realised by means of trace heating and operating the gasifier in combustion mode. Air can be fed to the riser through the nozzles at the bottom (primary air) or alternatively at several other axial positions (secondary air). Biomass is fed just above the nozzles or at a height of 1.5 meters. The used sand particles have an average diameter of 500 pm. (400-630 pm) Both nitrogen and air can be used to fluidise the auxiliary bubbling bed. In table 5 ranges of gasification operating conditions are given. Table 6 gives some typical mixed cup compositions of the produced gas.
457
Standpipe
Figure 2. ECN pilot scale (100 kgihr fuel) CFB gasifier, length = 6 meter, internal diameter = 0.20 m, temperature = 700 - 900 O C , diameter sand = 500 pm.
45 8
Table 5 . Ran Biomass feed Air feed Temperature Gas velocity Sand flux
es of asification o ratin conditions [kg/hr] 60-100 [nm3/hr] 80-120 700-900
[k (m2.s)]estimated
7-30
The riser is well equipped with on line measuring systems for temperature (axial), pressure (axial) and concentration (axial & radial). After the first cyclone gas mixedcup concentrationsare measured continuously. Also at this position tar samples are taken.
co
[vol%dry] c02 [vol%dryl H2 [~01%dry] CH4 [VOI%drydryl H 2 0 [vol% wet]
test a
test b
14.2
12.3 17.9
test c 11.7 17.8
13.3 3.8
10.1 3.7
16.4 11.9 4.0 12.4
****
****
Method of measurement At five different axial positions, specially developed probes have been installed to withdraw gas samples from the interior of the reactor. The probes are designed to be freely movable over the reactor diameter, so full radial profiles can be obtained at each axial position. A ceramic porous filter, at the tip of the probe is filtering char, sand and ash from gas sample. After passing through the probe the gas passes through the following pre sampling train (see Figure 3): '
(1) Solids (chars), .heavy tars and water precipitation in a glass bottle filled with
ceramic wool. (2) An ice bath to condense the majority of the water. (3) A paper filter to remove residual heavy tars and aerosol char components. nitrogen offgas
I
riser
paper liller
>lOO°C
solids precipitation
water condensation
Analyzer
offgas
Figure 3. Pre-sampling train
More information considering the method of measurement and particular aspects of the analysis of these measurements can be found in 35.36,37
.
459
RESULTS Figure 4 gives a typical example of longitudinal gas fraction profiles in the pilot riser measured during a standard gasification test. The plotted profiles are measured in the centre of the riser, which means that changes in gas fractions can be caused by both radial gas mixing and gas production. However, Figure 4 suggests (also considering Figure 6) that most of the gas production is located below 1.5 meters, which is equivalent with 25% of the total riser length. This indicates that the majority of the biomass devolatilisationprocess takes place in the bottom part of the riser. CO, H2 and CH4 fractions, all components that are formed during biomass pyrolysis and tar cracking, clearly increase in the first 1.5 meters, while C02, a combustion product, shows only a slight tendency to decrease. Oxygen was never detected, in runs with only primary air, which indicates that the combustion zone is located below 1 meter. The small increase of CO, H2 and CH4 after 2 meter is presumably caused by tar cracking. Measured radial gas hction profiles are plotted in Figure 5. It should be noticed that the reproducibility is quite good. Profiles of CO, C 0 2 and CH4 show a parabolic behaviour, while fractions near the wall can be up to 3 times higher than fractions in the centre. The profiles are not completely symmetrical; this can be ascribed partly to the method of measurement and the gas-solid flow 24*36. One could think of several hypothetical causes for the radial profiles: ( I ) A heterogeneous radial biomass distribution. More biomass near the wall will induce more gas production near the wall, which results in higher product gas component fractions. (2) A parabolic gas velocity profile. As a result of the velocity profile a residence time distribution arises i.e. longer residence time near the wall (less dilution with inert gas) results in higher gas fractions. Parabolic gas velocity rofiles are observed by a number of authors in comparable reactor configurations 15.16,1? . (3) Heterogeneous tar cracking. If the tar cracking rate is proportional to the solids hold-up (char) more cracking occurs near the wall.
In a future paper these hypothesises with regard to the radial gas fraction profiles will be checked and validated by a reactor model. Information concerning radial mixing in the riser is obtained from Figure 6, in which radial H2 profiles are given at three different axial positions. If it is assumed that only little gas production takes place after 2 meters of reactor length, radial mixing must be poor. On basis of the measured radial gas concentration profiles and the standard dispersion model (to describe gas mixing) it can be concluded that the radial PCclet number of the dilute region is in the order of 1000. Experiments showed that the C02 radial profile is almost flat during gasification runs in which air is supplied only through the bottom nozzles. However, in runs with secondary air, and during combustion experiments (see Figure 7) also radial CO2 profiles (differences) have been measured. This would indicate that the existence and the shape of the radial C02 profile is somehow related to the oxygen I fuel ratio. During normal gasification operating conditions, all combustibles and O2 are converted to C 0 2 and H20 after a short equivalent of reactor length, resulting in nearly flat radial concentration profiles for these components. This is supported by gas concentration measurements 0.8 meters above the primary air nozzles, where hardly any 0 2 is detected. Secondary air is injected near the wall, so that a local excess of oxygen in this region can cause deviations from the flat C02 profile. Measurements show that the C02 profile caused by secondary air injection is less pronounced than profiles of the other components.
460
.
0
,
0
.
,
1
.
,
.
,
3
2
.
,
.
1
5
4
8
axial co-ordinate [ml
Figure 4. Measured axial volume fraction profiles for COz, CO, Hz and CH4,fractions at the centre of the reactor are plotted (h-0.275).
.
20-
5
m
-
18
la:
g 14: 12-
f
u
.
E
*"
. I t
'
/
10-
c
:
.
g
8-
2a
-5 s
6:
.
4 '&
4:
c
2-
0
CH, ,
.
,
.
,
.
.
I
,
.
I
Figure 5. Measured radial volume fraction profiles for COz, CO, Hz and CH4 at a reactor length of 2.75 meters (h-0.275).
-
eg . 12-
10-
F
U
.c C .-
-8
. 8-
8-
c
a
5e
. 4-
7 !
0.0
.
, 0.2
.
, 0.4
.
, 0.6
.
,
0.8
.
,
1.0
d/D (-1
Figure 6. Measured radial volume fraction profiles for Hz at a reactor length of 2.75, 4.00 and 4.75 meters (h=0.275).
46 1
*18O I =>.,
'.+ ........
......................... ..........v.-....... 0, centre T ...... CO wall
.' ~
........
".
.............. ..
p"'
b
o2wall
Figure 7.Measured axial volume fraction profiles of C02 and 0 2 in the centre and near the wall of the riser during combustion experiments (1-3). CONCLUSIONS To study the gas mixing capacity of circulating fluidised bed (CFB) biomass gasifiers, radial and axial gas concentrationprofiles have been measured and interpreted in both a hot pilot scale biomass gasifier (100 kg/hr fuel) and a cold-flow set-up. The presented pilot plant data are unique in their sort and provide new insight in radial gas mixing in a CFB riser. In the pilot plant, strong radial gas concentration profiles, up to a factor 3 higher concentrationsnear the wall compared to the centre, have been measured. These profiles are obtained during gasification conditions (850 "C,and A9.30). However, measured CO,profiles are found to be flat under certain process conditions. In case of only primary air, the COz profiles are flat. On the other hand secondary air and combustion process conditions yield important radial C02 gradients. This indicates that the shape of the radial C02 profile is related to the oxygen to combustiblesratio. On basis of the pilot plant data and the standard dispersion model it can be concluded that the radial Ptclet number must be in the order of 1000 in the dilute zone of the riser. Such a value excludes the radial mixing of gases almost entirely.
SYMBOLS & DEFINITIONS
-
T temperature k = reaction rate constant, first order D = dispersion coefficient
["CI [Us] [m2/s]
C gas/tracer [mole/m3] concentration Gs = solids flux [kg/(m:.s)l Gf fluid flux [kg/(m .s)l Jo,JI Bessel function , first kind, order zero [-I (one)
--
z = average gas residence time in the riser g = -z
L
5
*
=z
[s]
[-I
Z
r 8=R L Q=R
V(E98)L peE = DE(E,O)
462
[-I
[-I [-I [-I
R = riser radius [ml D riser diameter [ml r/R = dimensionless radius [-I d/D dimensionless diameter [-I Aq,B,D,E = parameters in equations (see table 2 & 3) [-I V,U = gas velocity [ds]
-
-
subscripts r,O = radial direction z,E = axial direction
Pe'E = V(E98)L aD&m
[-I
x
Pe - V(E98)R i3 - D,<&e> V(E96)R Pe'e = aDi3cE9e)
[-I [-I
ae I=
oxygen supplied oxygen needed for stoichiomehiccombustion
h = root of Bessel function
[-]
r-1
REFERENCES 1. Dutch Ministry of Economic Affairs, derde energie nota, actie programma duurzame energie in opmars 1997-2000, The Hague, (1997) 2. Kersten S.R.A, Prins W, van Swaaij W.P.M, 'quasi'-equilibrium models for biomass gasification, to be published 3. Bridgwater A.V. Chemistry for the Energy Future, IUPAC Chemistry for the 21st century (1999) ,editors: Pamon V.N, Tributsch H, Bridgwater A.V. Hall D.0 4. Drift v/d A, Kersten S.R.A, Biomassa-conversie-eigenschappen bij vergassing, (1998), ECN-publication 5. Drift v/d A , Brandstoffen uit reststromen voor circulerend wervelbedvergassing (1998), NOVEM publication 355 197/4040 6. St&I K, Neergaard M, Nieminen J, Progress Report: Varnamo Biomass Gasification Plant, 1999 Gasification Technologies Conference San Francisco 7. Thumer F, Mann U, Kinetic investigation of Wood Pyrolysis, Ind. Eng. Chem. Res., 20(3), 482, (1981) 8. Font R, Marcilla A, V e r d ~E, Devesa J, Kinetics of the Pyrolysis of Almond Shells and Almond Shells Impregnated with CoClz in a Fluidised Bed Reactor and in a Pyroprobe 100, Ind. Eng. Chem. Res., 29(9), 1846(1990) 9. Chan W.R, Kelbon M, Krieger B.B, Modelling and experimental verification of physical and chemical processes during pyrolysis of a large biomass particle, Fuel, 64,1505(1985) 10. Wagenaar B, The rotating cone-reactor (1994), Ph.D. thesis University of Twente 11. Bliek A, Mathematical modelling of a cocurrent fixed bed coal gasifier (1984), Ph.D. thesis University of Twente 12. Johnson J.L, Fundamentals of coal gasification, in: chemistry of coal utilization (1981), editor: Elliott M.A 13. Werther J, Hartge E.U, Kruse M, Novak W, Radial mixing of gas in the core zone of a pilot scale CFB, in: CFB 111, editors: Basu P, Horio M, Hasatani, 593 (1990) 14. Amos G, Rhodes M.J, Mineo H,Gas mixing in gas-solids risers, Chem. Eng. Sci., 48(5), 943( 1992) 15. van Breugel J.W, Stein J.J, de Vries R.J, Isokinetic sampling in a dense gas-solids stream, Proc. Inst. Mech. Emgrs., (1969)
463
16. Yang Y.L, Jin Y, Yu Z.Q, Wang Z.W, Investigation on the slip velocity distributions in the riser of dilute circulating fluidised bed, Powder Technology, 75,68 (1 992) 17. Yang Y.L, Gautam M, Mei J.S, Gas velocity distribution in a circulating fluidised bed riser, Powder Technology, 78,221 (1994) 18. Nieuwland J.J, Delnoij E, Kuipers J.A.M, van Swaaij W.P.M, An engineering model for dilute riser flow, Powder Technology, 90, 115 (1997) 19. Grace, J.R, Tuot J, A theory for cluster formation in vertically conveyed suspensions of intermediate density, Trans. Inst. Chem., 57,49 (1979) 20. Venderbosch R.H, The role of clusters in gas-solids reactors (1998), Ph.D. thesis University of Twente 21. van Zoomen D, Proc. Conf. interaction between fluids and particles, Inst. of Chemical Engineers, 66 (1962) 22. Yang G, Huang Z, zhao L, Radial gas dispersion in a fast fluidised bed, in: Fluidisation IV, editiors: Kunii D, Toei R, 145 (1984) 23. Adams C.K, Gas mixing in fast fluidised beds, in: CFB Technology 11, editors: Basu P, Large J.F, 299 (1988) 24. Bader R, Findlay J, Knowlton, T.M, Gadsolids flow patterns in a 30.5 -cmdiameter circulating fluidised bed, in: CFB Technology 11, editors: Basu P, Large J.F, 123 (1988) 25. Martin M.P, Turlier P, Bernard J.R. Gas and solid behaviour in cracking circulating fluidised beds, Powder Technology, 70,249 (1992) 26. G a yh P, Giego L.F, Adhez, Radial gas mixing in a fast fluidised bed, Powder Technology, 94, 163 (1997) 27. Dry R.J, Radial concentration profiles in a fast fluidised bed, Powder Technology, 49,37 (1986) 28. Brereton C.M.H, Grace J.R, Yu J, Axial gas mixing in a circulating fluidised bed, in: CFB Technology 11, editors: Basu P, Large J.F, 307 (1 988) 29. Matsen J.M, Some characteristics of large solids circulation systmes, in: Fluidisation Technology, editor: Kreans D.L, 2, 135(1976) 30. Zhang W, Tung Y, Johnson F, Radial voidage profiles in fast fluidised beds of different diameters, Chem. Eng. Sci., 44,3045 (1991) 31. Godfoy L, Patience G.S, Chaouki J, Radial hydrodynamics in risers, Ind. Eng. Chem, Res, 38,81 (1999) 32. Patience G.S, Chaouki J, Solids hydrodynamics in the fully developed region of CFB risers, in: Fluidisation VIII, 33 (1995) 33. Abed R, The characterisation of turbulent fluid bed hydrodynamics, in: Fluidisation IV, editiors: Kunii D, Toei R, 137 (1984) 34. Jiang P, Inokuchi K,Ozone decomposition in a catalytic circulating fluidised bed, in: CFB Technology III, editors: Basu P, Horio M, Hasatani M, 557 (1990) 35. Moonen R.H.W, Radial and axial gas concentration profiles in a CFB biomass gasifier - measurementsand analysis (1 999), MSc. thesis University of Twente 36. Jansen M, Cold-flow circulating fluidised bed hydrodynamics (1 999), MSc. thesis University of Twente 37. Oosting T.P, Improving circulating fluidised bed operation - experiments and modelling (2000), MSc. thesis University of Twente
464
A Demonstration Project for Biomass Gasification and Power Generation in China C. Wu, X. Yin, S. Zheng, H. Huang, Y. Chen Guangzhou Institute of Energy Conversion, Chinese Academy of Sciences, 81 Xian Lie Road Central, Guangzhou 51 0070 China
ABSTRACT : Based on the experiences gained from an experimental biomass gasification-power generation system, a new middle-scale demonstration plant has recently been constructed in Sanya, Hainan Island of China. The plant consists of a circulating fluidized-bed gasifier0 gas-purifying systems, gas engine/generator sets, and waste-water treatment facility. The designed electric power output of t h s plant is 1.2 MW; total investment is 5.9 million RMB. Preliminary performance data of h s plant have been obtained. It is expected that considerable economic and environmental benefit will be achieved in this project.
INTRODUCTION DEVELOPMENT HISTORY The technology of biomass gasification and power generation has already been employed in China for some time. The conventional technologies utilize the rice hulls gasification and generation for the rice mill, and the scale of the moving-bed gasifier system is small (the largest capacity is only 200 kW),which leads to lower economic benefit and some pollution. Guangzhou Institute of Energy Conversion has designed a circulating fluidized-bed (CFB) gasifier for wood powder in 1992 reported previously [1,2]. Since then, a number of CFB gasifiers have been constructed in various locations in Chna [3];the gas produced is most ofien used for boiler firing in these cases. In 1998, an experimental biomass gasification and power generation plant was built in Putian, Fujian province of China. This plant employed a middle-scale circulating fluidized-bed gasifier processing rice hulls from a rice mill; and the producer gas is used to power internal combustion engine/generator sets for electricity production. This experimental plant was started up successfully and run in nonnal operation; the accomplishment of this project demonstrated the technological feasibility of such system, and provides technical support for larger power plant projects [4].
465
CONSTRUCTION CONDITIONS Sanya Timber Factory is a large-scale timber processing enterprise that locates in Sanya, Hainan Island of China. In this factory there are two product lines imported from Sweden for 30,000m3 shaving plank and 40,000m3mid-density fibre, and other workshops. About 100 tons wood waste is produced every day, which includes 35 tons wood powder, 10 tons leftover material, 10 tons sieving waste and 45 tons all lunds of bark. Due to the complicated sorts of timber waste, some of the wastes are used and others are disposed which causes severe problems of storage and pollution. On the other hand, timber processing is a high electricity consumption activity; this factory needs power about 5000 kW and the cost is up to ten million RMB per year. So to convert the wood waste into electricity for the factory’s self-consumption through biomass gasification-power generation technology has favourable conditions of environmental protection and energy conservation. Thus we selected this factory as the site for the new MW-scale biomass gasification and power generation plant.
DESIGN SCHEME
MAIN EQUIPMENTAND SYSTEM CONFIGURATION The characteristics of the CFB technology and relevant design and test methods have been discussed previously, e.g. [l-31. In the present project, the main equipment of the biomass gasification-power generation system includes a biomass conveyer and feeder, a CFB gasifier, a Venturi purifying tube, water scrubbers, a blower, gas engine/generator sets (made by Hongyan Machine Works, China), a waste-water treatment pool, etc. The flow diagram is shown in Fig. 1. The layout of the plant was decided according to practical conditions and is displayed as Fig. 2.
POWER PRODUCTION AND ELECTRICITY DISTRIBUTION
In order to get the utmost out of the generation facility and with consideration of practical situation of the power load distribution, the produced electricity will be distributed in the following manner: (1) Dividing the electricity output of the biomass power plant into two parts, each part is provided by three parallel gas engines and the total output of each part is 600 kW. (2) The power load of the timber factory which can be substituted is separated fiom the total load and equipped with bi-directional switch; it can make use of not only the electricity produced by the biomass power plant but also the outside power grid. (3) Dividing the electricity load of the timber factory that can be substituted into three parts, the capacity of each part is 500-600 kW, so the biomass power plant can keep higher operation ratio.
466
I
Fig. 1 Flow diagram of biomass gasification and power generation plant. (1 .Gasifier. 2. Primary cyclone. 3. Secondary cyclone. 4. Venturi scrubber. 5. Water scrubber. 6. Water scrubber. 7. Gas tank.)
TECHNICAL SPECIFICATIONS According to experiences gained from the experimental biomass gasification-power generation plant in Putian and the electricity consumption characteristics of the timber mill, this plant can reach the operation objectives as follows: (1) The total capacity is 1200 kW; the average load for long-term operation should reach 85% of the designed capacity. (2) The system failure rate is less than 10%; the operation shutdown rate is less than 10%. (3) The total utilization rate is about 70%. (4) The consumption of raw material is 1.2 kg/kW.h(dry); it is totally 30 tons each day.
467
Fig. 2 Layout of biomass gasification and power generation plant. (1. Feed material storage. 2. Control room. 3. Gasifier and gas cleaning equipment. 4. Gas enginelgenerator set. 5 . Electricity distribution room. 6 . Waste-water treatment facility.) LAND AND EMPLOYEE REQUIREMENT
The total area of this plant is about 1500 m2: Gasifier workshop: Gas engines workshop: Biomass storage field: Waste-water pool:
300 m2 300 m2 600 m2 300 m2
The total number of staff, including 3 administrators, is about 25 persons. They are divided into three groups to work in turn. Among each group, 2 persons are in charge of the gasifier, 4 persons are in charge of electricity production, and 1 person distributes electricity.
MODE OF ADMINISTRATION
In order to make sure the high efficiency operation of the power plant and get benefits, the plant should be administrated by one special organization or company which takes separate accounts and is responsible for its own profits or losses. So Guangzhou Institute of Energy Conversion cooperated with Sanya Timber Factory to invest together, and established a joint-stock company, which is a legally independent body;
468
and the rights and interests of both sides are decided by its contribution to the biomass power plant. It has been agreed upon that the Institute shares 62.8% stock and the Factory 37.2% stock. The joint-stock company is fully responsible for the financing, engineering construction, management and refunding of debt. The cooperative relationshps among the investors are regulated by a contract, and the operation should strictly follow the relevant regulations.
BUDGET
The total investment for this project is 5.9 million RMB including factory construction, equipment and other relevant costs; the detailed budget is shown in Table 1. Table I Budget for the 1200kW Biomass Power Plant.
Item 1. Building construction 1) Main workshop 2) Civil work 3) Foundation 4) Waste-water pool 5) Biomass storage field 2. Equipment manufacture 1) Gas engine 2) CFB Gasifier 3) Purifying system 4) Waste-water treatment 5) Water cooling tower 6) Conveyer 7) Electricity distribution board 8) Electric appliance and cable 3. Equipment installation 1) Gasification system 2) Gas engine 3) Power system 4. Technology and regulation 1) Gasification regulation 2) Generator regulation 3) Power distribution regulation 4) Patent cost 5. Management cost Total
Quantity 300 m2 300 m2 300 m3 600 m2 6 sets 1 set 1 set 1 set 2 sets 1 set 1 set
S~~(IO~RMB) 70 25 10 10 10 15 350 165 55 20 35 5 5 15
50 20
5 10 5 100 5 10 5 80 50 590
PRELIMINARY RESULTS
Construction of the Sanya biomass gasification-power generation plant was 469
completed in late 1999; and the plant was started up in early 2000. During the test period, the CFB gasifier and gas cleaning system are operated normally; and the gas engine/generator system is operated at about 60% of the rated load, providing part of the electricity demand of the Sanya Timber Factory. In Table 2, some operation and gas quality data are given. The cold gas eficiency of the gasifier is about 70-80%. According to manufacturers specifications, the efficiency of the gas enginelgenerator system is 23.5%. So the wood fuel to electricity efficiency of this plant is estimated to be 17-19%.
Table 2 Gasifier Operation and Gas Quality Test Data. Feed rate of wood (kgh) Feed rate of air (Nm3/h) Gasifier temperature Gas composition (vol.%) H2
co co2
CH, C2H6 C2H2 N2 Lower heating value of gas ( K J / N ~ ~ )
1500 1650 775
885 1350 800
885 1610 940
7.59 24.83 13 5.91 0.3 0.27 48
6.33 18.66 11.57 5.58 0.11 0.27 51.8
7.04 18.29 13.76 3.63 0.6 0.55 56
6360
5300
5010
COST-BENEFITANALYSIS
COST The electricity cost from the biomass gasification-power generation plant includes: Raw material cost: This part of cost includes the raw material purchase, collection and transport. According to the contract, the raw material is provided by the timber mill and the cost of biomass wastes is about 0.06YkW.h. Workers’ wages: Thls part of cost includes the wages for the administrators, operation workers and other labors. Equipment maintenance expense: This part of cost includes the purchase of spare parts and damageable material, and maintenance fee. Plant administration expense: this part of cost includes the running expenses for daily official business. The detailed analysis is listed in Table 3:
470
Table 3 Operation Cost. Item 1.Workers’ wages 1)Administrator 2)Operation workers 3)Other labors 2. Maintenance cost 1)Gasificationand purge 2)Generator maintenance 3)Generator fittings 4)Lubricating oils 5)Others 3.Administration cost 1)Officialbusiness expenses 2)Reception fee 4.Raw material cost Total
Unit price 20000Ylyear 12000Ylyear 10000YIyear 5OOYIday 5OOYIday
6000Ylmonth 3OOOYlmonth
Sum( 104y/year) 37.2 6.0 25.2 6.0 28.5 2.0 1.5 10.0 5.0 10.0 10.8 7.2 3.6 40.0 116.5
OUTPUTAND GROSS PROFIT Decided mainly by the power load characteristics of timber processing factory, the biomass power plant is required to operate except the overhaul of production line once a year. So the designed power output is up to 6.65 million kw.h/year. The current electricity price in Sanya is 0.55 Y/kw.h, so the production value of the power plant is about 3.48 million Y/year and the gross profit is 2.32 million Ylyear. The profit forecasting is shown in Table 4. Table 4 Profit forecasting (1200 kW). Item 1.Operation days 2.0peration time 3.Self-comsuptionload 4.Power output 5.Self-comsuptionpower 6.Power on sale 7.Electricity price 8.Production value 9.Gross profit
Unit Quantity Dayslyear 330 Hoursfday 24 kw 40.0 104kw.h/year 665.3 lo4kw.h/year 3 1.6 lo4 kw.h/year 633.7 0.55 Yh.h 104Y/year 348.5 lo4Y/year 232.0
Method of calculation
24~330~1200~0.7 24x330~40
633.7x0.55 348.5-116.5
DISCUSSION Our tests on the Sanya biomass gasification-power generation plant so far have yielded positive results. The CFB technology is suitable for gasification of powdered wood material, and proves reliable in continuous operation under different load conditions. After passing the gas cleaning system, the producer gas can be used in the 47 1
Hongyan type gas engine to produce electricity. Some operational problems like waste-water pollution have to be monitored in long-term tests. Base on the current price of biomass wastes in China, the electricity production cost in this project is around 0.2 Y RMBkW.h . The economic return from this demonstration project will reach about 2.4 million Y RMB/year, and all capital investment may be paid back in about 3 years. When the demonstration plant operates under designed conditions, about 10 thousands tons of biomass wastes will be utilized every year. Consequentially, with the advantage of no net C02 release, more than 10 thousands tons of C02 can be avoided. Comparing with coal combustion, a great deal of flue gas, such as NOx and SO2 can be alleviated. In the long term, the environmental benefit from the application of biomass gasification and generation technology should be much more obvious.
SUMMARY
An MW-scale biomass gasification and power generation plant has been constructed in Sanya, Hainan Island of Chma. During preliminary tests, it was learned that the operation of the CFB gasifier and gas cleaning system is satisfactory; the gas is of adequate quality and can be combusted in the Hongyan type gas engine for electricity production. The electricity generated is currently used to supply part of the electricity need of the Sanya Timber Factory. In more long-term tests, the technical feasibility and economic and environmental benefit of the biomass gasification and power generation technology may be demonstrated in this project. LITERATURE C. Wu, B. Xu, Z. Luo and X. Zhou, Performance analysis of a biomass circulating fluidized bed gasifier. Biomass and Bioenergy, Vo1.3, No.2, pp.105-110, 1992. B. Xu, C. Wu,Z. Luo, H. Huang and X. Zhou, Design and operation of a circulating fluidized-bed gasifier for wood powders. in "Advances in Thennochemical Biomass Conversion" (ed. A.V. Bridgwater), pp. 365-376, Blackie Professional Publishers, Glasgow, UK. 1994. C. Wu, P. Liu, Z. Luo, B. Xu and Y. Chen, The scale-up of biomass circulating fluidized bed gasifier. The 6th China-Japan Symposium on Fluidization, Oct. 9-11, 1997, Beijing, China, pp.196-200. C. Wu, S. Zheng, Z. Luo and X. Yin, The status and future of biomass gasification and power generation system. China-EU Renewable Energy Technology Conference, Mar. 1999, Brussels.
472
Pressurised gasification of biomass and fossil fuels in fluidised bed gasifiers, hot gas cleanup using ceramic filters and pressurised product gas combustion Wiebren de Jong, Omer Unal, Peter Hoppesteyn, Jans Andries and Klaus R. G. Hein Section Thermal Power Engineering Department of Mechanical Engineering and Marine Technology Deljl University of Technology Mekelweg 2, NL-2628 C D D e p , The Netherlands Phone: +31 15 2786751, FAX: +31 15 2782460, e-mail:
[email protected]
ABSTRACT Gasification of biomass and older fossil fuels, hot gas cleanup using a ceramic filter and combustion of LCV product gas in a combustor is performed using a 1.5 MWh test rig at Delft University and a 10-50kWth at Stuttgart University (IVD) in the framework of pilot plant research on efficient, environmentally acceptable large scale power generation systems based on fluidised bed gasification technology. The influence of operating conditions (pressure, temperature, stoichiometric ratio) on gasification, gas clean up efficiency and combustion characteristics (gas composition, conversion grades) are studied. The gasifiers are operated at pressures in a range of 1.5 - 10 bar and maximum temperatures of circa 900 "C. The Delft gasifier has a 2 m high bed zone (diameter of 0.4m) followed by a freeboard approximately 4 m high (diameter of 0.5 m). The IVD gasifier has a diameter of 0.1 m and has a total reactor length of 4 m. Both gasifiers are equipped with a hot gas cleanup ceramic filter and a pressurised combustor. Measurements are used for validating a mathematical model for emission of environmentally harmful components like fuel-nitrogen derived species. Results obtained are presented and analysed. A practically solids free fuel gas of a heating value acceptable for combustion in a gas turbine combustor is obtained after hot gas cleanup and at gasifier stoichiometric ratio's in the range of of 0.3 - 0.7. Carbon conversions are well above 80%. Fuel-nitrogen conversion is in-line with other pressurised investigations with bottom feeding of biomass.
473
INTRODUCTION Power and heat production by the process of pressurised air blown fluidised bed gasification of solid fuels is a promising technology with respect to emissions and efficiency. These process characteristics depend strongly on parameters like fuel type, pressure and air stoichiometry. There is a relation between these variables and the actual process configuration, like e.g. staged combustion, advanced PFBC IGCC, or the ‘British Coal Cycle’. In the framework of co-operation between the partner institutes IVD (University of Stuttgart) and the section Thermal Power Engineering (Technical University Delft, TUD) TUD has carried out experimental work at IVD. These experiments have been performed with the advanced staged combustion (DWSA) installation. The maximum thermal capacity of this test rig is ca. 50 kW. Also measurements have been performed using the Delft pressurised fluidised bed (PFBG) test rig with a maximum thermal capactity of ca. 1.5 M W . As fuels pelletised Miscanthus Giganteus (PFBG experiments) and German Brown Coal (Hambach open mine) and pinewood (DWSA test rig) have been chosen. Brown Coal forms a significant part of the electricity generation capacity in Germany (ca. 28%) and also in some Eastern European Countries, as well as e.g. Australia. Biomass has been chosen as it is becoming increasingly relevant for e.g. the Dutch energy supply situation. The Dutch government has decided that by the year 2020 10% of the energy supply should be based on renewables of which the main part will be biomass, see e.g. Kwant & Leenders [I]. With energy supplied from biomass as a renewable source there is practically no net COz emission, as the C02 released to the atmosphere will be taken up by plants in a relatively short time scale. By substitution of coal or other solid fossil fuels in general by biomass, net CO1 emissions will be reduced significantly. In the experiments pinewood was chosen as biomass model component. Pinewood contains practically no sulphur, has a very low nitrogen content and also contains practically no alkali nor heavy metals and is scientifically interesting for so called zeroreference experiments. The research work described in this paper was primarily directed toward emission studies and the main goals were: (1) to investigate the conversion of the main components in the fuel to LCV gas species and (2) to study the fuel nitrogen conversion to NH3 and HCN during pressurised fluidised bed gasification (PFBC) coupled to NOx formation in the combustion.
Of course, also other aspects of pressurised fluidised bed gasification were studied, like tar composition in the generated low calorific value (LCV) gas and carbon conversion. The ignition of the LCV gas during pressurised combustion of the LCV gas was also of interest. Besides the fuel, the main process parameters were the applied air stoichiometry and operating pressure.
TEST FACILITIES AND MAIN EXPERIMENTAL VARIABLES Figure 1 shows the pressurised fluidised bed test installation (DWSA) at IVD, University of Stuttgart. The DWSA is used for several years in the framework of research in the field of conversion of solid fossil fuels under well-defined reproductive process conditions and on a small scale (50 kWth maximal), see e.g. Nagel et al. [2].
474
"....
I
I
Fig. 1 Schematic of the DWSA test rig. Figure 2 presents the PFBG test installation at Delft university. ~.
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.1
Gas Turbine Corn bustor
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.
"
",
Gasifier
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@$ Fuel Air 4 t Steam
% S.A. Ceramic Filter
1
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Vessel gas -I
I
G.A.: Gar Analyrls S.A.: Solld A n a l y i l r
Fig. 2 Schematic of the PFBG test rig. Table 1 shows the range within which DWSA and PFBG process variables can be varied.
475
Table I Operating range (gasification) of the DWSA and PFBG test rig ~~~
Variable
Range DWSA
Range PFBG
Pressure (MPa) Temperature ("C) Primary air Stoichiometry, h (-) Fluidisation velocity ( d s ) Fuel
0.12 - 1.6 750 - 1000 0.3- 1.O 0.1 - 1.0 Coal, Brown Coal, Biomass
0.3 - 0.8 750- 1000 0.3-1.0 0.5 - 0.8 Coal, Brown Coal, Biomass
In table 2 the main dimensions of both gasifiers are shown.
Table 2 Main dimensions of the DWSA and PFGB gasifier DWSA Bed diameter (m) 0.10 Max. bed height (m) 1.0 Freeboard diameter (m) 0.177 Freeboard height (m) 3.0
PFBG 0.38 2.0 0.485 4.5
The DWSA installation can be divided into two main parts. The first part consists of an air preheater, fluidised bed reactor, solid fuel dosing vessel with on-line mass determination system and a hot gas cleaning section, containing a cyclone and a ceramic candle filter (Schumacher type). In the fluidised bed reactor the solid fuel is gasified with air to produce a low calorific value (LCV) gas that is cleaned of fly ash and unreacted solid ciirbonaceous material. Air and also additional nitrogen can be preheated and is introduced into the reactor by four nozzles just above the distributor plate. The reactor is electrically heated in order to maintain a constant temperature over bed as well as freeboard section. The solid fuel is fed into the bed section in the bottom part just above the distributor by a screw feeder from beside. The hot gas cleaning section ensures a good gas-solid separation efficiency, with filter temperatures of about 500 "C. The second part of the test rig consists of a combustion air preheater, a specially designed LCV gas burner, a flue gas cooler and a pressure control valve. The LCV gas combustor is situated in a water cooled pressure vessel. This swirl-diffusion combustor has a centric position in the ceramic combustion chamber. From the central symmetry axis of the burner to the outer wall the gas flows are in annular spaces whereby in the inner annulus primary air is added with swirl to the LCV gas and in the outer annulus secondary air is added. The analysis of the produced LCV gas is performed directly behind the ceramic candle filter by means of continuous on-line O2 (paramagnetic), CO- and COz (NDIR) analyses In addition, H2, CO, CH4 and N2 concentrations are measured off-line by means of a gas chromatograph. An FTIR is used for measurement of NH3, HCN, CO, COZ, C&, CzH4, CZHZ, HCl, COS and H20.
476
Analysis of the combustion exhaust gases directly behind the combustor is performed by means of continuous on-line analysis for 02,CO, COz, NOx, N20 and s02. Besides gas analysis, also characterisation of solids has been performed. Samples of bed material, cyclone and filter ash have been taken. Proximate and ultimate analyses have been carried out as well as heating value determination. The test rig at Delft university is an aidsteam blown Pressurised Bubbling Fluidised Bed Gasifier (PFBG) with a ceramic channel-type filter and a modified pressurised ALSTOM Typhoon gas turbine combustor for the Low Calorific Value (LCV) gas produced. Table 2 shows the main design data. Compressed air and pressurised steam enter the gasifier through a central nozzle in the distributor plate after preheating between inner and outer vessel. Fuel, bed material and additive are fed from big bags on a conveyor belt and transported into a double valve lock hopper system followed by screw feeding into a vessel. From there, the material is fed pneumatically into the bed through a feed point in the bottom plate and directed toward the central nozzle. Bed contents can be kept constant by a removal system at the bottom of the reactor. The freeboard is well insulated and only contains probes. The ceramic filter consists of three honeycomb-like elements that are cleaned on-line one at a time using pulses of heated nitrogen. The combustor is of AGT Typhoon type and uses preheated air (350 "C) to combust the hot LCV gas. Gas sampling with respect to the PFBG installation is done after the ceramic filter unit, before and after the combustor. An off-line operated FTIR with a heated 2 m gas cell is used to measure N20, NO, NO2, NH3, HCN, CO, C02, S02, COS, C&, C2H4 and HCl. A GC is used to measure C1-C=, aliphatic hydrocarbons, Ar, C02, CO and N2 offline. CO, C02, SO2 are measured on-line using NDIR analysers. A micro-GC is applied for on-line H2 analysis. Paramagnetism based analysers are used for on-line O2 measurement. An NDUV analyser is used for NO,. A novel tar sampling method has been used, developed by Brage et al. [ 1 I]. Samples have been analysed by GC. Solids are sampled isokinetically before and after the ceramic filter unit. The experiments concerning the DWSA test rig have been performd with brown coal (Hambach open mine) and crushed pelletised pinewood sawdust. The experiments with the PFBG described in this paper have been carried out with pelletised Miscanthus Giganteus. Table 3 gives the main characteristics of the fuels. As can be seen, the pinewood has very low contents of nitrogen and sulphur. It is also observed that the heating value of brown coal is somewhat higher. Also, the ratio fixed carbon:volatiles is different for the fuels involved, so that a different behaviour of the fuels is expected during gasification.
477
Table 3 Fuel analyses (Raw basis) Brown Coal (Hambach)
Pinewood (crushed pellets)
Miscanthus Giganteus (pellets)
36.0 44.1 15.6 4.3
15.5 74.8 9.5 0.3
16.6 71.5 9.1 2.8
53.9 36.0 4.8 0.6 0.4 0.1
46.5 46.6 6.5 0.03 0.1 0.0
42.8 47.2 6.3 0.5 0.2 0.2
20.9
18.6
17.4
Proximate analysis: Fixed carbon (mass%) Volatiles (mass%) Moisture (mass%) Ash (mass%) Ultimate analysis: C (mass%) 0 (mass%) H (mass%) N (mass%) S (mass%) C1 (mass%) Heating value: HHV (MJkg)
EXPERIMENTAL RESULTS THE D WSA EXPERIMENTS Six experiments have been performed using the DWSA installation. The average length of stable operation was between 3.5 to 4.5 hours. Almost 1.5 hours were necessary to reach steady state operation conditions. Table 4 gives an overview of the main experimental results obtained using the DWSA test rig, regarding process conditions, LCV gas composition, heating value, carbon conversion and cold gas efficiencies. The main variables are the fuels applied, brown coal (BC) as a fossil fuel and pinewood (PW) as biomass species. For these fuels the air-stoichiometry was the most important process variable. With the pinewood experiments also pressure was varied. The LCV gas produced from the pinewood experiments at 0.51 MPa is somewhat lower in calorific value as that from the brown coal experiments at practically the same pressure and air-stoichiometric values and bed temperatures. Compared on a purge N2 free basis, the water concentration in the LCV gas from pinewood is higher, whereas the H2 concentration is lower, although there is less moisture and more elementary H in the wood as compared to brown coal. The light hydrocarbon concentrations are higher in the pinewood based LCV gas, whereas the CO concentration is significantly lower under the same conditions. A consequence of the above described differences is that the ignition of the gas produced from pinewood is more difficult than that of brown coal. The differences in the LCV gas composition are attributed to the structure in which C, H and 0 are bound in the fuels. Biomass consists of cellulose, hemi-cellulose and lignin. Whereas in older fossil fuels like brown coal more aromatic structures (pyridinic, pyrrolic) are present. This leads to different flash pyrolysis behaviour, the
478
initial step together with drying in the fluidised bed gasification process. A different yield of initial products results from this different behaviour. A sound mathematical description of this process is necessary to be able to model the whole gasification process. This is subject of study at Delft university, see Leentjes [3]. Carbon conversions observed are good and well above 90%.
Table 4 Overview of the DWSA gasification experiments 1
2
3
4
5
6
0.5 1 79 1 BC 3.7 7.3 0.30
0.5 1 802 BC 2.5 8.4 0.51
0.5 1 858
0.5 1 824 W 3.0 9.0 0.48
0.15 792
BC 2.2 9.2 0.66
0.5 1 782 PW 3.3 5.8 0.30
1.30 Total air stoichiometry; h tot (-) 2.5 Nitrogen flow to gasifier; $rn,NZ(kgh) LCV gas flow; ~ , L C V 12.8 (kgh) LCV gas composition 13.4 CO (vol%, wet) 10.4 H2 (vol%, wet) 2.1 C& (vol%,wet) 0.1 C2H4(vol%, wet) 10.9 C 0 2 (vol%, wet) 5.2 H 2 0 (vol%, wet) 57.3 N2 (vol%, wet) 0.5 Ar* (vol%, wet) 2290 NH3 (vol%, wet) 3.91 Higher heating value; HHV (MJ/Nm3) 94 Carbon conversion (%) (solid catch basis) 55.5 Cold gas efficiency 70 Fuel-N conversion to NH3 (%) 5 Fuel-N conversion to char-N (%)
1.31
1.47
1.29
1.29
1.46
2.8
2.7
7.2
3.0
1.4
13.6
14.1
16.3
14.9
6.5
8.9 7.0 1.4 0.1 11.8 5.3 64.8 0.5 1339 2.65
6.1 4.5 0.6 0.1 12.6 6.4 69.0 0.6 770 1.65
8.0 5.2 3.0 0.5 9.7 8.2 65.1 0.3 176 3.21
5.6 3.8 2.0 0.3 13.7 11.4 62.7 0.5 149 2.20
7.3 5.3 2.1 0.7 13.3 10.6 60.2 0.5 176 3.21
98
99
100
100
100
56.0 60
39.6 41
69.7 100
46.9 100
54.2 100
2
1
0
0
0
Experiment Pressure; P (MPa) Bed temperature; Tb ("C) Fuel Fuel flow (raw, kgh) Primary air flow (kgh) Primary air stoichiometry;
(-1
PW 1.5 3.6 0.52
* calculated value In the pinewood 'experiments, the fuel nitrogen is almost completely converted into NH3, with HCN and solid bound nitrogen being negligible considering experimental errors. The conversion of fuel nitrogen into NH3 with the brown coal experiments is 479
significantly lower compared to pinewood gasification, as also indicated in the literature with respect to pressurised bottom fed fluidised bed gasification, see e.g. Kurkela 141. The background of this different fuel nitrogen release behaviour can be explained (to a certain extent) by the different nature of the nitrogen bound in the chemical structures of the fuels, see e.g. Leppalahti & Koljonen [ 5 ] and Zhou [ 6 ] . In biomass, nitrogen is mainly incorporated in the form of peptide bounds (in e.g. amino-acids and protei'ns). In older fuels, like brown coal and (sub)bituminous coal, the appearance of the nitrogen species is more abundant in pyrridinic and pyrrolic structures. There are, however, also investigations which show different, significantly lower conversions of fuel nitrogen into NH3 with biomass as fuel, see e.g. Chen [7]. The understanding of the mechanisms of devolatilization is still not completely satisfying. Differences can possibly also be attributed to the gaseous environment in which the primary fast pyrolysis takes place, either oxygen rich or lean.
THE PFBG EXPERIMENTS Gasification results The experimental results obtained using the Delft PFBG test rig are presented in the form of graphs. The fuel used for the gasification experiments is pelletised Miscanthus in all cases presented. Dolomite is added to the fuel in a mass ratio of ca. 0.03-0.05. Steam to air ratio is 0.04-0.11 in these tests Figure 3 shows an increase of the concentations of light hydrocarbons, methane and ethylene, with decreasing air stoichiometry. The values are reasonably in line with the results obtained in the experiments with the DWSA installation indicated in table 4 for pinewood biomass (corrected to dry and purge nitrogen free values). Figure 4 shows an increasing trend of the higher heating value of the produced gas with decreasing air stoichiometry. The trend is in-line with earlier research concerning Miscanthus and Miscanthudcoal gasification experiments performed with the PFBG test rig, see e.g. de Jong [8]. The gas quality with respect to it's higher heating value was sufficient for stable pressurised combustion in the downstream ALSTOM Typhoon gas turbine combustor, see Hoppesteyn [9].
480
-
Pg
6
CI
5.5
i
5
N
z
7 @[CH4l8O0.5Mpa A[CH~]~R~~R. [CM] 8 a7 pa O[QH4J80.5Mpa A [QH4] 8 0.4 Mpa CI [QH4] 0 0.7 Mpa
4.5
a m L
4
a 3.5
4
P 3
3- 2.5
g
2
T 1.5
3 1 u 0.5
2
0 0.2
0.25
0.3
0.35
0.4
0.45
0.5
0.55
0.6
(-1 Fig. 3 Light hydrocarbon species concentration versus air stoichiometry. 4
++
m
--t
$ 3 Y
2.5
a,
a B> 2
*HHV 0 0.5 MPa
m
.-5C 1.5
WHHV00.4MPa
0
c $
.-cm
1
I 0.5
0.2
0.25
0.3
0.35
0.4
0.45
0.5
0.55
0.6
(-1 Fig. 4 Higher Heating value of the gasification product gas
One of the major problematic groups of components in gas produced by biomass gasification is tar. They contribute to fouling of equipment (e.g. gas engine or turbines) and to emissions in gas cleaning andlor combustion processes. These components have been defined to be organic aromatic species with a molecular weight higher than benzene, see e.g. Simell et al. [lo]. Figure 5 shows an increase of the specific tar concentrations of Polyaromatic Hydrocarbons (PAH) and Phenols in the produced LCV gas. These compounds have been quantified by means of the novel solid phase adsorption technique, developed at KTH Sweden, see Brage et al.[ 111. Especially the contribution of Phenols appear to be important at lower Air Factors. For (pressurised) 48 1
gasification using steam as (co-)gasifying medium the contribution of Phenols is also reported to be significant by Milne et al. [12]. Reproducible measurement of Benzene, Toluene and Xylenes with the s.p.a. sampling technique did not appear to be possible. The PAH’s analysed range from Indene, Naphtalene to Pyrene, with Naphtalene being the major species.
10000 a,
+ PAH 8 0.5 MPa
r
A PAH B 0.4 MPa
c.
e!
a,
P
rn PAH 8 0.7 MPa
4 2 1000 .o
0 Phenols B 0.5 MPa
a
A Phenols B 0.4 MP6
n
0 Phenol6 8 0.7 MPa
f8 E
Y
5s t
100
C
C
8L
3 10 0.2
0.25
0.3
0.35
0.4
0.45
0.5
0.55
0.6
1 (-1
Fig. 5 Higher hydrocarbon concentration versus air stochiometry of the gasification product gas of the PFBG gasifier. Figures 6 and 7 show the experimental PFBG results with respect to fuel nitrogen to Ammonia and Hydrogen Cyanide. These species are known precursors for NOx formation under e.g. gas turbine combustion conditions, which is a problem when dry, high temperature gas cleaning is applied, see e.g. Hoppesteyn [9]. From the results it can be concluded that a major part of the fuel-bound nitrogen is converted to Ammonia. This has also been indicated before, in the description of the pressurised fluidised bed pinewood gasification using the DWSA test rig, although somewhat lower conversion values are observed in the PFBG tests. The fuel nitrogen to NH3 conversion values are comparable to and in-line with values reported by VTT, where a slightly smaller scale pressurised fluidised bed is operated (ca. 500 kWth), for experiments with straw, a fuel quite comparable to Miscanthus, see e.g. Kurkela et al. [13]. Using straw with dolomite as additive they found values in the range of 60-71% fuel nitrogen conversion to NH3, at air stoichiometry values between 0.28 and 0.3 1.
482
-
h
8
5000
500
4500
450
4000
A [NH3] 0 0.4 MPa
L
o 3500
P
400 350
@ [NH3] 8 0.7 MPa 0 [HCN] 0 0.5 MPa
3000
300
A [HCN] 8 0.4 MPa
'9 2500 2000
250 200
7
1500
150
4
1000
100
500
50
~
n
g
P 2
-.
0
0.2
0.25
0.3
0.35
0.4
0.45
E
Fz
0 0.5
0.55
0.6
h (1
Fig. 6 NOx precursor species concentration versus air-stoichiometry in LCV gas produced by the PFBG gasifier 100
20
90
18 A Fuel-N to NH3 0 0.4 MPa
80
70
g
16
Fuel-N l o NH3 0 0.7 MPa 0 Fual-N to HCN 8 0.5 MPa A Fuel-N to HCN 0 0.4 MPa
14
z
12
50
10
p
8
-
2 40
3
g
60
Y
v
0
zI
30
6
20
4
10
2
1
u.
LL
0
0
L
0.2
0.25
0.3
0.35
0.4
0.45
0.5
0.55
0.6
h (-1
Fig. 7 Fuel nitrogen conversion to NH3 and HCN for the PFBG gasifier.
Ceramic filter hot gas cleaning results The ceramic filter unit in the Delft PFBG test rig has been operated for more than 100 hours under gasification conditions. They had to be exchanged once in the period for which the data presented here have been obtained. This was necessary, as due to opening of the unit to exchange a probe severe air leakage caused filter fire and
483
cracking. All filter elements of the second set have been equipped with thermocouples then. The filters have been operated during long stable period set points, of which figure 8 shows an example. Stable base-line pressure drop is observed, with values of the filter pressure drop between 10 and 16 mbar. Filter temperatures during stable set points (two periods in this case) of 650-700 "C are typical for all experiments carried out. The figure shows a transition from 0.7 MPa to 0.4 MPa, characteristic for a change of load in gasification (the air stoichiometry of the gasifier was kept constant for both set points). Gas cleaning efficiencies of ca. 99.95%have been obtained, with typical filter outlet dust loads of 5-10 mg/Nm3 LCV gas. These values are acceptable for gas turbine (combustor) operation considering also the sub-micron particle size of solids permeating the filter, see e.g. [14]. The values are also well below Dutch emission standards for power producing stations as well as waste incinerators, see e.g. Bergsma et al. [15]. 750 700 650 600 550 CI 500 450 3 400 350 E 300 250 200 150 100 50 0
30 28 26 24 22
g 8 g 8 e 8 g S g S S 8 i : g0
20 18 16 14 12 10 8 6 4 2 0
CI
jj
5
-
h
&sG$$4i$~K~oG F
F
time
Fig. 8 Characteristic temperatures and pressure drop behaviour concerning hot gas filtration for the PFBG test rig.
Gas turbine combustor results The combustion of the LCV gas generated by the gasifier and cleaned by ceramic filters is performed using a commercial downscaled ALSTOM Typhoon gas turbine combustor. The combustor is operated with swirling LCV gas and primary air flows. Swirl is used in these types of combustors to create an area of relatively low pressure in the primary combustion zone to generate a recirculating flow. This ensures rapid and intensive mixing of fuel, air and hot product gases, which enhances both flame stability and combustion efficiency, c.f. e.g. Beer [16]. For all experiments reported here, high combustion efficiencies of 99.4%and higher have been observed, see Hoppesteyn [9]. Gas sampling has been performed in the combustor by means of a quench-cooled probe.
484
Temperature profiles have been measured by a probe where the influence of radiation on the measurement has been minimised. A 2-D CFD model has been set up using FLUENT4.5 in a joint EU JOULE project with FLUENT and ALSTOM. As turbulence models the k-&model and Reynolds Stress model (RSM) have been applied. As chemistry models a chemical equilibrium model has been applied and on the other hand two models describing finite reaction chemistry, i.e. the laminar flamelet model and the reaction progress variable model. The comparison between experiments and the numerical results from the three chemistry models show that the chemical equilibrium model is sufficient to predict the combustion of LCV gas at elevated pressures, since deviation from chemical equilibrium is small due to the fast reactions. Hence no improvements are expected and have been observed from kinetically limited models. The RSM with constants C1 and C2 in the pressurestrain term proposed by Gibson and Younis [17] seems to yield the best predictions, however, the influence of the type of turbulence model (RSM or k- E) on the species concentrations and temperature predictions is not very large. One major topic of the work is an extension of the CFD code with a post-processor for the fuel nitrogen to NOx conversion. A reduced kinetic scheme has been obtained which describes quite well the combustion emission behaviour with respect to NOx formation. Figure 9 shows a comparison between experimentally observed NH, conversion to NO and modeling with this modified post-processor. The agreement is reasonably good for experiments at 0.5 and 0.7 Mpa. Significant deviation between model and experiment is seen for an experiment at 0.4 MPa. This is attributed to a measurement error at that pressure. A recommendation is to perform more experiments at this or lower pressure. ...............
...................................................
..
.
...
70
g
60-
Y
0 f
m
50-
I
400 .-
8
1
ln
30-
C
8
20-
0.3
0.4
0.6
0.5
0.7
0.8
Pressure [MPa]
Fig. 9
Measured and calculated fuel NH3-NO conversion in an LCV flame in experiments with the Delft PDU operated at different pressures
485
CONCLUSIONS AND FUTURE ACTIVITIES The test rig at IVD (50 kWth) and the Delft PDU (1.5 MWth) were operated successfully with respect to pressurised gasification, high temperature gas cleanup and combustion of biomass (wood and Miscanthus) and brown coal derived LCV gas. Carbon conversion and main gas composition were similar to data provided in literature by VTT, regarding their 500 kWth pressurised fluidised bed gasifier. Fuel-bound nitrogen was converted mainly to NH3 during gasification. For biomass the conversion values were significantly higher than for brown coal, which is also inline with VTT findings. A clear understanding of the nitrogen release behaviour is still not complete, as differences exist in literature regarding different gasifier set-ups, which can be attributed to different devolatilisation behaviour of the solid fuels. The gas cleaning by hot gas ceramic filtration was good, dust loads after the filter were below emission standards and fouling rules; stable combustion of the LCV gas was ensured without significant fouling. The performance of the gas turbine combustor in the Delft PDU with respect to main combustion product formation and fuel NO, was quite well predicted by FLUENT CFD modelling with a modified post-processor, although validation at pressures lower than 0.5 MPa is recommended. Further research work is necessary to extensively determine the influence of the biomass fuel type used and synergistic effects of brown coalkoal as additive to biomass with respect to Carbon conversion, tar formation and NO, precursor formation. Also, gasification modeling is further optimised and validated with these and more experimental data to describe the basic processes of drying, devolatilisation, partial combustion and gasification. Also, research attention will be paid to the fate of alkali’s and heavy metals, both model studies and experimental validation of the models will be performed. With respect to combustion, the research focus will be on improvement of the models regarding tars and their effect on emissions of CO, soot and unburned hydrocarbons (UHC). Part of that research will be the testing of a combustor integrated in an entire 500 kWth small scale gas turbine set-up.
REFERENCES 1. Kwant K.W. & Leenders C. (1999) Development of green energy market in the Netherlands and the perspectives of biomass. In: Proceedings of the fourth biomass conference of the Americas, (Ed. By R.P. Overend & E. Chornet), pp. 1629. Elsevier Science Ltd., Oxford. 2. Nagel H., Spliethoff H.& Hein K.R.G.(1997) Untersuchungen zum Einfluss des Hybridkonzeptes auf den Betrieb einer Druckwirbelschicht. In: Proceedings of the VGBconference: “Forschung der Kraftwerkstechnik 1998”, (2.5, pp. 1- 18.Essen. 3. Leentjes B. (2000) Kinetic modelling of afluidised bed gasifier. MSc. Thesis EV2008, Technical University Delft.
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4.
5.
6. 7.
8. 9. 10.
11. 12.
13.
14.
15.
16. 17.
Kurkela E. (1996) Formation and removal of biomass-derived contaminants in fluidised-bed gasification processes. VTT publications report No. 287, VTT, Espoo, Finland. L Leppalahti J. & Koljonen T. (1995) Nitrogen evolution from coal, peat and wood during gasification: literature review. Fuel Processing Technology, 43, 1-45. Zhou J. (1998) Fuel-bound nitrogen evolution during biomass gasification. Dphil thesis, University of Hawaii. Chen G. (1998) The reactivity of char from rapid heating processes under pressure; the role of the time-temperature-environmenthistory of it's formation. DPhil thesis, Royal Institute of Technology Stockholm. de Jong W., Andries J. & Hein K.R.G. (1999) CoaYBiomass cogasification in a Pressurised Fluidised Bed Reactor. Renewable Energy, 16, 1110-4. Hoppesteyn P.D.J. (1999) Application of Low Calorific Value Gaseous Fuels in Gas Turbine Combustors. DPhil thesis, Technical University Delft. Simell P., Stahlberg P., Kurkela E., Albrecht J., Deutsch S. & Sjostrom K. (2000) Provisional protocol for the sampling and analysis of tar and particulates in the gas from large-scale biomass gasifiers. Version 1998. Biomass & Bioenergy, 18 (I), 19-38. Brage C., Yu Q., Chen, G. & Sjostrom K. (1997) Use of amino phase adsorbent for biomass tar sampling and separation. Fuel, 76 (2), 137-42. Milne T.A., Abatzoglou N. & Evans R.J. (1998) Biomass Gasifier Tars: Their Nature, Formation and Conversion, IEA Biomass Utilization Task XIII, "Thermal Gasification of Biomass" activity report NRELRP-570-25357, USA, Golden. Kurkela E., Laatikainen-Luntama J., StAhlberg P. & Moilanen A., (1996) Pressurised j7uidised-bed gasification experiments with biomass, peat and coal at V7T in 1991-1994, part 3. Gasification of Danish wheat straw and coal. VTT publications report No. 29 1, VTT, Espoo, Finland. Ytsma S.D., Rijpkema L.P.M., van Loo, S., Schenk E.P. & Kiel J.H.A. (1998) Stookgasreiniging bij vergassing van biomassa-state of the art (in Dutch), Dutch Agency for Energy and Environment (NOVEM) report nr. 9829, The Netherlands, Utrech t . Bergsma G.C., Crouzen H.C., de Weerd G. & van der Werff T. (1999) Beperking van emissies naar de lucht bij conversie van biomassa naar elektriciteit en warmte (in Durch), Centrum voor energiebesparing en schone technologie, The Netherlands, Delft. Beer J.M. (1974) Combustion Aerodynamics. In: Combustion Technology: some modern developments, (Ed. By H.B. Palmer & J.M. BeBr), pp. 62-89. Academic Press Ltd., New York. Gibson M.M. and Younis (1986) Calculation of swirling jets with a Reynolds stress closure, Physics of Fluids, 29,38-48.
ACKNOWLEDGEMENT The work described in this paper is funded in part by the European commission in the framework of the JOULE R&D programme (contracts JOF3-CT95-0018 and JOR3CT95-0027) and the Dutch Governmental Agency for Energy and Environment (NOVEM, EWAB R&D programme: 355196/1150). Global Filter Systems Inc. and Foster Wheeler Energia Oy gave support in designing the filtration vessel and the pulse cleaning system. The authors wish to thank Dip].-Ing. H. Nagel for making it possible to perform measurements at IVD. 487
Tri-Generation from Biomass and Residues; Options for the Co-Production of Fischer-Tropsch Liquids, Electricity, and Heat C . Daey Ouwens, H. den Uil, and H. Boenigter Netherlands Energy Research Foundation (ECN), Unit ECN Biomass, P. 0.Box 1, 1 755 ZG Petten, The Netherlands
ABSTRACT Biomass is considered as an important energy source for this century. Several liquid fuels can be produce from biomass, the most important being methanol and Fischer-Tropsch (FT) liquids. The production of FT liquids is considered to be the most optimal route for producing as these fuels are directly applicable in the existing infrastructure and are free from sulphur and aromatics. The most optimal chain to produce these liquids is tri-generation, i.e. the combined production of liquid fuels, electricity, and heat in a Biomass Integrated Gasification Combined Cycle (BIGCC) system combined with a ‘once-through’ Fischer-Tropsch process. In large advanced plants (1000 MW*) the overall efficiency for tri-generation from biomass is higher than for conventional BIGCC: 59 vs. 50%. At a biomass feedstock price of 3.6 Euro/GJ, liquid transportation fuels can be produced for 0.45 Euro/litre, which corresponds to 47 Euro/ton avoided COz emission. These costs are sigmficantly below the value of 95 Eurohon for a BIGCC. For small-scale systems the local situation will determine the cost figures. Within the EU and Dutch governmental objectives to increase the contribution of renewable energy, tri-generation might prove to be the optimal option for the use of governmental stimulants and tax regulations.
INTRODUCTION Biomass and residue (‘waste’) based energy currently contributes to about 3% of the primary energy consumption in the European Union (EU). This share will have to be tipled within the next decade to fulfil the EU objective of 8.5% energy from biomass and residues in 2010. It is also a prerequisite to enable a reduction of greenhouse gas emissions as documented in the Kyoto-protocol. Within the framework of renewable energy sources, biomass is considered as an important, and maybe the most important, energy source for this century [1,2]. Discussing the introduction of biomass, we have to realise that the attractiveness as a large-scale, modern energy source has come up quite recently. This means that most of the applications are still in a first stage of development. Commercial implementation 488
of biomass technologies on the European energy market has remained modest, despite remarkable efforts in research, development, and demonstration in the last two decades. Only large-scale combustion significantly contributes to the energy demand and mainly residual wood is used as biomass fuel at present. The residue (‘waste’) policy of the European Union is characterised by strict emission requirements on landfilling and (thermal) processing and by landfilling bans on high caloric residues. To reach the C 0 2 reduction goal also residual streams have to be brought in for energy generation. An important aspect of biomass liquid (bio-)fuels can be produced fi-om this renewable source [3]. This paper discusses installations for the combined production of liquid fuels and electricity and heat: tri-generation.
ADVANCED OPTIONS: TRI-GENERATION Tri-generation is an advanced syngas application in which (liquid) fuels are produced in addition to the direct generation of electricity and heat. In the case of tri-generation biomass and residues are gasified and between the biomass gasifier (including gas cleaning and compressor) and the combined cycle (CC) a once-through conversion unit to produce the fuel is placed (Figure 1). Inputs can be various biomass feedstocks and residual streams.
Biomass (S residues) ‘green’ Diesel
T
Heat
0 FIscher-Tropach Synthesis
Figure I . Schematic representation of the tri-generation principle.
A once-through process will be used in combination with tri-generation as in this way the product output of the plant can be optimised for market prices or needs. The co-production of electricity, heat, and liquid fuels leads to higher efficiencies and lower overall costs. In the case of a once-through process and tri-generation between 30 to 50 % of the energy of the gas is converted into liquids and the rest is converted into electricity and heat in a combined cycle. About 30% of the energy content of the gas converted into liquids is released in the form of heat, which can be used to generate additional electricity with an integrated steam cycle. 489
CONCEPT For large-scale biomass conversion (400-1000 MW&)the biomass Integrated Gasification Combined Cycle (BIGCC) is a well-known concept [4]. Compared with combustion the system has some clear advantages: 0
0 0
8
A high overall efficiency for the production of electricity of more than 50% can be obtained. Only a small heat exchanger needed for the cooling of the syngas after the gasifier. The gas cleaning is also smaller as gas is cleaned before utilisation in the IGCC. Very low emissions can be obtained for compounds like NO,, SO2, HC1, dust, etc.
For this concept atmospheric gasification, with air as the gasification medium, has been chosen. In the future pressurised systems may be preferred. All main components of such a system do exist on the market: gasifier, cleaning system, compressor, Fischer-Tropsch unit, gas turbine, and the steam cycle. Large-scale, as well as smallscale (15 to 100 MW*) systems belong to the possibilities. PREREQUISITES The combination of high value outputs, scale-effects and flexibility in general, may lead to a very efficient, cost-effective and clean overall system. Cost-effective means that the electricity produced has to compete with electricity from coal. For fair comparison the cost of ‘clean’ electricity generation must be considered with a similarly low environmental burden. For instance, costs for removal or storage of the emitted carbon dioxide from coal plants have to be included. For electricity this implies an allowable cost price of 0.05 EurokWh. This value is the EU strategic goal and seems achievable for large advanced systems [5] where electrical efficiencies between 50 and 60% can be obtained [ 6 ] .In the case of the produced liquid fuels, cost-effective means that the cost price must correspond to the cost prices of fossil fuels, which will be used in the future in cars with low or very low emissions. For the liquid fuels FT diesel is taken as example and the cost price would have to be 0.34 EuroAitre (of diesel). The other prerequisite is the fact that society not only expects competitive, but clean systems and products as well. One has to look at the whole chain of production and conversion of biomass to get a clear picture of the environmental consequences. For the feeding material there is a strong preference for ‘woody’ or ‘grassy’ materials [7]. The emission of the power plant has to be low and the strict Dutch rules for waste incineration installations are taken as a point of departure [5]. The wholeintegrated system has the further advantage that it produces renewable products for which there is a market today: electricity, heat, and FT-liquids. Furthermore, the liquid fuels form an attractive energy carrier and storage medium. CHOICE OF LIQUID FUEL Possible products and processes that are considered are the production of methanol (and ethanol) and hydrocarbons with the Fischer-Tropsch (FT) synthesis. Also chemicals like dimethylether (DME) and dimethoxymethane (DMM) can be produced.
490
A consideration opposing the methanol route is that it does not lead to a product for direct application as fuel (additive), because it must fnst be converted to methyl tertiar-butyl ether (MTBE). In Europe there is no infrastructure for the large-scale use of MTBE, whle in the USA (California) a ban on MTBE is expected. Furthermore, methanol from syngas has the same properties as ‘normal’ methanol. On the long-term methanol is a fuel to be used in fuel cells. In 1999, in the Netherlands a government-initiated study was completed to determine the optimal route for producing energy camers from renewable sources (e.g. biomass) [S].The production of synthetic transportation fuels (especially diesel, and to a lesser extent gasoline) with the Fischer-Tropsch technology was considered as a promising route. In the evaluation aspects were included concerning technology, economy, existing infrastructures, and COz reduction potential. The FT process produces a diesel-like middle distillate fuel with a high-cetane number. Some special aspects of the oil are: 0
0 0 0
0
0
Free of sulphur and aromatics; in contrast to conventional fuels from fossil origin. High energy density (between 30 and 40 MJ per litre). Suitable for long distance transport and long term storage. Directly applicable in the existing infrastructure (distribution network and vehicles). FT fuels can be blended with the conventional fuels to meet the increasingly strict European Union fuel specifications. FT oils can be used - in the future - in fuel cells.
Primarily inspired by these arguments ECN decided to concentrate their research on the Fischer-Tropsch option for the short-term production of renewable transportation fuels.
FISCHER-TROPSCH PROCESS The FT process is well known and already applied on a large scale [9,10,11,12]. Currently, the two players that operate commercial Fischer-Tropsch plants are Shell and Sasol. In the Sasol and Shell plants gasification of coal and partial oxidation of natural gas, respectively, produce the syngas for the FT synthesis with well-defmed compositions. Shell operates the SMDS (Shell Middle Distillate Synthesis) process in Bintulu, Malaysia, which produces heavy waxes with a cobalt catalyst in multi-tubular fixed bed reactors. Sasol in South Africa uses iron catalysts and operates several types of reactors, of which the slurry bubble column reactor is the most versatile ( i e . applied in the Sasol Sluny Phase Distillate; SSPD). SYNTHESIS
The Fischer-Tropsch (FT) synthesis involves catalytic reactions in which CO and HZ are reacted to form mainly aliphatic straight-chain hydrocarbons (C,H,). The kind of liquid obtained is determined by the process parameters (temperature, pressure), the kind of reactor, and the catalyst used. Typical operation conditions for the FT synthesis are a temperature range of 20O-35O0C and pressures of 15-35 bar, depending on the 49 1
process. The subsequent chain-growth is comparable with a polymerisation process resulting in a distribution of chain-lengths of the products. The distribution of the products depends on the catalyst and the process operation conditions (temperature, pressure, and residence time). In general the product range includes light hydrocarbons (C, and C2), LPG (C3-C4),gasoline (C5-C,,), diesel (C$-Cz5),and waxes (>CI9).Besides these straight-chain hydrocarbons also branched hydrocarbons, unsaturated hydrocarbons (olefins), and primary alcohols are formed in minor quantities. In the (exothermic) FT reaction one mole CO reacts with two moles of H2 to afford a hydrocarbon chain extension. The oxygen from the CO is released as product water: CO
+
2H2
3
-CH,-
+
H,O
AH = -165kJlmol The reaction implies a H$CO ratio of at least 2 for the synthesis of the hydrocarbons. When the ratio is lower it can be adjusted in the reactor with the catalytic Water-Gas Shft (WGS) reaction: CO
+
H,O 3 CO, AH = -42kJ/mol
+
H,
When catalysts are used with WGS activity the water produced in the reaction can react with CO to form additional H2. In this case a minimal HJCO ratio of 0.7 is required and the oxygen from the CO is released as C02: 2CO
+
H, -CH,AH = - 204 kJ/mol
+
CO,
(3)
CATALYSTS Several types of catalysts can be used for the Fischer-Tropsch synthesis - the most important are based on iron (Fe) or cobalt (Co). Cobalt catalysts have the advantage of a higher conversion rate and a longer life (over five years). The Co catalysts are in general more reactive for hydrogenation and produce therefore less unsaturated hydrocarbons and alcohols compared to iron catalysts. Iron catalysts have a hgher tolerance for sulphur, are cheaper, and produce more olefin products and alcohols [13]. The lifetime of the Fe catalysts is short and in commercial installations generally limited to eight weeks.
NETHERLANDS ENERGY RESEARCH FOUNDATION (ECN) The Netherlands Energy Research Foundation (ECN) is the leading institute for renewable energy research in the Netherlands. One of the ECN priority research areas is biomass, in which the unit Biomass focuses on the thermal-conversion of biomass and residues (B&R) to generate electricity and heat, liquid and gaseous fuels, and fine chemicals [14]. ECN Biomass develops and operates installations for thermal conver-
492
sion of biomass and residues, for example a 0.5 MW Circulating Fluidised Bed gasifier, a 0.35 MW fixed bed gasifier, a 15 kW pyrolysis reactor with a thermal tar cracker, and a test facility for gas cleaning with a gas engine. Within the Fischer-Tropsch research ECN Biomass concentrates on the definition of the gas cleaning with respect to the typical B&R impurities, llke N H 3 , HCl, HCN, H2S, COS, tars (heavy organic molecules), soot, and alkali metals. Traces (< ppm) of these compounds can already be a poison for the Fischer-Tropsch catalysts. For the implementation of B&R and Fischer-Tropsch ECN its strategy is on the demonstration of integrated systems to reduce the time necessary to realise a first full-scale installation for conversion of biomass and residue, gas cleaning, and Fischer-Tropsch synthesis. To achieve this ECN focuses on two lines of development: 1. Large-scale dedicated installations for gasification of imported biomass. 2. Small- to medium scale installations for the conversion of local biomass and residues.
LARGE SCALE INSTALLATIONS When the purpose of the installation is to produce green electricity from biomass the scale will have to be large - estimated is 400 up to 1000 MW thermal input. This is imposed by the necessity of large-scale import at relative high costs. Due to the economy of scale cost-effective products can be made even with the high feedstock price and is it also economic to install extensive gas cleaning to remove all impurities. In this development ECN Biomass co-operates with Shell Global Solutions and the integration is focused on the SMDS cobalt catalyst fixed-bed FT technology. In t h s ‘large-scale’ approach ECN is performing System Studies to determine the technical and economical optimal process. This joined project is partly funded by the Dutch government (Stichting Duurzame Energie; SDE). A follow-up project for the demonstration on bench-scale will start by the end of 2000.
SMALL- TO MEDIUM-SCALE FISCHER-TROPSCH The utilisation of local biomass and residues (construction and demolition wood, park and garden residues) can be decentralised in most cases, imposed by fuel availability and logistics. In this case the size of the conversion installations is generally limited to small or medium scale (15-100 MW thermal input). For these installations the cost of products distribution is a major parameter in the economics (Figure 2). For example, the cost to transport one PJ of oil is approximately ten times lower than to transport the same energy content as (natural) gas. On its turn transport of gas is ten times cheaper than electricity, which on its turn is ten times cheaper to be transported than heat. For small- and medium scale B&R fuelled installations the production of heat as major product is therefore the least attractive, unless they are fully integrated with other installations on site or deliver to a city heat gnd. The economic most attractive option is to produce liquid transportation fuels.
493
10000
r
se e
t: 1 0
44L
I000
100
70
I
Liquid Fuel
Natural Gas
Electricity
Heat
Figure 2. Indicative presentation of the relative costs for transport of energy.
Choice of Catalyst
The product gases from the conversion of B&R are generally not well defined and may contain variable impurities, due to the varying composition of the B&R feed. An indicative composition for the main constituents of a product gas is shown in Table I.
Table I . Indicative composition of the product gas fiom an air-blown CFB gasifier with a wood biomass fuel [4]. Gas
CO
H2
C02
H20
CH4
C2H4
N2
traces
r%vOii
17.2
13.3
12.1
13.7
2.8
0.9
39.7
0.3
On this scale an extensive gas cleaning section (necessary to protect the catalyst) will be too expensive. Therefore, an operational argument to select Fe catalysts is to reduce the impact of impurities. The Fe catalyst is less sensitive compared to Co and the decrease in activity is less dramatic as the normal life span is already relatively short. A cheap catalyst with a higher replacement frequency is preferable. Furthermore, the use of an iron catalyst is also preferred considering the H2/C0 ratio of 0.7 to 1.2 for B&R product gas. Choice of Reactor
The slurry bubble reactor is becoming the preferred type for B&R based FT [13]. A slurry reactor is a vertical vessel in which the catalyst is suspended in the reactor solvent. Typically the catalyst is present in 10 to 35 wt?? and both types of catalyst can be used. The solvent is an oil or wax of heavy hydrocarbons (the products of the FT synthesis). The feed gas is fed into the reactor through the bottom and the unconverted syngas and the gaseous hydrocarbons emerge from the top. The liquid products are filtered fFom the catalyst and drained. The use of a slurry reactor offers several advantages: ( i ) good heat transfer and thus a homogeneous reaction temperature, ( i i ) high through-put, (iii) high single-pass catalyst productivity, (iv) possibility to regenerate
494
catalyst during operation, (v) relatively simple construction and operation, and (vi) low investment costs [ 151.
Choice of Operation The production of hydrocarbons with the Fischer-Tropsch synthesis from B&R product gases can be economic feasible. To reduce investment costs the FT synthesis should be operated in once-through mode without an expensive recycle and have a high conversion to optimise the economics. The unconverted syngas and small (gaseous C1-C4) hydrocarbons are used to produce electricity and heat (pi-generation). The electricity can be used in the process or supplied to the grid, while the heat is preferably integrated or supplied to installations on site because of the costs of transportation.
ECN Activities ECN Biomass is currently realising a lab-scale FT bubble slurry reactor to be integrated with the available thermal conversion installations and gas cleaning facility. The reactor should be operational at the end of 2000. It will be a challenge to define operational limits for a FT process that runs with B&R derived syngases of variable compositions. Changes in the syngas compositions might origin from different feed (mixtures) used for the gasification, to (small) discontinuity in the feeding, or to variations in the gasification process. Especially in smaller scale installations the latter are more likely to occur as the ‘buffering’ capacity of the installation is smaller.
ECONOMICS An economic benefit is that the FT fuels are of higher quality than conventional fuels,
as they contain no sulphur and aromatics, and burn cleaner. Higher prices for the FT product can therefore be reasoned, maybe not (yet) for the large consumer vehicles, but especially for niche applications where stricter fuel regulations are imposed by (local) governments (e.g. for city buses). Furthermore, installations solely fuelled with biomass may benefit from tax regulations, as FT fuels produced from biomass are renewable (i.e. ‘green’). With more tightened regulations, the fossil fuels may even become more expensive. Biomass is limited available in Europe and the generation of green energy (liquid fuels and electricity) will depend on large-scale import. Therefore, available biomass should be utilised in the most (cost) efficient way. Within the framework of the COz reduction policy this can be quantified as the costs per avoided ton COz emission.
COSTS FOR LARGE-SCALE To come to demonstration and in a later stage to realisation of commercial plants the costs are a very important aspect. Large-scale systems are therefore the most attractive as they benefit from the economy of scale. An economic evaluation is made for future and more advanced plants [4,16].
495
Energy balance For a 1000 MWh gasifier and with values for efficiencies are used that seem achievable for future systems, an overall energy balance of the tri-generation plant is made. The product gas from the gasifier contains 80% of the energy from the biomass with a composition as shown Table 1 . Note that the CO and H2 contain only -70% of the energy of the gas. The CH4 and C2H4,both inert for the FT synthesis, contain the remainder. For the FT synthesis is assumed a CO conversion of 80% and a selectivity of 80 wt% for the liquid hydrocarbons (C5+). The liquid products contain 28% of the energy and a significant amount of the energy is released as heat from the exothermic reactions. With the unconverted gases and the C1-C4 FT hydrocarbons electricity is generated in the CC. Additional electricity is generated from the heat streams from the gasification and the FT synthesis. The overall efficiency of the tri-generation process is 59% (excluding the plants own electricity consumption).
Economic evaluation The determination of the scale of the units and the investment costs is based on the overall energy balance. The following considerations are made: For the feeding material 3.63 Euro/GJ has been taken. Th~sfigure is hgher than most other studies use but it seems realistic for large-scale imported biomass. The investment costs for the gasifier with gas cleaning are taken as 230 EurokWh, It is assumed for simplicity that the gas cleaning in fiture advance systems is Sufficient for the FT synthesis. For the investment for the Fischer-Tropsch synthesis unit is used 10,200 Euro/barrel per day capacity (BPD).This value is based on 30% of the cost of a FT-unit with natural gas gasification of $30,00O/BPD. For this large CC investment costs of 570 EurokW, are used. The cost for fossil electricity has been taken as 0.032 EurokWh, while for green electricity a price of 0.050 EurokWh is used (the EU strategic goal for renewable electricity). Furthermore are used: depreciation 13.1% of the investment cost, operational & maintenance (O&M) costs 5%, and 8000 operational hours. When all the costs are taken together, a value of 0.45 Euro/litre is calculated for the production costs of the FT liquids (based on the net electricity production and when the electricity is sold as green). This value is higher than the cost price of fossil transportation fuels of 0.34 Euro/litre, but sufficiently below the market price (in The Netherlands in October 2000) of 0.91 Euro/litre to have room for tax incentives from the government. In the light of the EU objective to increase the amount of energy from biomass and to reduce the C02 emissions the cost per ton avoided C02 are important indicators to determine the changes of implementation and receiving tax beneficence. When fossil prices are used for both the electricity and the FT liquids the cost per ton avoided C02 emission are calculated to be 47 Eurolton C02 [171. For comparison, the costs per ton C02 for a BIGCC without FT unit are 89 Eurohon. This significant difference is partly due to the higher overall efficiency of the tri-generation process (59 496
vs. 50%). However, the lower investment costs for the FT unit, compared to the CC, are the major contributor. Using part of the gas in a FT unit and a smaller CC results in lower overall investment costs (when the CC is still large enough to benefit from the economy of scale). The main losses in the process are in the form of heat. The efficiency increases significantly, and the cost of the FT liquids decrease, when this heat can be utilised in the process or supplied to a city heat grid.
COSTS FOR SMALL-SCALE For small-scale systems the local situation will determine the targets of the costs prices for electricity and the fuels [18]. For example, gate fees for residue (‘waste’) processing and local heat demands for city heating have to be considered. In general, higher investment costs are allowable as residues are less expensive than biomass (or even have a ‘negative’ cost price).
CONCLUSIONS The production of Fischer-Tropsch liquids (diesel) is today the optimal route for producing energy carriers from biomass. The FT diesel is directly applicable in the existing infrastructure and is free from sulphur and aromatics; in contrast to fossil diesel. For the demonstration of integrated systems ECN focuses on two lines of development: (1) large-scale dedicated installations for imported biomass and (2) small- to medium scale installations for local biomass and residues. For small-scale systems for the processing of local biomass and residues the local situation will determine the cost figures. With a large-scale (1000 MWm) biomass plant and with biomass feedstock prices of 3.6 Euro/GJ, liquid transportation fuels can be produced for 0.45 Eurollitre. This corresponds to cost per ton avoided C 0 2 emission of 47 Eurohon C02, significantly below the costs per ton C 0 2 for the BIGCC. Within the EU and Dutch governmental objectives to increase the contribution of renewable energy, tri-generation might prove to be an optimal option for the use of governmental stimulants and tax regulations.
REFERENCES AND NOTES 1. 2. 3.
4.
Shell (1996) The evolution of world’s energy system. Johansson T., B., Kelly H., Reddy A.K.N. & Williams R.H. (June 1992) Renewables for fuels and electricity, UNCED. Larson E.D. & Jin H. (1999) Biomass Conversion to Fischer-Tropsch Liquids: Preliminary Energy Balances. In: Proceedings of the 4th Biomass Conference of the Americas, Vol., pp. 3-853. Faaij A. & Van Ree R. (1998) Long term perspectives of biomass integrated gasification with combined cycle technology; costs and efficiency and a comparison with combustion, Novem, Utrecht, The Netherlands, EWAB report 9840.
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5. 6.
7. 8.
9.
10.
11.
12. 13.
14. 15. 16.
17.
18.
New standards for biomass (Autumn 1999) Dutch Ministry of Environment (proposal). The main cause of the high efficiency is the use of a gas turbine with intercoolers, recuperation, reheater, and a 100°C higher turbine inlet temperature than currently applied. Daey Ouwens C. & Faaij A. (2000) Sustainable Energy: New Challenges for Agriculture and Implications for Land Use, Wageningen, to be published. GAVE final report (in Dutch), NOVEM Publication Centre, 1999. (a)Hunt D. (1983) Synfuels handbook, Industrial Press Inc., New York. McGraw-Hill. (b) Encyclopaedia of Science and Technology (1992) 7th ed., McGraw-Hill Inc., New York. Larson E.D. (June 1999) Advanced Technologies for Biomass Conversion to Energy. In: Proceedings of the 2nd Olle Lindstrom Symposium on Renewable Energy, Bio-Energy, Royal Institute of Technology, Stockholm, Sweden. (a) Hedden K, Jess A. & Kuntze, T. (1994) From Natural Gas to Liquid Hydrocarbons. In: Edrol Erdgas Kohle, part 1; 110,7/8, pp. 318-321. (b) Idem, (1994) part 2; 1 10,9, pp. 365-370. (c) Idem, (1995) part 3; 111,2, pp. 67-71. (d) Idem, (1997) part 4; 113,12, pp. 531-540. Daey Ouwens C. (1999), Cost of bio-oil produced by the Fischer-Tropsch process in a BGICC, Report Eindhoven University of Technology. RenTech Inc. (July 1999) Gas to Liquid Home Page: www.gastoliauids.com. Downloaded from this page the “Complete Howard Weil Fischer Tropsch ‘GTL’ Technology Report”. Netherlands Energy Research Foundation (ECN):www.ecn.nl. Bhatt B.L., Frame R., Hoek A., Kinnari K., Rao V.U.S. & Tungate F.L. (1995) Topics in Catalysis 2,2, 235-257. Daey Ouwens C., Faaij A. & Ruyter H.P. (2000) Flexible, competitive production of electricity, heat, bio-fuels and ethanol by tri-generation, In: Proceedings of Sevilla, to be published. Numbers used: COz production from natural gas: 56.1 kg C02/GJ, electricity: 90.5 kg C02/GJ (62% efficiency from natural gas), fossil oil: 73.3 kg C02/GJ, and liquid transportation fuels 8 1.4 kg C02/GJ. Larson E.D. & Jin H. (1999) A preliminary assessment of biomass conversion to Fischer-Tropsch cooking fuels for rural China. In: Proceedings ofthe Fourth Biomass Conference of the Americas, Addendum, Overend R.P. & Chornet E. (eds), Oakland.
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Concept for a Decentralised Combined Heat and Power Generation Unit for Biomass Gasification I. Romey a, M. Adorni a, J. Wartmann a, G. Herdin b, R. Beran ', K. Sjostrom d, Ch. Ros6n a University of Essen, Technology of Energy Supply Systems and Energy Conversion Plants, Universitats Str. 15, 451I 7 Essen, Germany Jenbacher AG, Achenseestrasse I-3,6200 Jenbach, Austria Technical University of Graz, Institute for Internal Combustion Engines and Thermodynamics, Kopernikusgasse 24, 8010 Graz, Austria Royal Institute of Technology,Department of Chemical Engineering and Technology / Chemical Technology, Teknikringen 42, 100 44 Stockholm,Sweden
ABSTRACT The development of an improved process for decentralised combined heat and power generation from biomass gasification was the main goal of h s project. Based on gasification tests performed in a bench-scale pressurised fluidised gasifier a concept for an allothermal gasification unit was made. Air-blown as well as steamblown tests were carried out considering three different kinds of biomass (birch, salix, and crushed pelletised straw) against a wide spectrum of operation conditions with different bed-materials (silica sands, magnesite and dolomite). The aim of the gas quality optimisation was the production of a fuel gas (syngas) with a h g h Hz and low tar content suitable for a gas engine. The tar content was measured with SPA method and a gravimetric method. The tests were performed at the Royal Institute of Technology, Stockholm, in close co-operation with the University of Essen. Jenbacher AG designed a new type of cylinder head for the direct feeding of the hot pressurised syngas into the gas engine in co-operation with the T e c h c a l University of Graz. Taking the results obtained into account a concept of decentralised combined heat and power generation (CHP) unit with an electrical power output of 1 MW was set up. INTRODUCTION Biomass is an environmental benign source of energy and a C02-neutral fuel for heat and power generation. Wood can be used either directly from forestry or also as waste product from wood processing industry, e.g. saw dust. The advantage of gasification
499
compared to combustion is that generally a higher overall efficiency cap be achieved
PI.
Often air is preferred as gasification agent, however, steam offers some advantages despite a higher apparatus expenditure. The syngas has a high hydrogen content and a high heating value, because it is not diluted with the nitrogen from air. Therefore, the design was based on the idea of an allothermal steam gasification process. The biomass fuel should come from the local area as a long distance transport can be economically not justified due to its small heating value. Thus, the economical size of a plant is directly coupled with the availability of the biomass fuel nearby. Hence, a decentralised combined heat and power generation (CHP) unit with an electrical power output of 1 MW, was chosen. GASIFICATION TEST THE EQUIPMENT
The biomass gasification tests were performed in a laboratory development unit (LDU). The pressurised fluidised bed reactor with the connected high-temperature filter is shown in Figure 1. Several reconstructions have been made to give today's design with a maximum operating pressure of 3.0 m a . The LDU was constructed for a feeding rate up to 15 kgh of biomass fuel using different kinds of fluidising and gasifymg agents such as air, steam and carbon dioxide. The maximum temperature is 1000°C for the reactor and 500°C for the filter. The fuel hopper and feeder has a fuel capacity of 120 litres. The reactor is fed from the top directly into the fluidised bed through a cooled and Coollng water
I
Flow chart LDU Date: QQ-0525
draw :J. Warimann and C. Rosbn
Figure 1 Laboratory Development Unit LDU 500
insulated pipe (see Figure 2). The 1.5 meter long reactor consists of two different zones, the reaction zone at the bottom and the freeboard in the upper part. The diameter of the reaction zone and of the freeboard are 0.144 m and 0.2 m, respectively. The inner reactor wall and the bottom which is designed like a perforated cone are made of Inconel steel. The reactor temperatures and pressure drops are continuously monitored. Thermocouples measure the temperature at nine levels and the pressure drops in the bed and over the whole reactor are controlled by pressure pipes. The preheated fluidising agents pass through the perforated cone and fluidise the bed material. The syngas exits the reactor at the top and is transported via a heated pipe to the filter. In the high temperature filter the syngas is cleaned from ash and char dust particles by three 1 m long Inconel steel tube socks that are heated to prevent condensation of volatiles and tar. The pressure valve is the last heated part, placed after the secondary reactor, before the syngas is cooled in two water coolers at atmospheric pressure. Most of the tar and the produced water are condensed in the two water coolers. The remaining aerosols are trapped in two 5 p filters. The analysis of the syngas is made with a gas chromatograph, connected on-line with analyses every ten minutes. II
7
! ! I
i
! ! !
i
i
! !
I
Heater
Figure 2 Pressurised Fluidised Bed Reactor
50 1
Table 1 Ultimate Analysis of the Different Biomass Fuels
Rawmateria' Straw pellets Salk Birch wood
C
H
N
0
S
HHV**
(wt%mt)
(wt%mf)
(wt%mf)
(wt%mf)
(wt%@
MJkg
38.86 41.75 44.4
0.24 0.05 <0.1
19.1 19.2 19.3
45.4 50.0 49
6.1 5.9 6.1
1.7
0.7 <0.2
*mf: moisture free; ** maf moisture and ash free
The three biomass fuels available for this project (straw pellets, salix, birch wood ) were tested at the LDU. The drymg procedure for the fuels included oven-drymg at 105°C followed by an open storage to reach equilibrium moisture content with the
surrounding air. Thereby, a very low moisture content can be achieved. The final treatment of the biomass is milling and sieving, through which a particle size distribution of 1.0-3.2 mm was obtained. A high biomass moisture content influences the biomass flow quality negatively during the final fuel treatment steps as well as during the feeding to the gasifier. The characterisation of the biomass fuels is shown in Tables 1 and 2. Table 2 Proximate Analysis of the Different Biomass Fuels
Raw Material Straw pellets Salix Birch wood
Moisture 12* 8.6* 7*
Fixed carbon Volatiles 11.9 14.-9
80.4 82.6
n.a.
n.a.
Ash 7.7* 1.6* 0.2*
*average of different analyses RESULTS FROM THE GASIFICATION TESTS Overall 60 different gasification tests were performed. The steam gasification tests were carried out in temperature range of 700 to 850°C and mainly at a pressure of 0.5 MPa. At the air gasification higher temperatures were achieved (700 to 930°C) at pressure levels of 0.5 MPa and 1.0 MPa. Finally some tests with steam and air were done. The preliminary evaluation has shown that the tar yield was mainly influenced by the temperature. An increase of the gasification temperature reduced the tar content of the syngas, especially at high feeding rates. The tar yield was not influenced significantly by the two different gasification agents. The use of calcined dolomite as catalytic bed material influenced the tar yield positively, however, it must be taken into account that the partial pressure of steam and hydrogen is limited, because both components poison the catalyst. It was observed that calcined dolomite reacted with the steam or with hydrogen to hydroxides and the particles agglomerated. When magnesite was used the same effect was observed at a temperature of about 45OOC at which the filter system worked. The entrained bed material led to a filter obstruction. The results from the steam gasification have shown that it was possible to achieve high hydrogen content but the feeding rate was lower compered to the air gasification 502
test due to a slower char reaction kmetic. Higher feeding rates were feasible when air was admixed. In Table 3 some typical results from the gasification tests were shown. In all tests birch was used as biomass fuel.
Table 3 Typical Syngas Composition from Different Tests
Pressure in MPa
Incom 14 0.5
Bed material
MgO
Test name
Bed temperature in "C Fuel feeding rate in kg D S h
900
Air tests* Steam tests** Air*/Steam Incom Incom IncS 12 IncS 10 IncS 19 IncS 20 19 16 0.5 0.5 0.5 1.o 0.5 0.5 Silver- Silver- Dolo- Silver- Olivine Olivine sand sand sand sand mite sand 780830 800 700 800 700-750 780
4.25
3.16
2.41
Total tar in g/Nm
9.9*
17*
18.2*
N2
H2 CH4 C2H2 C2H4 C6H6
43.9 19.8 17.4 9.8 6.8 0.4 1.6 0.34
h O2 gasification
0.18
co co2
1.25
45'33 **
1.04
4.75
6.55
125**
29.1*
27.1*
Syngas composition in % vol (moisture free): 55.3 48.1 3 1.87 17.6 17.4 16.67 6.16 26.12 15.4 16.6 24.0 31.88 20.55 3.7 8.4 48.67 55.80 23.16 5.8 7.4 8 5.07 6.5 0.3 0.6 1.0 0.36 0.55 1.6 1.2 1.53 0.36 0.83 0.30 0.32 0.13 0.04 0.07 0.2 1
0.2
Steam mass flow 11 in k g h Syngas HHV in 6.2 7.2 11.9 7.6 MJ/Nm3 *recalculated to air gasification; ** nitrogen free and dry
8.3 9.6
36.66 23.34 14.48 12.98 9.9 0.54 1.62 0.46 0.14 0.12(fix) (fix) 8.5 8.0 9.1
9.8
The tar samples were taken from the syngas at the gasifier outlet for about one hour after stable operational conditions were achieved. Two different kinds of tar measurement were done. The solid phase adsorption (SPA) was a device for tar sampling developed at the Royal Institute of Technology, which measured chromatographable tar (e.g. neutral compounds and phenols). Chromatographable tar or "light" tar can be combusted without problems and usually does not condense in the piping leading to the gas engine. The second gravimetric tar sampling device measured "light" as well as "heavy" tars. It must be considered that tar was defined as a mixture of organic components with a boiling point above 78°C. Hence, benzene was not included.
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CONCEPT FOR COMBINED HEAT AND POWER GENERATION HEAT SUPPLY OF THE GASIFIER The heat supply of the LDU was done electrically. Thereby, the entering gasification agent was heated. For the concept it is chosen to operate with a heat exchanger located in fluidised bed. The required heat is provided by combustion of a share of the syngas. To achieve a good efficiency the amount of gas for the combustion has to be minimised. Therefore, the air should be pre-heated as far as possible.
Biomass Steam Syn-gas
........ Air
--
*
Flue Gas
Figure 3 Heat Supply of the Gasifier The concept of the heat supply is shown in Figure 3. The syngas is combusted with the pre-heated air. The hot flue gas is cooled down to a sufficient inlet temperature for the heat exchanger by admixing some of the flue gas from the heat exchanger outlet. To compensate pressure losses a compressor I blower is necessary. The outlet temperature of the heat exchanger in the gasifier is given by the gasification temperature and an assumed pinch of 50 K. The feasible inlet temperature depends on the high temperature properties of the chosen material. The temperature should be as high as possible to reduce the mass flow through the heat exchanger and to reduce its sue. Another way to control the temperature could be an appropriate excess of air but this would lead to a reduced availability of the heat exchanger caused by higher material stress due to the oxygen content in the gas. The heat of the leaving flue gas is used in a heat recovery system where among other streams the air for the combustion is pre-heated. The heat exchanger in the fluidised bed consists of several hairpin tubes, whch are vertically immersed into the fluidised bed (Figure 4a). This design was selected to keep the material stress on the surface of the heat exchanger small (Figure4b). To increase the availability the ends the of hairpin pipes should be extra shielded by additionally welded metal plates. The pipes are concentrically distributed over the whole cross section of the fluidised bed to ensure an uniform heat transfer in radial direction (Figure 4c). To enable an uniform heat transfer in axial direction the length of the heat exchanger pipes should comply with the height of the bed. However, in a small zone at the bottom of the bed no pipes can be placed due to highly erosive impact of the penetration jets of the fluidisation agent entering the gasifier through a multi-orifice plate distributor. 504
Biomass
Syngas
pq
I
0000 (00000)
v .....
Figure 4 Heat Exchanger Design The design and the geometrical dimensions of the fluidised bed gasifier result from several boundary conditions: (1) The desired ratio of height to diameter of the fluidised bed. (2) A sufficient heat transfer must be ensured, i.e. the heat exchanger agent must fit into the fluidised bed. (3) The ratio biomass char to bed material has to be a certain range. According to the bench-scale reactor at Royal Institute of Technology in Sweden a ratio of height to diameter of the fluidised bed of 1 is chosen. The ratio of biomass char to bed material is between 10-15 % talung into account the experiments. Higher ratios could lead to bed-disturbances. These boundary conditions lead to the design described above and to a calculated fluidised bed diameter of approx. 1.5 m. The minimum steam mass flow to ensure a stable fluidisation results from the bed diameter, the gasification pressure and the minimum fluidisation velocity, whch depends mainly on the particle size. A large diameter, a high pressure and a large particle size leads to an elevated mass flow. The estimated steam mass flow for fluidisation always has to be equal or larger than the required steam mass flow for the gasification process. The particle size does not affect only the steam mass flow but also the heat transfer between the heat exchanger and the fluidised bed. The heat transfer against particle size shows a minimum between of 1.5 - 2.0 mm in this case. At smaller particle sizes the heat transfer increases quickly. DESIGN PRINCIPLE A concept for a CHP unit with an electrical power output of about 1 MW was set-up. It is thermodynamically optimised taking into account the experiences from the
505
gasification tests. The following boundary and operational conditions lead to a good efficiency. (1) The gasification temperature and the steam mass flow should be as low as possible. The gasification tests have shown that feasible operational conditions were achieved at temperatures of about 850°C and higher. The necessary steam mass flow results from the fluidisation conditions. (2) The gasification pressure should be as small as possible. Hence, the gasification pressure is determined by the required inlet pressure of the syngas at the gas engine and the pressure losses caused by heat exchangers and piping. (3) The gasifier has to be supplied with heat for whch syngas is combusted. To achieve a good overall efficiency the amount of syngas for combustion has to be minimised. Therefore, it is necessary to pre-heat all streams leading to the gasifier and the combustor as far as possible. Thus, a complex heat recovery system is inevitable. Based on preliminary results the electrical efficiency of the CHP application is calculated to 25 - 30%. The gasifier has an efficiency of 70% considering the amount of the combusted syngas for the heat supply. The heat that could not be recovered for the gasification process is used for generation of district heat. This concept is not optimised with regard to economic aspects. BRIEF DESCRIPTION OF THE SYSTEM As shown in Figure 5 the biomass is gasified with steam at 850°C. Discharged bed material and char dust are separated in a cyclone from the product stream and recycled to the fluidised bed. The gasification tests at LDU showed that the tar content in the fuel gas was too high for direct feeding into the gas engine. Therefore, a tar cracking device is necessary. For the same reasons the ammonia content in the syngas must be reduced, too. Both can be done applying a catalytic reactor. Subsequently, the syngas is cooled in several steps. Afterwards, the gas stream is divided, whereby a small share (approx. 20%) is reheated to a temperature of approx. 840°C and fed to the combustor. The remaining 80% of the syngas are cooled down to the required engine inlet temperature of approx. 7OoC and finally combusted in the gas engine. In the combustor for the heat supply the reheated syngas is burned with the pre-heated air (at approx. 850OC). The hot flue gas and recycled cold flue gas are mixed to achieve the desired inlet temperature for the heat exchanger (approx. 1l0OOC). Leaving the heat exchanger with a temperature of approx. 900°C the flue gas is used to pre-heat the fluidising agent, the air (approx. 85OOC) and syngas for combustion. The exhaust gas of the gas engine (450 - 5OOOC) is additionally integrated in the heat recovery process for preheating of the different streams and for steam generation.
506
Figure 5 Concept of the CHI’ Unit DESIGN OF THE GAS ENGINE The research at Jenbacher AG during the past years has shown that a sufficient performance output and an appropriate cost benefit relation can be realised by a turbo charged gas-otto-engine. For the combustion of the syngas a gas engine designed for natural gas was adapted. The fuel feeding to the cylinder had to be changed to obtain a homogeneous mixture formation that was important to reduce NOx emissions and to prevent knocking. The Technical University Graz simulated the fuel mixing process inside the cylinder with a 3-dimensional CFD tool. The grid model for the CFD tool is shown in Figure 6. Different concepts have been investigated. It was chosen to operate with direct ignited combustion. Thereby, the ignition of the airhyngas mixture was carried out by a spark plug installed directly into the cylinder head. The syngas was a h x e d to air in the d e t pipes. For the tests a modified series cylinder head was designed and machined (Figure 7). This type of cylinder head had all design features of the concept developed by the Technical University of Graz. The tests with the research engine are ongoing.
SUMMARY The gasification tests have shown that stable operational condition were achieved above 850°C at a pressure of 0.5 m a . It was observed that gasification temperature had an important influence on the tar yield of the syngas.
507
The use of olvine sand as bed material was possible without any restriction due to temperature and pressure. Magnesite was limited to high temperature and the use of gasification agent air. At low temperatures around 400-500°C it showed tendencies to build hydroxides whch led to agglomeration of particles. dolomite had similar problems at a pressure of 1 MPa with the gasification agent steam in combination with temperatures above 700°C. Hydrogen is a strong catalyst poison over 650°C for dolomite. The gasification of biomass with steam offers advantages regarding the gas composition and the heating value. However, a high expenditure for the heat supply of the gasifier and the heat recovery system are necessary. The required heat for gasification rises with increasing gasification temperature. In the presented concept the heat supply is done by combustion of a share of the syngas. A complex heat recovery system for heat integration is necessary to achieve good efficiencies, for which high temperature heat exchangers are required. The allothermal gasification with steam has particularly advantages if hot temperature heat sources fiom other processes can be integrated. The district heat generation is done at a high temperature level. If the emphasis of such a plant would be the electricity generation than the district heat generation could be replaced by a small steam cycle.
Figure 6 CFD Grid Model
Figure 7 Cylinder Head
ACKNOWLEDGEMENT
The concept of a combined heat and power station presented is based on results of the EU funded project JOR3-CT98-0291. LITERATURE 1. TAB-Arbeitsbericht Nr. 49, Monitoring "Nachwachsende Rohstoffe", Vergusung
und Pyrolyse von Biomassen.
508
Biomass Power Generation: Sugar Cane Bagasse and Trash L. Waldheim', M. Monis' and M. Regis Lima Verde Leal2 1 TPS Termiska Processer AB, 61 1 82 Nykoping, Sweden Copersucar Technology Center, C.P. 162 - Piracicaba SP-Brazil
ABSTRQCT A huge potential for power generation from waste fuels exists within the sugar cane industry. 1 200 million tonnes of sugar cane is harvested annually, whch corresponds to a worldwide electricity production potential of 40 000 MW or 300 T W a n n u m in the eighty countries where sugar cane is grown on a significant basis. The main objective of Project BRA/96/G31 - Biomass Power Generation: Sugar Cane Bagasse and Trash, is to evaluate and develop the technology required in the complete fuel-to-electricitychain; starting with cultivation and recovery of sugar cane by-product fuels to electric power generation with advanced systems (e.g. BIG-GT, Biomass Integrated Gasification - Gas Turbine) integrated with the sugar mill. Although the main activities within the project are focused on Brazil, it is an international effort aimed towards the optimisation of energy production from sugar cane biomass. In addition to the global potential of sugar cane biomass to increase the power output from industry, the emitted greenhouse gases will be reduced by millions of tonnes; as a result of the substitution of fossil fuels, increased generation efficiency and stopping the current practice of burning of cane fields prior to harvesting. The Global Environment Facility (GEF), through UNDP, has co-funded the project and, for the activities starting in the year 2000, the European Union and the Swedish national energy adrmnistration also provide financial support. The project work plan covers the following five topics:
(1) Agronomic routes to green cane harvesting with trash recovery. (2) Availability and quality of sugar cane trash. (3) Use of bagasse (i.e. the dry pulp remaining after the juice has been extracted from sugar cane) and cane trash as fuels in a BIG-GT process. (4) Integration of a BIG-GT system in a typical sugar mill. (5) Environmental impact. Various cane harvesting and trash recovery procedures have been developed and evaluated, not only in relation to the economy of the recovery operation, but also in relation to the quality of the trash as well as weed control aspects. The potential for reduction in the emission of greenhouse gases by changing from traditional harvesting 509
to trash recovery has been estimated. Pilot plant gasification tests have been performed on bagasse, and similar tests on trash are planned to start this year. On the basis of the results obtained so far, initial studies on the integration of a BIG-GT plant in a sugar mill and optimisation of the overall mill facilities have been made. Thls paper presents the status and future plans of the project and hlghlights some results and conclusions reached.
INTRODUCTION Development in the sugar cane industry is driven by the wish to convert the fueI value of the biomass fuels produced (bagasse and the almost equal amount of cane trash, comprising tops and leaves separated at harvest) to a high quality energy carrier such as electric power. However, driving forces also include establishing a more sustainable, environmentally friendly, large-scale, agronomic production chain and improving social conditions in the areas of cultivation. Recovery of cane trash implies a change from traditional harvesting methods; which normally consists of setting huge areas of cane fields ablaze prior to the harvest, and then employing seasonal workers to cut the cane stalks manually. To recover the trash, a new so-called “green mechanical harvesting” scheme will have to be introduced. By recovering the trash in this manner, the production of local air pollutants, as well as hydrocarbons contributing to adverse climatic change, from the fires will be avoided, and working conditions of the labourers will be improved. All-year-round work will have a social impact as the seasonal labour that normally moves around the large agricultural areas with the harvest can be located in one region. To realise h s vast agro-fuel potential, research and development of the entire system fiom the fuel supply chain, the thermal process, its integration in the sugar mill’s combined heat and power (CHP) system to export to the grid must be carried out and demonstrated at 111 scale. An initiative in this direction is Project BR4/96/G31 - Biomass Power Generation: Sugar Cane Bagasse and Trash. In July 1997, Copersucar and the UNDP signed a contract covering the activities of this project. The work is coordinated by the Brazilian Ministry of Science and Technology. The duration of the project was estimated to be 30months with a budget of US$7.4 million; split roughly equal between Copersucar and the GEF (through UNDP). During year 2000, financial support for an extension of the work was received from the European Union’s ENERGIE program and the Swedish National Energy Administration (STEM).
PROJECT OBJECTIVE The main objective of the project is to evaluate and develop the technology required in the complete fuel-to-electricity chain, starting with cultivation and recovery of sugar cane by-product fuels to electric power generation with advanced systems (e.g. BIG-GT, Biomass Integrated Gasification - Gas Turbine) integrated with the sugar mill. 510
In the area of sugar cane harvesting, the objectives are to lmprove the equipment and procedures for recovering the cane trash, whilst also studying side effects such as decrease in the use of herbicides. A second objective is to make realistic assessments of the cane recovery potential, fuel quality and the associated cost. In the area of gasification, the objective of the project is to test and improve the performance and availability of an atmospheric-pressure circulating fluidised bed (CFB) gasification system operating on agrofuels. The scientific and technical objectives for the gasification tests are; to establish, by testing, suitable operating conditions for cane trash he1 in order to maximise the performance and availability of a large-scale demonstration plant. Although the main activities of the project will take place in Brazil, the project is an international effort aimed towards the optimisation of the energy potential of sugar cane worldwide. Copersucar Technology Center (CTC) of Brazil will work mainly on sugar cane cultivation and harvesting as well as on trash recovery procedures. CTC is the technology branch of Copersucar, a cooperative of 36 sugadethanol mills in Brazil producing three million tonnes of sugar, 3.2 million cubic metres of ethanol and crushing 65 million tonnes of cane per season. Most of Copersucar’s affiliated mills are in the State of Sgo Paulo, Brazil. CTC will also take the lead in the integration studies of the BIG-GT plant with the sugar mill. The gasification development will be the responsibility of TPS of Sweden. This work is an extension of TPS’s work in the GEF-supported Brazilian Wood project (WBP project) [l]. TPS will also participate in the study of the integration of the BIG-GT plant with a typical sugar mill. TPS is a Swedish R&D company that has developed a proprietary gasification process that can be used in conjunction with a gas turbine, this process will be demonstrated on wood fuel in the 8 MW ARBRE plant in the UK [2]. A similar project of 32 MW is planned for Mucuri, Bahia, Brazil. Within the part of the project supported by the EU, Danish experiences fiom straw handling and feeding will be applied to sugar cane trash by Thomas Koch Engineering A/S. In particular, these experiences include feeding to a large scale CFB reactor by a feeder which incorporates a proprietary plug formation arrangement to permit the plug properties to be controlled. Tests with gasifier feed systems will be made at a scale of 1 tonnehr. The objective of this testing, which is expected to start in the latter part of 2000, is to compare the feeding characteristics of sugar cane fuels with that of straw to permit the design of feeder systems with capacities up to 20 tonneshr.
WORK PLAN
The project work plan covered the following five main elements: (1) Evaluation of agronomic routes to green cane harvesting with trash recovery. This part of the project focuses on green harvesting and trash recovery routes; including trash recovery in a dry cleaning station or recovery by baling in the field, bale processing at the mill, and the costs associated with various trash recovery routes. This part of the project will generate data from testing of the relevant machinery. (2) Evaluation of sugar cane bagasse and trash availability and quality. 51 I
This part focuses mainly on the trash, as the availability and quality of bagasse is already well understood. The studies will establish the theoretical quantity of trash available from normal and developed cane varieties, the amount that can realistically be recovered (taking into account also the impact of leaving trash in the field) and the fuel quality aspects of the trash. (3) Bagasse and trash atmospheric pressure fluidised bed gasification tests. In this part of the project, tests will be performed in a 2 M W air-blown atmospheric pressure CFB gasification pilot plant at TPS. The tests are required to establish suitable operating conditions for the BIG-GT system, to generate data for modelling p q o s e s , and to establish contaminant (i.e. tar, alkalis, etc.) content in the gas. This data will be input information for a design study to establish the gasification and gas cleaning parts of the BIG-GT installation to be used in the sugar mill integration study. (4) Integration of the BIG-GT system with a typical mill. The focus of this part of the project is the integration of the BIG-GT installation in the sugar mill. ( 5 ) Identification and evaluation of environmental impact. The environmental impact, in terms of emissions to atmosphere, on soil, on terrestrial biological environment as well as socio-economic impacts will be evaluated.
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CURRENT PRACTICES IN SUGAR CANE HARVESTING AND IN SUGAR MILL ENERGY MANAGEMENT The sugar cane harvesting period, with few exceptions, is five to six months, which is also the seasonal period for the sugar mill. Normal practice, with a few exceptions (e.g. in Cuba), is to set the fields on fire prior to harvesting, and then harvest the sugar cane stalks manually or automatically. This practice leads to the release of a huge amount of fumes and other contaminants detrimental to the local air quality. Also, the tines release unburned hydrocarbons which contribute to the greenhouse effect. The harvested sugar cane is transported to the sugar mill and, during the milling process, sugar is extracted for use or for fUrther processing by fermentation to yield alcohol. As a side product, bagasse, at a moisture content of about 50%, is recovered for use as a fuel. Brazilian sugar mills, similar to mills throughout the world, have process steam requirements of approximately 500 kg of steam per tonne of cane processed. To produce this steam (typically at 22 bad300"C) nearly all bagasse produced (typically 140 kg dry soliddtonne cane processed) is consumed. This steam production is sufficient to produce all the electric and mechanical power required to run the mill and thus, fuel availability and power and thermal energy requirements are balanced. However, because the boilers normally have low efficiency and their steam conditions are conservative, the power to heat ratio is low. The reason for this is that there is little incentive to improve the energy conversion system because bagasse is considered to be a waste material and an excess of this material when entering the off-season will require that it be disposed of at cost. In addition, institutional barriers have not given sufficient incentive to produce more electric energy than required in-house. 512
Beet sugar factories and corn ethanol distilleries are much more efficient in energy use than sugar cane mills because they have to purchase fossil fuels. It is known that the present steam consumption of the average sugar cane mill can be considerably reduced by using the process technology available in beet sugar factories and corn ethanol distilleries. ' I h s improvement in process technology will allow the application of new and more efficient technologies, e.g. more efficient boilers at more favourable steam conditions or the even more efficient BIG-GT plant.
GREEN HARVESTING, STATE-OF-THE-ART The trash recovery part of the project was divided into the following four routes (shown schematically in Fig. 1):
(1) Whole cane harvesting; loading and transporting cane and trash; cleaning and trash recovery at the mill. (2) Whole cane harvesting; cane picked up, chopped and cleaned in the field; transporting clean cane; baling and transporting trash to the mill. (3) Chopped cane harvesting; cane cleaned and loaded during harvesting; transporting clean cane; baling and transporting trash to the mill. (4) Chopped cane harvesting with extractor off; cane and trash loaded during harvesting; transporting cane and trash; cleaning and trash recovery at the mill.
Whole green cane harvesling
I-()-[
RquteA
Cane and trash
1
RoyteB
Roule C
Cane and trash picked-up from the
1
harvester
Route D
harvester
loaded into truck
I
Fig. 1 Selected harvesting routes for unburnt cane.
The cane quality and the trash quantity recovered in practice, as well as the cost for its recovery, will be different for each scheme. The knowledge and experience
513
available at the start of the project was insufficient to have a good idea of the potential of trash yield, or to select any one of these routes as being superior.
TRASH AVAILABILITY FIELD TESTS The first activities of this part of the project were directed at the assessment of cane biomass quantity and quality in the cane field prior to, and after, harvesting. The influence of cane variety, age, cut, soil and climate was also investigated. During the 1997/98 harvesting season, field tests were performed on routes C and D with the combined harvester AustoftA7700, with the trash extractor on and off. During the 1998/99 harvesting season, field tests were made with the machines involved in routes A and B. The potential of sugar cane residues is approximately 14% of the mass of stalks. This means that for each tonne of sugar cane stalks, there is 140 kg of dry residues. This potential is only slightly lower than the bagasse yield itself and therefore the biomass potential can, theoretically, be almost doubled by using green harvesting.
TRASH RECOVER Y Trash Recovery in the Field aser Harvesting In 1991, Copersucar began a project to test some balers to determine their performance. The results of the tests, and the operational problems encountered, indicated that the baling system producing rectangular baler was the best; firstly, because of the higher operational baling capacity; secondly, because it deals better with the trash and pieces of cane; and thirdly, because of better space utilization by the bales of the transportation truck. Difficulties in recovering a large amount of small rectangular bales from the field, and their stacking in the truck, indicated that large rectangular bales should be used.
Trash Recovery and Processing at the Mill Site An important item in both routes A and D is the cane dry cleaning station. A prototype for 250 tonnes of cane per hour was designed and built, and tested in one of Copersucar mills since 1994. Trash separation efficiency at the mill is 55 to 60% when operating on chopped green cane. It is expected that a trash separation efficiency of greater than 70% can be achieved.
AVMLABILITY OF TRASH On the basis of the data from the various tests, the net recovery factors were estimated for the four routes A to D. It should be remembered that these are, in many cases, very preliminary tests and that there is still room for further improvements in the recovery of trash. It is important to remember that whatever the method of trash separation from the cane, a certain amount of vegetal impurity (trash) will remain with the cane and it will be crushed with the cane at the mill. This vegetal impurity should be considered in the industrial process as it will influence the amount of bagasse produced. 514
Considerations for the estimate of future residues availability: All unburned harvesting areas will be mechanized. Unburned sugar cane harvesting: 100% in Sgo Paul0 State and 50% in the rest of the country. The State of Silo Paulo, in accordance with environmental laws, will harvest almost all its cane unburned. Approximately 55% of Brazil’s sugar cane area can be mechanized, and therefore harvested unburned. Potential of material (trash) produced: 140 kg dryhonne cane (average value). Efficiency of the harvester in the separation of trash in the field: 68% with the extractor on and 5.5% with the extractor off (values obtained from test results). Trash separation efficiency at the mill: 55 to 70% (average value obtained from test results with the dry cleaning station. The station was improved and will be retested). Trash recovery efficiency with balers: 56 to 84% (value obtained from baling test results, for the condition of maximum recovery efficiency, with the raking of two rows of trash over one row, in areas of unburned cane harvesting, with the trash left on the field). Table 1 presents the final dry biomass availability from residues from harvesting for the main producing areas of Brazil. It is estimated that the potential of agricultural residues for the sugar cane produced in Brazil is approximately 40 million tonnes, most of it nowadays being burned before harvesting. Considering that, in the future, the majority of the sugar cane production will be harvested without burning and, taking into account the recovery factors of cane trash, there will be approximately 20 million tonnes biomass available, without considering the bagasse which amounts to an additional 40 million tonnes. There is also potential for another 10 to 15 million tonnes if mechanised harvesting is adopted completely in the areas where only partial use of th~smethod was assumed in the table. To put this number into perspective, the sugar cane industry can realistically provide, as a side product, biomass fuel quantities comparable in magnitude (i.e. equivalent to 20 to 25 million toe) to 50% of the biomass energy usage in 1995 in the 15 member states of the European Union.
GASIFICATION OF SUGAR CANE BAGASSE AND TRASH The work contracted to TPS as part of this project included characterisation of the gasification properties of bagasse and cane trash fuels, and bench-scale gasification tests. Three pilot plant tests on bagasse pellets were performed during 1998 and 1999. On the basis of the results of these tests, conceptual engineering of bagasse-fuelled BIG-GT power plants, integrated into a sugar mill or as stand-alone units was performed. Success of the tests on bagasse led to an extension of the project. Pilot plant tests on loose trash material will be performed during 2000 and 2001. Support for this activity, and also for the WBP project, has been received from both the EU and Statens Energhyndigheten (STEM) in Sweden.
51s
Availability of Dry Biomass Residues from Sugar Cane Harvesting Table 1 (green leaves, dry leaves and tops) in Brazil, compared to Bagasse Availability (million tonnes). Green Real availability Availability Crushed Potential of cane * dry residues harvest (%) of dry residues of dry bagasse 181.5 25.4 100 13-15 25 50 16-17 35 35.0 249.7 Center South ** 50 2 7 7.2 51.9 North-Northeast Area Siio Paulo
Brazil 301.6 42.2 * source: UNICA - 1997/98 harvesting- season ** including S b Paulo.
80
18-19
42
DESCRIPTION OF THE PILOT PLANT Fig. 2 is a schematic flowsheet of the TPS 2MW atmospheric pressure CFB gasification pilot plant. A
GAS
I
FU
Fig. 2
Schematic flowsheet of TPS’s 2 MW CFB gasification pilot plant.
The fuel is fed from a hopper to a weigh belt conveyor which measures the feed flowrate, it then passes through a rotary valve system equipped with sealing air and into the gasifier through a screw feeder. The screw feeder controls the fuel feed flowrate. The gasifier is of CFB type. Medium-sized fuel particles and bed material elutriated from the gasifier are captured in the solids separators whch are placed at the exit of the gasifier, and recycled to the bottom section of the gasifier. At the
516
bottom of the gasifier, a sparger type distributor provides primary air to the fluidised bed. An ash drain is located below this distributor. More details of the gasifier design can be found in Ref. 3. The gas leaving the gasifier’s secondary solids separator enters the bottom section of the “tar cracker”. The cracker is of CFB type, it operates in a similar manner to that of the gasifier. The gas leaving the cracker’s secondary solids separator passes through heat exchangers before it enters a “cold cyclone”. The gas leaving the cold cyclone can be flared. Downstream of the cold cyclone, the gas passes through a filter and a wet scrubber.
TEST PROGRAM AND RESULTS The test program included three tests, a first so-called “shake-down’’ test to establish typical operating conditions, followed by two tests to yield steady-state results. Throughout the tests, the pelletised bagasse was found to have excellent fuel feeding properties which resulted in a high and even fuel feeding during all three tests, thus providing good plant availability. Availability of the gasification system was logged at between 9 1 to 99% during the three weeks duration of the tests. In the first test on bagasse, the gasifier was first operated at low temperature. The temperature was then gradually increased until agglomeration was provoked. The purpose of this step-up in temperature was to validate the projections made from bench-scale tests on the safe operating window to avoid ash related problems. The results were in line with projections, and in the following tests, the operating temperature was kept below a threshold temperature, and no agglomeration occurred. The high chemical reactivity of the organic part of the bagasse results in a high carbon conversion to gas, above 95%.The carbon content of the bottom ash was low. As the gas cleaning is achieved in a separate stage, the operating conditions of the gas cleaning is decoupled from the gasifier’s operating conditions. A reasonably low tar content of the product gas was achieved at temperatures close to the upper limit, whilst still not interfering with gasifier operation. The composition and heating value of the gas generated (Table 2) was typical for the pilot plant operating on a dry biomass fuel. Because of the high heat losses in the pilot plant, the lower heating value of the gas is lower than in the bench-scale tests, where electrical heaters are used for compensation, or in a full-scale plant where heat losses are relatively lower. The conversion of nitrogen to ammonia is high, 60 to 90%. To reduce NOx emissions from a gas turbine plant to acceptable levels, the ammonia is removed from the product gas in a wet scrubber located immediately upstream of the gas compressor. HCN formation is low, much less than 1%. The main objective of the tests, to show that sugar cane bagasse can be used as feedstock in the gasification process, was achieved. The other objectives; to find a stable operating regime and to validate the data and parameters used for modelling and scale-up, were also achieved.
517
Table 2
Composition of Gas Generated in the TPS Pilot Plant. Gas component
% volume
10.0 56.4 12.7 3.7 16.7 0.5 LHV
4.3 MJ/Nm3
BIG-GT INTEGRATION WITH TYPICAL MILL STEAM ECONOMY IN THE SUGAR MILLS
Brazilian sugar mills, similar to mills throughout the world, have process steam consumption of approximately 500 kg of steam per tonne of cane processed. By generating steam at 22 bar/30O0C, sufficient electric and mechanical power (using back-pressure turbines) to run the plant is aclueved and nearly all bagasse produced is consumed. Thus, fuel availability, and power and thermal energy requirements are balanced. Typical Mill
In this project, “Typical Mill” was used as a reference. The Typical Mill has the following operating conditions: Daily milling Annual cane processing Milling rate Pol, % cane (sucrose content) Fiber, % cane Annual bagasse production, 50% moisture Annual trash availability, 15% moisture Sugar production Alcohol production Process steam conditions Steam consumption Milling season Off-season
7 000 tonnes 1 300 000 tomes 292 tonnesh 14.1% 13.8% 345 800 tonnes 120 000 tomes 400 tonnedday 353 000 litredday 2.5 bar, saturated 500 kg stedtonne cane 4 457 hours 3 164 hours
The existing boilers in the sugar mill are two 54 tonnesh bagasse-fuelled boilers, each capable of generating 100 t o m e s h at 22 bar, 300°C, and, apart from the mechanical drives in the mill, a 3 300 kW turbogenerator.
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Economising on Steam The potential for decreasing steam usage in the sugar mill, and hence also to reduce the cost, was studied. Two levels of savings; a decrease in steam consumption from the conventional 500 kg/tonne of cane down to 340 and 280 kg/tonne of cane were considered. The total investment for the 340 kghonne of cane case was estimated to be US$4.5 million and the process steam consumption for an ethanol production of 50% hydrated and 50% anhydrous ethanol was 341 kg steadtonne cane. The total investment for the 280 kg/tonne of cane case was estimated to be US$7 million and the process steam consumption for an ethanol production of 50% hydrated and 50% anhydrous ethanol was 287 kg steadtonne cane. The review made by CTC in 1999 indicated that the moisture content of bagasse during the season and also off-season was consistently 50%. However, it was concluded that the moisture content of the trash during the harvest season was 15%, and by re-watering it was increased to about 50% during the off-season. This means that during the season, operation on trash without drymg is feasible. A dryer has to used in the BIG-GT plant and t h s equipment should be capable of handling both bagasse and trash, or a mixture of both, in the range of moisture contents outlined above. CTC opted for a flash type dryer, which has been successfully used in several mills to dry bagasse using boiler flue gas as drylng agent. This type of dryer is believed to have a low investment cost and it will be easier to integrate with the BIG-GT plant, although a separate hot gas generator is required to supplement the flue gases. Use of Bagasse and Cane Trash CTC analysed the alternatives of BIG-GT - mill integration based on the preliminary BIG-GT data provided by TPS. To organize this evaluation, some basic assumptions were made:
(1) BIG-GT plant: based on the General Electric (GE) LM 2500 gas turbine. (2) HRSG pressure/temperature (bad'C): 82/480, 22/300 and 2.5haturated. (3) Mill process steam consumption levels (kg steadtonne cane): 500,340 and 280. Two operating modes were considered:
(1) Independent thermal power plant. (2) Cogeneration, full or partial. BIG-GT System Process Description, Input Data and Assumptions The fuel is gasified in an air-blown CFB gasifier operating at slightly above atmospheric pressure and at about 850°C. Bottom ash is continuously withdrawn from the gasifier. The raw gas is catalytically cleaned from tar by hot gas cleaning in the "tar cracker". The hot fuel gas is cooled to enable removal of dust, nitrogen-containing compounds, chlorides, alkalis, etc. in the secondary gas cleaning stage. A baghouse filter operating at about 180°C efficiently removes the fly ash dust. 519
A wet scrubber, in which a solution of ammonium sulphate circulates, cools the gas and causes most of the water vapour to condense, whilst at the same time it absorbs any remaining contaminants in the gas, primarily ammonia. The condensate is cleaned by physical and biological treatment. The cleaned fuel gas is pressurised in a gas compressor prior to enterhg the gas turbine combustion chamber. The performance of the GE LM 2500 gas turbine was calculated using the Simulationmodel developed by TPS in the WBP project. Heat to the steam system is provided by the flue gas from the gas turbine to the heat recovery steam generator (HRSG) and by the cooling of the fuel gas. The flue gas from the HRSG can be used to dry the fuel in an integrated dryer. A steam condition of 60 barI5OO"C was chosen for the independent operation mode. In cogeneration mode, three steam conditions were specified by CTC: high-pressure (80 bar/48O0C), medium-pressure (22 bar/300°C) and low-pressure steam (2.5 bar/saturated). The independent mode is essentially an adaptation of the WBP project combined-cycle plant concept operating with bagasse and trash. In independent mode, for one BIG-GT module based on the LM 2500 gas turbine and a HRSG operating at 60 bar/500°C, a net electric output of 33 W e , and a net electric efficiency of 40% can be achieved for bagasse; but with cane trash, the efficiency on an LHV basis is 34 to 37%, the lowest value associated with a dry fuel during the off-season period when low level heat recovery in the dryer is no longer appropriate. Cogeneration mode, either full or partial, requires a careful evaluation of the alternatives to find the best technicaYeconomic option. In cogeneration mode inside a sugar mill, two cases, using all three steam pressure alternatives, were studied, namely integrated drymg or external drymg. The net electric output of one BIG-GT module is 16 MWe with a total efficiency of 78% and a power to heat ratio of 0.37 to 0.38 if a fuel dryer is integrated in the BIG-GT plant using HRSG exhaust flue gas. if the drying is performed outside the BIG-GT unit, the power to heat ratio decreases to 0.33, as more heat is recovered in the HRSG. The integration of the dryer in this case must be evaluated for the sugar mill concerned, as integration of the dryer in the BIG-GT unit shifts fuel usage from an external dryer to the mill boiler to meet the overall steam demand. The steam consumption in the sugar mill was originally 500 kg/tonne cane milled. It was found that the reference system, which was in balance in terms of steam consumption and drive requirements in relation to the bagasse produced, could not be improved by utilising a BIG-GT system because of the lower steam production as a result of using the gas turbine, although the electric energy would increase. This would otherwise set the plant off balance. Thus, it was concluded that the introduction of BIG-GT technology, and its integration into the sugar mill would catalyse steam economy investments in the mill itself. After the first screening step, based on preliminary heat balances and engineering judgement, six systems were selected. In the calculations, it was assumed to have year-round production in the BIG-GT plant, cane trash would be mixed with bagasse and any steam in excess of the mill need would be used in a new condensing backpressure turbogenerator at the appropriate pressure level. The results are shown in Table 3. The results show that the power production is similar in all cases, except when two BIG-GT modules in parallel are used, in which case it is almost doubled. The first 520
step of steam economy (i.e. reducing the steam consumption of the mill from 500 kg steadtonne cane to 340 kg steadtonne cane) will allow an increase from no exported power to 170 kWtonne cane. Apart from steam savings, thls would not require any change in the mill steam system, as increases in steam pressure or new boilers do not increase the power export to any significant amount. In the second step of steam economy (i.e. reducing linther the steam consumption to 280 kg steadtonne cane) and when a sufficiently large BIG-GT unit can reliably substitute the mill boilers completely for steam production, a M e r increase in exported power to 290 k W t o n n e cane could be achieved. In this case, the choice was to use the single module system with the lower pressure, as the mill boilers will not require replacement, and the HRSG and the new turbogenerator can be designed for lower pressures, giving lower total investment cost. Table 3
Condition
Boiler+HRSG
- 20T340 Boiler+HRSG Boiler+HRSG Boiler+HRSG 2 BIG-GTs HRSG only New boiler + HRSG New boiler + HRSG
Results from Analysis of a Typical Mill.
Steam HRSG Mill Exported Exported consumption steam steam power, power, (kg/tonne condition condition season off-season cane) (bar1"C) (bar1"C) (MW) (MW)
Specific power export (kW tonne cane)
340
221300
221300
27.9
27.9
163
280 280 340
221300 821480 821480
221300 221300 221300
27.9 27.5 28.1
27.9 29.2 29.3
163 165 168
280
821480
221300
43.3
58.5
29 1
340
821480
821480
29.3
29.3
172
280
821480
821480
29.3
29.3
171
Even if the case of using two BIG-GT systems had the best overall data, the mill would, during the season, be reliant on the new technology, and one system alone would not generate steam enough to maintain the needs of the mill. Therefore, if during the season one unit would be out of operation for whatever reason it would have a large mpact on the mill if a mill boiler was not on standby all the time. A standby boiler would reduce the benefit of this system. The single module system 20T340 therefore seems good from the point of view of constructing a first demonstration plant. The case where two BIG-GT units are used is of interest to look at in future calculations. This system gives better fuel utilisation and improved net power export, it therefore concerns a mature technology of long-term interest.
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ENVIRONMENTAL IMPACT OF GREEN HARVESTING AND IMPROVED ENERGY CONVERSION In March 1998, the evaluation of the atmospheric impact was concluded. Three main aspects were analysed: (1) Energy balance/C02 net emission.
(2) Methane and other greenhouse gases emissions. (3) Particulate emission. The results show the extraordinary contribution that the new technology may bring to reducing greenhouse gas and particulate emissions. The main factors are the new amount of biomass available for power production, the higher efficiencies in conversion and, to a lesser extent, the avoided emissions as a result of less sugar cane burning. The final results are shown in Table 4. Table 4
Reduction in Emissions of Greenhouse Gases (methane, CO, NOx) due to the Partial Harvesting of Non-burnt Cane.
Emitted gas Carbon monoxide Methane Methane (as COz equivalent)
co
NOx particulates
Difference: Future - Today (kghonne cane) 85 - 137
- 0.03 (wind tunnel) - 1.2 (IPCC)
Brazil: 300 x lo6 tonnes canelyear (million tonnes) 25 -40 0.01 - 0.4
0.7 (wind tunnel) - 30 (IPCC)
0.2 - 9
- 1.8 (wind tunnel) - 4.1 (IPCC)
0.5 - 1.2 0.03 - 0.1 0.4 - 0.6
- 0.1 (wind tunnel) - 0.3 (IPCC) -1.3 to 2.7
CONCLUSIONS For each tonne of sugar cane stalks, 140 kg of dry residues can be recovered. This potential is only slightly lower than the bagasse yield itself and therefore the biomass potential can, theoretically, be almost doubled by using green harvesting. Trash separation efficiency at the mill is 55 to 60% when operating on chopped green cane. It is expected that a trash separation efficiency of greater than 70% can be achieved. It is estimated that the potential of agricultural residues for the sugar cane produced in Brazil is approximately 40 million tonnes. Considering that, in the future, the majority of the sugar cane production will be harvested without buming and, taking into account the recovery factors of cane trash, there will be approximately 20 million tonnes biomass available, without considering the bagasse which amounts to an additional 40million tonnes. There is also potential for another 10 to 15 million tonnes if mechanised harvesting is adopted completely in the areas where only partial use of this method was assumed.
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In the tests in the gasification pilot plant, the excellent feeding properties of bagasse resulted in a high and even fie1 feeding, providing plant availabilities ranging from 91% in the first test, to 96 to 99%. A high carbon conversion to gas, above 95%, was achieved. The carbon content of the bottom ash was low. The main objective of the tests, to show that sugar cane bagasse can be used as feedstock in the gasification process, was achieved. The other objectives; to find a stable operating regime and to validate the data and parameters used for modelling and scale-up, were also acheved. In the BIG-GT - mill integration study, independent thermal power plant and cogeneration, full or partial, modes were evaluated. It was found that the reference system, which was in balance in terms of steam consumption and drive requirements in relation to the bagasse produced, could not be improved by utilising a BIG-GT system because of the lower steam production as a result of using the gas turbine. This would set the plant off balance. Thus, it was concluded that the introduction of BIG-GT technology, and its integration into the sugar mill would catalyse steam economy investments in the mill itself. In independent mode, for one BIG-GT module based on the GELM2500 gas turbine and a HRSG operating at 60 bar/500°C, a net electric output of 33 W e , and a net electric efficiency of 40% can be achieved for bagasse, but with cane trash, the efficiency on an LHV basis would be 34 to 37%. Of the cases investigated, the results show that the power production is similar in all cases, except when two BIG-GT modules in parallel are used, in which case it is almost doubled. The first step of steam economy (i.e. reducing the steam consumption of the mill from 500 kg stedtonne cane to 340 kg steadtonne cane) will allow an increase from no exported power to 170 kWh/tonne cane. Apart from steam savings, this would not require any change in the mill steam system, as increases in steam pressure or new boilers do not increase the power export to any significant amount. In the second step of steam economy (i.e. reducing further the steam consumption to 280 kg steadtonne cane) and when a sufficiently large BIG-GT unit can reliably substitute the mill boilers completely for steam production, a further increase in exported power to 290 kWh/tonne cane could be achieved. The environmental impact investigation showed the extraordinary contribution that the new technology can bring to reducing greenhouse gas and particulate emissions. The main factors are the new amount of biomass available for power production, the hgher efficiencies in conversion and, to a lesser extent, the avoided emissions as a result of less sugar cane burning.
REFERENCES 1 Waldheim L. & Carpentieri E. (1998) Update on the Progress of the Brazilian Wood BIG-GT Demonstration Project. ASME Turbo Expo 98, Special Biomass Section. Stockholm, Sweden. 2 - 5 June. 2 Rensfelt E. & Everard D. (1998) Update on Project ARBRE: Wood Gasification Plant Utilising Short Rotation Coppice and Forestry Residues. Seminar on Power Production from Biomass 111. Espoo, Finland. 14 - 15 September. 3 Rensfelt E. (1997) Atmospheric Gasification - The Grkve Plant and Beyond. International Conference on Gasification and Pyrolysis of Biomass. Stuttgart, Germany. 9 - 1 1 April.
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Ammonia formation and NO, emissions, with various biomass and waste fuels at the Vamamo 18 MWthIGCC plant Barbara Goldschmidt' , Nader Padban2, Michael Cannon3, Greg Kelsal14, Magnus Neergaard', Krister Sthl6, Ingemar Odenbrand7 Sycon Energikonsult AB, SE-205 09 Malmo, Sweden Dept. of Chemical Engineering II, Chemical Center, Lund University, P. 0.Box 124, SE-221 00 Lund, Sweden ALSTOM Power UK Ltd.,P. 0.Box 1, Lincoln LN2 5DJ UK ALSTOM Power, Technology Centre, Cambridge Rd, Whetstone, Leicester, LE8 4LH) UK Sycon Energikonsult AB) SE-205 09 Malmo) Sweden Sydkraj AB, SE-205 09 Malmo, Sweden Dept. of Chemical Engineering II, Chemical Center, Lund University, P. 0.Box 124, SE-221 00 Lund, Sweden
ABSTRACT The long-term operation experience from the Varnamo IGCC plant includes extensive information on flue gas emissions. A range of biomass fuels, including forest residue, willow, straw and RDF, have been gasified and burnt during a demonstration period of six years. Willow and straw have been used both as sole fuels and as mixtures with wood fuel. RDF has been used in mixtures with wood fuel. Among other things, the nitrogen chemistry during gasification and combustion has been studied during the demonstration period. The conversion from fuel nitrogen to ammonia in the gasifier was found to be around 60-65% for all fuels. The conversion was independent of fuel type and fuel nitrogen content. No dependency on operational parameters of the gasifier was found either. The conversion from ammonia to nitrogen oxides in the gas turbine was found to be dependent on concentration. At higher ammonia levels the conversion ratio was lower. The conversion ratio was 60-70% for the standard fuel mixture of bark and forest residue, for which the plant was designed. Fuels with hlgher nitrogen content had lower conversion ratios, the conversion ratio for wheat straw being around 30%. The total conversion from fuel nitrogen to nitrogen oxides was around 40% for mixed bark and forest residue. For wheat straw it was 2025%.
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INTRODUCTION The biomass fuelled IGCC demonstration plant at V h a m o consists of a 18 MWm pressurised circulating fluidised bed gasifier, a 4,2 MW, gas turbine and a 1,8 MW, steam turbine. The low heat value gas produced in the gasifier is cooled in a gas cooler and cleaned in a candle filter at a temperature of 350-4OO0C before it is combusted in the gas turbine. The flue gas from the gas turbine is used for the production of steam in the heat recovery steam generator. The steam is superheated, together with steam from the gas cooler, and supplied to the steam turbine. Besides electricity, 9 MWm heat for district heating purposes is produced. An extensive demonstration and development programme has been carried out at the Vamamo plant during 1996-2000 (I). The operation has been very successful, and the overall experience from the demonstration programme is that the IGCC technology with pressurised gasification of biomass fuels is a highly feasible technology for efficient conversion of solid biomass fuels to electricity. The best proof of the successfid operation is the achieved number of operating hours. The Vamamo plant has been in operation for 8500 hours (in gasification mode). The total gas turbine operation time on 100% low heat value gas from the gasifier is 3600 hours. During the demonstration period, experience has been gained in a number of areas. Areas of interest include general technical development, fuel flexibility, environmental performance and the evaluation of costs of fbture plants. The investigation of environmental performance of the plant has focussed, among other things, on the nitrogen chemistry during gasification and combustion. Results from this investigation are presented below.
EXPERIMENTAL Fuel samples were analysed by independent laboratories using standard methods (C, H and N using Leco elementary analyser, heating value according to IS0 1928-1976). The ammonia content of the low heat value gas from the gasifier was analysed after the hot gas filter at the operating temperature of around 35OoC, using a continuous online instrument based on lnfrared spectroscopy (Opsis). Ammonia was also analysed by Lund University, using a Fourier transform infrared spectrometer. Additionally hydrogen cyanide, as well as ammonia, was sampled and analysed by VTT, using methods described elsewhere ('I.The three different methods of ammonia analysis showed good agreement. Nitrogen oxides in the flue gas were analysed continuously by a multi-component analyser based on d a r e d spectroscopy (Mekos).
RESULTS GASIFICATION
A number of biomass fuels and refuse derived fuels (RDF) have been gasified and subsequently combusted in the Vamamo plant. The fuels cover a broad range of nitrogen content, from 0,1% in sawdust to 1% in wheat straw.
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midue
1596RDFh bark
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Fig. 1 Typical nitrogen content in tested biomass fuels (including mixtures with RDF). The nitrogen content of the design fuel of the Vamamo plant, a mixture of bark and forest residue, normally lies within the range 0,2-0,4%, depending on the relative amounts of bark and forest residue. Also the nitrogen content of other fuel mixtures show some variation, especially when fuels with relatively high and low nitrogen contents are used together as in mixtures of RDF and sawdust. The values shown in Figure 1 are typical average values. The nitrogen content of the sawdust and cellulose chips fuels is close to the detection limit of the nitrogen analysis method used. Nitrogen contents ranging from 0,06% to 0,16% have been measured in these fuels. The fuel nitrogen is converted mainly to ammonia and molecular nitrogen in the gasifier. The ammonia content of the low heat value gas is very much dependent on the nitrogen content of the fuel, as shown i Figure 2. It is thus is much lower with sawdust than with wheat straw.
500 0 Cellulosechips Bark and forest or sawdust residue
Willow
Bark
16%RDF in bark
2630%ROFifl sawdust
Wheatstnw
Fig. 2 Typical ammonia concentration in the gas, with different biomass fuels. 526
The production of hydrogen cyanide is usually very low in biomass gasification, unlike in coal gasification where hydrogen cyanide can be one of the major nitrogen containing compounds. At the Varnamo plant hydrogen cyanide levels of 3-6 ppm have been found with cellulose chips, and levels of 20-30 ppm with mixed bark and forest residue. This corresponds to 1-2% of the ammonia levels with the same fuels.
COMBUSTION The Vamamo gas turbine is an almost standard ALSTOM Power Typhoon gas turbine, with modified combustors and burners. In brief, the combustor design is based on the conventional diffusion flame prmciple. The combustor has contra-rotating primary air and fuel addition to promote good shear layer mixing. In addition, a portion of the fuel is injected through plain holes in the burner head to avoid excessive swirl momentum of the fuel gas. The combustion in the turbine has always been reliable. The relatively low heat value of the gas produced in the gasifier has not presented any problems. Due to the low heat value of the gas, the combustion temperature of the Vamamo gas turbme is not high enough to produce any significant amounts of nitrogen oxides (NO,) from the nitrogen and oxygen in the combustion air (thermal NO,). Therefore almost all of the NO, emissions originate from fuel nitrogen (fuel NO,), so that the NO, emission increase with increasing nitrogen content of the fuels, as shown i Figure 3. - "
250
200
e
100
50
0 Cellulose chips Bark and forest or sawdust residue
Bark
Willow
15% RDF in bark
25-30% RDF in
Wheal straw
sawdust
Fig.3 Typical NO, emissions, with different biomass fuels. (At 14-16% 0 2 (dry vol), O2levels depending on gas turbine load.) The main NO, compound is nitric oxide (NO). Due to the rather moderate combustion temperature with the low heat value gas, some of the NO, is emitted as nitrogen dioxide (NOz). The NOz emissions are strongly influenced by gas turbine load, as shown in Figure 4. At low load a relatively larger amount of total NO, is in the form of NOz. This is due to the fact that the compressor is not throttled at part load by guide
527
vane control, resulting in higher air to fuel ratios and thus lower temperature levels at part load. (02levels may rise from 14% at full load up to 16% at part load.)
-
45 -
18 NO2
4 -
c
40 ~-
-
Gas turbine load j
4
- 4.5 -- 4.0
4
35
*
30
*
3
'6
25 20-
--
33
-
3.0
4
-4
..r
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;
- 2*5
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i?
z
8
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8
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1999-01-27 1999-01-28 1999-01-28 1999-01-29 1999-01-29 1999-01-30 1999-01-30 1999-01-31 1999-01-31 12:oo:oo oo:oo:oo 12:w:oo oo:oooo 12:oo:oo oo:oo:oo 12:oo:oo 0o:oo:oo 12:oo:oo
Fig. 4 NO2 ratios at full gas turbine load and 75% load. In the normal operation mode of a base load IGCC plant, i. e. full load, the NO2 emissions amount to around 10% of total NO, with the Varnamo gas turbine. At 75% load, the NOz emissions are 20-25%,and at 50% load around 35% of total NO,. The emissions of nitrous oxide (N20) are normally very low, below or close to the detection limit (1 ppm). As noted above for NOz, more NzO is formed at low temperature levels. At 50% gas turbine load, NzO emissions of a few ppm may be found.
NITROGEN CONVERSION
Gasification The conversion of fuel nitrogen to ammonia has been calculated for a number of steadystate periods with different fuels and different operational parameters of the gasifier (temperature, pressure, load, etc.). The calculations suffer from two difficulties, one being the uncertainty of the fuel nitrogen analysis, and the other the uncertainty concerned with the measurement of the biomass fuel flow. The sampling and flow measurement problems associated with solid fuels are well known. The biomass flow, which is needed for the conversion calculations together with the flow of the low heat value gas, was measured and calculated in three different ways. All three different methods of fuel flow measurement suffered from some uncertainty. Therefore the conversion ratio was found to cover a rather broad range of 40-80%. The average conversion ratio was 60-65%, and was independent of fuel type and fuel nitrogen content. No dependency on operational parameters of the gasifier was found either. This is probably due to the rather limited operating range of the gasifier.
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The ammonia content has been measured at gasifier pressures covering the range of 10-21 bar. No significant influence of gasifier pressure on ammonia formation could be seen in this pressure range. Periods of low ammonia content at low pressure could be explained by higher equivalence ratios. (The normal equivalence ratio was around 0,3. The fuel moisture content was in the range of 5-20%.) The higher equivalence ratio results in a lower heat value of the gas, due to dilution and due to increased oxidation of the gas. The decrease in heat value is accompanied by a decrease in ammonia content. The gasification temperature was varied by some twenty degrees around 930-940°C (gas temperature). At thls temperature the biomass conversion is close to loo%, and practically all of the nitrogen is released into the gas phase. The variations in ammonia content with temperature were probably caused by variations in equivalence ratio, more than by the temperature itself. The conversion ratio from fuel nitrogen to ammonia did not change sigmficantly. Tests with steam injection into the gasifier were also performed, with mixed bark and forest residue (12% moisture in fuel) and with bark only (8% moisture in fuel). The variations in steam injection corresponded to around half of the fuel moisture content. No influence on the nitrogen conversion ratio was seen in these tests.
Cornbustion The conversion of ammonia to NO, decreased sigmficantly with increasing ammonia content. This is seen in Figure 5 , where the NO, emissions increase by only around 50% when the ammonia concentration is increased from 2000 to 4000 ppm. The NO, emissions are expressed as mg/MJ fuel, to eliminate the influence of variations in flue gas volume and oxygen level depending on gas turbine load. (1 ppm at 15% O2 corresponds to 1,7 mg/MJ.) 500.0 450,O
400.0 350,O
-3 -
3
300.0
2 250.0 z
m E 200,o
150,O
100.0 50,O
0
500
1000
1500
2000
2500
3000
3500
4000
ppm NH3 in gas
Fig. 5 NO, emissions as a function of ammonia content during combustion.
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4500
The conversion of ammonia to NO, has also been calculated for a number of steadystate periods with different fuels and different operational parameters of the gasifier, as well as different gas turbine loads. Also these calculations suffer from the difficulty of fuel flow calculation, as the flue gas flow is not measured but must be calculated from the biomass flow. The conversion ratio was calculated to be 5040% (average conversion ratio 6070%) for the standard fuel mixture of bark and forest residue, with an ammonia content of 1000-1500 ppm in the low heat value gas. Bark and willow, with ammonia contents around 2000 and 3000 ppm, had conversion ratios of 40-60% (average 50%). Wheat straw, with ammonia contents around 4000 ppm, had a conversion ratio of only 20-40% (average 30%). Total nitrogen conversion In the calculation of the overall conversion from fuel nitrogen to NO, the uncertainty of the fuel flow is eliminated, as the flue gas flow is also calculated from the fuel flow. The total conversion ratio was 30-50% (average 40%) for mixed bark and forest residue. For bark and willow the conversion ratio was 30-40% (average 35%) and for wheat straw even lower, 20-25%. DISCUSSION
Nitrogen conversion during gasification In gasification and pyrolysis of highly volatile coal, peat and biomass, most of the fuel nitrogen is converted to ammonia and the formation of HCN is almost negligible. For the conversion of fuel nitro en to ammonia, conversion ratios between 10 and 90% are reported in the literature (3-8,. The conversion seem to be affected by the operational conditions, and the results from bench-scale experiments with biomass show that the conversion ratio increases with increasing temperature and generally is higher during gasification than during pyrolysis (4, ', ** 'I. The results from gasification experiments with sawdust and bark at the 90 kWm gasifier of Lund University (Io3 ") show a relationship between the conversion of fuel nitrogen to ammonia and the total carbon conversion of the gasifier. At high equivalence ratios (ER) and temperatures (ER > 0,35,temperature > 900°C), when a carbon conversion close to 100% is achieved, independent of type of fuel, the conversion of fuel nitrogen to ammonia reaches a maximum value of around 50-60%. llus is comparable to the conversion ratios found at the Vamamo plant. At low ERs and temperatures (ER 0,25-0,30, temperature 80O-85O0C) conversion ratios as low as 15% have been found ('I). This is probably due to the fact that the unreleased part of fuel nitrogen in lab-scale experiments becomes accumulated in non-gaseous products of thermal conversion, such as tars and char. There is not much data available on the influence of pressure on nitrogen release during pyrolysis and gasification. However, according to Nichols et al. (12) the concentration of ammonia increases with increasing pressure in a high temperature entrained flow reactor, whilst the conversion of fuel nitrogen to molecular nitrogen is not influenced. Studies on the pyrolysis of peat and bark have shown that the conversion of fuel nitrogen to ammonia is dependent on how the nitrogen is bound in the fuel (3* 13).
530
Accordingly, more ammonia is released from easily decomposed aliphatic nitrogen structures than from more stable aromatic structures. The results from the Vzirnamo gasifier, that none of the above mentioned parameters i. e. temperature, ER, pressure and type of fuel, have much influence on the conversion of fuel nitrogen to ammonia, can probably be explained by the high level of carbon conversion. The thermal conversion of fuels in the Vamamo gasifier, within the examined ER and temperature ranges, produces no char and the conversion of the fuel to tars is very low (less than 1%). In addition there is no indication of the existence of nitrogen compounds in the tar. Thus, the conversion ratio of the Vamamo plant probably presents the maximum possible conversion of fuel nitrogen in a large-scale gasification plant. The fact that this conversion is below 100% can probably be explained by selective non catalytic reduction reactions, involving the NO, and the ammonia produced in the oxidising and reducing zones of the gasifier respectively.
Nitrogen conversion during combustion The gas turbine of the Vamamo IGCC plant was the first, and so far the only one, to be operated on low heat value gas from biomass gasification. No data from similar plants is therefore available for comparison. However data from rig tests, performed by Combustion Technology Centre, agree very well with the results obtained with the Varnamo gas turbine (I4). The gas turbine combustor design for these tests was the same as the one used for the Vamamo gas turbine. The rig tests were performed with a synthetic low heat value gas mixture, with different levels of ammonia. With ammonia free gas, almost no thermal NO, was obtained. The measured NO, concentrations were below 10 ppm (15% O2 dry vol), consistent with the results from the V h a m o plant. The NO, formation was dependent on the ammonia level and the conversion ratios were very close to the conversion ratios found in the Vamamo tests, as can be seen in Figure 5.
Fig. 6 NO, emissions as a function of ammonia content during combustion.
53 1
NO, REDUCTION IN FUTURE BIOMASS AND WASTE FUELLED IGCC PLANTS
The conversion af fuel nitrogen to NO, presents a problem when biomass or waste with a high nitrogen content is used as a fuel. This is a common problem for both conventional boiler plants and IGCC plants. In boilers, NO, reduction can be achieved by primary measures, eg. fuel staging or by selective non catalytic NO, reduction (SNCR) with ammonia injection in the boiler, or selective catalytic NO, reduction (SCR) with ammonia injected in the flue gas duct after the boiler. The SCR technology is applicable also to IGCC plants, and has been used with natural gas fired gas turbines for a number of years. The development of low NO, combustion systems for low heat value gases containing high levels of ammonia is clearly an option. This is being investigated already, as described below. The fact that only around half of the ammonia in the low heat value gas is converted to NO, in the Vamamo gas turbine, indicates that some “low NO,” activity is already taking place. Tlus might be further optimised to achieve an even higher conversion of ammonia to molecular nitrogen. At the operating temperature of the gas turbine, around 1000°C, SNCR reactions may take place between ammonia and NO,. The options to reduce NO, emissions from “second generation” biomass IGCC plants are described below, Ammonia removal before combustion
The first option is to remove the ammonia from the fuel gas before combustion. The advantage of ammonia removal at this stage is that the volume of the gas leaving the gasifier is much smaller than the flue gas volume. Conventional ammonia removal methods, involving condensation and absorption in water, may be used in bioniass IGCC plants. These methods are associated with cooling and potential loss of sensible heat from the gas. Also the absorbed ammonia has to be disposed of. Another option is to destroy the ammonia in situ, in the gas. Several methods, thermal as well as catalytq are presently being developed. Tests which used actual fuel gas, carried out at atmospheric pressure on a side-stream experiment at the Vamamo plant, showed that around 80% of the ammonia could be removed using a selective catalytic oxidation (SCO) approach based on porous aluminium oxide catalysts. In addition, laboratory scale experiments have shown the potential of the SCO approach to remove up to 95% ofthe ammonia in the gas (”). Combustion NO, control
Additional measures to control NO, emissions are based on the further development of the gas turbine combustion system to minimise the conversion of ammonia to NO,. Two such approaches are being developed currently by ALSTOM Power, based on staged air addition or catalytic combustion. The first of these approaches uses controlled air addition to provide a fuel rich primary zone followed by fuel lean intermediate and dilution zones within the gas turbine combustor. The fuel rich primary zone, with an equivalence ratio of about 1,4, promotes the conversion of the ammonia contained in the fuel to nitrogen rather than to NO,. The intermediate zone completes the combustion of the fuel to achieve low CO emissions at an equivalence ratio near unity, whilst the dilution zone provides the correct temperature profile to suit the inlet to the gas turbine. The primary zone of the combustor is fully impingement cooled. This ensures that the local equivalence ratio
532
near to the combustor walls remains close to the optimum vallue of around 1,4 (in comparison, the Varnamo gas turbine uses film cooling in the primary zone). The concept thus provides the double benefit of reducing ammonia conversion to NO, compared with conventional stoichiometric diffusion flame processes, and enhancing load tumdown capability. An extensive test programme was carried out, based initially on a generic combustor with a diameter of 125mm (I6-I8) and subse uently within the framework of the Cleaner Coal Power Generation Group in the UK 20). The aim of this latter programme was to produce a gas turbine combustion system bill of materials for an intermediate size gas turbine with a nominal power output of around 40 MW,. Two prototype combustors, at a diameter of 300were designed, manufactured and tested. Both combustors performed well, reducing the NO, levels substantially from those found in the Vamamo IGCC plant, without compromising other aspects of combustor performance. For example, at an ammonia in fuel gas level of 1000-1500 ppm (typical of that found in barwforestry residues gasification, such as the Vamamo IGCC plant whch utilises hot gas clean-up) the NO, emission was between 70 and 80ppm (15% O2 dry vol). The corresponding ammonia to fuel NO, conversion ratio was between approximately 25 and 30%. Whilst this reduction in conversion ratio is encouraging when compared with the 5040% conversion ratio noted earlier in this paper for the Varnamo plant at same ammonia level, it is still in excess of the legislative limits for natural gas fired gas turbines likely to be in force when “second generation” biomass IGCC plants are constructed. Further development work would be required to reduce NO, emissions using the staged air combustor design principle towards these levels. However, air staged combustion, together with SCO clean-up of the fuel gas, may well be capable of achieving environmentally acceptable NO, levels. A second approach, which shows considerable promise to achieve very low levels of NO,, is that of catalytic combustion. Although still at the laboratory scale, almost complete ammonia conversion to nitrogen has been achieved in recent catalytic combustion studies as part of a UK Foresight Challenge project (’I). This process has patent coverage in Europe, Japan and the USA. It is based on the use of a 2% rhodidalumina catalyst, again using staged air addition to control the reaction conditions to produce a pseudo-two stage process (22). The first stage is to selectively oxidise the ammonia to NO, in a limited oxygen environment, followed by catalytic reduction of the NO, by carbon monoxide andor hydrogen contained in the biomass fuel gas. The next stage in the development of thls approach is to carry out tests at representative gas turbine pressure and velocity conditions using the 2% rhodidalumina catalyst impregnated on a cordierite monolith. This will include testing of a full-scale Typhoon gas turbine combustor.
CONCLUSIONS From the tests performed at the Varnamo plant unique information has become available on the nitrogen chemistry during both gasification and combustion of various biomass fuels in a complete IGCC plant. The conversion from fuel nitrogen to ammonia in the gasifier was found to be around of 60-65% for all fuels. The conversion ratio was independent of fuel type and fuel nitrogen content. No dependency on operational parameters of the gasifier was found. The conversion from ammonia to nitrogen oxides in the gas turbine was dependent on the concentration of the ammonia. At higher ammonia levels the conversion ratio was lower. The conversion ratio was 60-70% for the standard fuel mixture of bark and
533
forest residue, for which the plant was designed. Fuels with higher nitrogen content had lower conversion ratios, the conversion ratio for wheat straw being around 30%. The total conversion from fuel nitrogen to nitrogen oxides was around 40% for mixed bark and forest residue. For wheat straw it was 20-25%. Further development work, which shows considerable promise to reduce the NO, emission from “second generation” biomass IGCC plants, is in progress. This includes: 0
0
0
the removal of ammonia prior to the gas turbine combustor in a SCO process based on porous aluminium oxide catalysts air staged combustor design with a fuel rich primary zone to minimise the conversion of ammonia to NOx catalytic combustor design using a 2% rhodium/alumina catalyst
ACKNOWLEDGEMENTS The work described in this paper was supported in part by ElectricitC de France, Elforsk, Elkraft, Foster Wheeler, the European Union, the Swedish National Energy Administration, the UK Department of Trade and Industry and the UK Engineering and Physical Science Research Council. The authors wish to thank ALSTOM Power, Lund University, Sycon Energikonsult and Sydkraft for permission to publish this paper. The contribution of colleagues in the above organisations to the work reported in this paper is gratefully acknowledged. REFERENCES 1. 2. 3. 4. 5.
6. 7. 8.
9. 10. 11. 12. 12.
Stiihl K., Neergaard M. and Nieminen J. Final Report from the Vlrnamo Biomass IGCC Demonstration Programme, presented at 1st World Conference and Exhibition on Biomassfor Energy and Industry, Sevilla, Spain 5-9 June 2000. Stiihlberg P., Lappi M., Kurkela E., S h e l l P., Oesch P. and Nieminen M. Espoo: VTTEnergy, 1998.45 p. C app. 45 p. (VTTResearch Notes 1903). Lepptilahti J., Fuel, 1995,74, 1363-1368. RosCn E., Bjornbom C., Brage G., Yu Q. and Sjastrom K., In: Chartier P., Ferrero G. L., Henius U. M., Hultberg S., Sachau J. and Wiimblad M. (Eds): Biomassfor Energy and the Environment, London: Pergamon, 1996, pp. 1296-1300. R o s h E., Bjombom C., Yu Q. and Sjostrom K. ‘In Bridgwater A. V. and BoocockD. G. B. (Eds): Developments in Themochemical Biomass Conversion, London: Blackie Academic & Professionals, 1997, pp. 817-827. Leppalahti J., Bioresource Technology, 1993,46. Kurkela E. and StAhlberg P. Fuel Proc. Technol., 1992,31,23. Li C. Z., Kelly M. D. and Nelson F. P., The Australian Symposium on Combustion and the Forth Australian Flame Days, p. 1-6, Nov-1995, Gawler, Australia. Hayhurst A. N. and Lawrence A. D., Combust. Flame, 1995, 100,591-604. Bjerle I., Padban N., Wang W. and Ye Z. PFB Co-Gasification of Fuel Blends, Coal/Textile; Biomass/Textile Wastes, EUR 1928541EN. Wang W., Padban N., Ye Z., Andersson A. and Bjerle I. Ind. Eng. Chem. Rex, 1999,38,4175-4182. Nichols K. M., Hedman P. 0. and Smoot L. D. Fuel, 1987,66,1257. Hamiiltiinen J. P. and Ah0 M. J. Fuel, 1996,75, 1377-1386. 534
14. Al-Shaikhly A. F., Mina T. I. and Neergaard M. 0. Gas Turbine and Aeroengine Conference and Exposition, The Hague, ASME 94-GT-438. 15. Leppalahti J. Development of Selective Oxidation Technology for NO, Emission Reduction in Gasification Power Plants, Final Report for Contract JOR3-CT970157, in preparation. 16. Kelsall G. J., Smith M. A., Todd H. and Burrows M. J. Gas Turbine and Aeroengine Conference and Exposition, Orlando, ASME 91-GT-384, June 1991. 17. Kelsall G. J., Smith M. A. and Cannon M. F. Gas Turbine and Aeroengine Conference and Exposition, Cincinnati, ASME 93-GT-413; published in the Journal of Engineeringfor Gas Turbines and Power, Vol 116, July 1994. 18. Kelsall G. J. and Cannon M. F., 1995, Power Productionfiom Biomass ZZ, Espoo, Finland, March 1995. 19. Kelsall G. J. and Whinfirey J. K. Report No. COAL R0103, UK Department of Trade and Industry, Crown Copyright, June 1997. 20. Kelsall G. J., Whinfrey J. K. and Armstrong S. J. Report No. COAL R0149, UK Department of Trade and Industry, Crown Copyright, June1998. 21. Catalytic Combustion of Nitrogen Bearing Low Heat Value Gaseous Fuels for Power and Heat Generation, ETSUProject Profile 230, Dec 1998. 22. Burch R. and Southward B. W. L. Journal of Chemical Communications, p. 703, 2000.
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Tar Formation in the 18 MWth Biomass-IGCC plant in Varnamo and in a 90 kWth Pressurised Fluidised Bed Gasifier at Lund University Nader Padban’, Soren Hansson’, Magnus Neergaard3, Krister St&h14, and Ingemar Odenbrand’ Dept. of Chemical Engineering II, Chemical Center, Lund University, P. 0.Box 124, SE-221 00 Lund, Sweden 2 Sycon Energikonsult AB, SE-20.5 09 Malmo, Sweden Sycon Energikonsult AB, SE-20.5 09 Malmo, Sweden Sydkrajl AB, SE-205 09 Malmo, Sweden Dept. of Chemical Engineering II, Chemical Center, Lund University, P. 0.Box 124, SE-221 00 Lund, Sweden
’
ABSTRACT The analysis, formation and degradation of aromatic hydrocarbons and tars during air gasification of different biofuels was investigated. The gasification experiments were performed in two different pressurised fluidised bed gasifiers. The first one is an 18 MWh PCFB biomass gasifier based on IGCC technology located in Vamamo, Sweden. The second one is a 90 kWh PFB gasifier at the department of Chemical Engineering 11, Lund University, Sweden. Contents of the light aromatics such as benzene and toluene were monitored by both intermittent and on-line methods. Polyaromatic hydrocarbons (PAHs) were collected by using different sampling methods and were then quantified with GC and GC-MS instruments. Contents of a total number of 80 different PAHs in the gas and fly ash were analysed and compared for different fuels. The content of the tar in the product gas was in the range between 1000-4000 and 10000-20000(mg/Nm3)for Viirnamo and LU-gasifier, respectively. Characterisations made by GC-MS showed that, independently from the kind of the fuel gasified, the tars consisted from similar constituents. The type of fuel and operation conditions influenced the intra-relationshipbetween different tar compounds. Equivalence ratio (ER) and the temperature were two key parameters, whch influenced the amount and the characteristics of the tars in the product gas. Based on both qualitative and quantitative GC-MS analyses the behaviour of the N and 0 containing tar compounds is studied. Although the gas has contained significant amounts of tars, excellent combustion of these has always been achieved and no deposits formed in neither filter nor piping system.
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INTRODUCTION Thermochemical conversion of woody biomass can result in a h g h amount of tars in the product gas. The amount of the tar produced is in general dependent on gasifier type and process condition. In an updraft gasifiers the conversion from fuel to tars could be as high as 10 and 20%. The gas from downdraft gasifiers is usually tar-free but these gasifiers suffer from the low total carbon conversion. The fluidised bed gasifiers show an intermediate behaviour between the up-and down draft gasifiers. The contents of the tars in the product gas from gasification processes must be reduced to a level that is acceptable to the specific application. The extent to which the reduction must be driven is a matter of process and end-user demand for gas quality. In conventional scrubbing systems, the product gas is quenched producing a wastewater laden with tars that must be cleaned before discharge fiom the plant. In addition to the waste water clean up issue, wet scrubbing can leave a small fraction of the tars in the product gas as a fine aerosol mist that is difficult to remove, and can create problems with downstream equipment such as compressors or turbines. Such tar deposition has been observed during integrated system testing of the biomass gasifier-compressor-turbine system at Battelle[l]. For the use of producer gas in e.g. gas engines the tar concentration in the gas must be reduced to levels below 100 mg/Nm3[2]. Hasler et al. [3] gives the same limit for the tar content in the gas for internal combustion (IC) engines. To achieve these low levels of tar concentration in the gas very efficient tar reducing processes, in integration with the gasifiers, are required. For systems based on hot gas cleaning, the tars are only considered problematic if they condense during the transport from the gasifier to the end-user equipment, e.g. the gas turbine. Problems with hot gas filters are reported from laboratory and pilot scale studies at VTT and CTDD (Coal Technology Development Division) showing that tar materials can cause high-pressure drop across the filter [4]. Nevertheless the suspected cause of the high pressure drop for these cases appears to be associated with properties of the filter cake rather than internal blinding of filter elements. There is not much data available in the literature concerning acceptable tar levels for gas turbines. However the successful experiences from Varnamo IGCC demonstration plant shows that the gas turbine can handle several g/Nm3 of tars (naphthalene and heavier compounds) without showing any tendency for sooting, deposit formation and volatile organic compound (VOC) emissions. One other problematic impact of the tars withm gasification systems is the condensation of heavy PAHs on fly ash particles. Since most of PAHs have hazardous properties, study of the PAH capture on fly ash becomes of vital importance. To be able to predict the behaviour of tars in gasification systems detailed information about their amount and composition is an important necessity. The objectives of this study have been the following: a) sampling, characterisation classification and quantification of tars, b) investigation of the relationship between the tar formation and the parameters such as fuel properties, equivalence ratio (ER), temperature and pressure, and c) the effect of different parameters on tar composition. In the studies of tar composition the oxygenated- and nitrogen containing tars were of special concern.
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EXPERIMENTAL EQUIPMENT Vdrnamo IGCCplant Sydkrafl AB has built the world's first complete IGCC Power Plant, which utilises wood as fuel. The plant is located in Varnamo, Sweden, and the technology used in the power plant is based on gasification in a pressurised circulating fluidised bed gasifier. The gasification technology is developed in co-operation between Sydkrafl AJ3 and Foster Wheeler Energy International Inc. The plant at V h a r n o produces about 6 MW electricity to the grid, as well as 9 MW heat to the district heating system in the city of Vtimamo, from a total fuel input equivalent to 18 MW. The normal wood fuel is dried in a separate fuel preparation plant, using a flue gas dryer, to a moisture content of 5 - 20 %. A simplified process diagram of the gasification plant is shown in (Figure 1). R*R
n
Figure 1: Process diagram
The dried and crushed wood fuel is pressurised in a lock-hopper system to a level which basically is determined by the pressure ratio of the gas turbine, and is fed by screw feeders into the gasifier a few meters above the bottom. The operating temperature of the gasifier is 950 - 1000°C and the pressure is approximately 17 bar. The gasifier is of a circulating fluidized bed type and consists of the gasifier itself, cyclone and cyclone return leg. The three parts are totally refractory lined. The fuel is dried, gasified and pyrolized immediately on entering the gasifier. The gas transports the bed material and the remaining char towards the cyclone. In the cyclone, most of the solids are separated fiom the gas and are returned to the bottom of the gasifier through the return leg. The recirculated solids contain some char, which
538
is burned in the bottom zone where air is introduced into the gasifier. The combustion maintains the required temperature in the gasifier. After the cyclone, the gas produced flows to a gas cooler and a hot gas filter. The gas cooler is of a fire tube design and cools the gas to a temperature of 350 - 400" C. After cooling the gas enters the candle filter vessel where the particulate clean-up occurs. Ash is discharged from the candle filter, as well as from the bottom of the gasifier, and is in the meantime cooled and depressurised. The gasifier is of an air-blown type. Thus about 10 % of the air is extracted from the gas turbine compressor, further compressed in a booster compressor, and finally injected into the bottom of the gasifier. The gas generated is burned in the combustion chambers and expands through the gas turbine, generating 4.2 MW of electricity. The gas turbine is a single-shaft industrial gas turbine. The fuel supply system, fuel injectors and the combustors have been redesigned to suit the low calorific value gas (5 MJ/nm3). The hot flue gas from the gas turbine is ducted to the heat recovery steam generator (HRSG), where the steam generated, along with steam from the gas cooler, is super-heated and is then supplied to a steam turbine (40 bar, 455"C), generating 1.8 MWe. The start-up phase was completed during spring 1996 and following that a demonstration programme was launched, which will continue until June 2000. During this period advantages and possible limitations of the new technology are evaluated. Specific areas of interest includes environmental issues, fuel flexibility and production costs in future facilities in addition to the technical development and improvements of the plant. The accumulated operating experience amounts to about 8500 hours of gasification runs and about 3600 hours of operation as a fully integrated plant as per the end of 1999. The test runs have been very successful and the plant has been operated on different wood fuels as well as straw and RDF. One of the last tests includes operation on 100 % straw, which was accomplished without disturbance or problems. . A more detailed description of the plant can be found elsewhere [7,8].
PFB gasifier at Lund University The pressurized bubbling fluidized bed gasifier at Lund University is a 90 kWm consisting of four main units: a bubbling, pressurized fluidized bed gasifier, a hot gas filtration unit, a catalytic conversion unit and a pressurized feeding unit. The test rig is designed for operation at temperatures between 850 and 950°C and pressures between 5 and 20 bar. The fluidized bed reactor consists of a tube, 102 mm in inner diameter having a total length of about 3.3m. Air as a fluidising agent is fed into the reactor bottom through a funnel shaped inlet. The hot product gas enters at the top of the filtration unit. The gas is then cooled and subsequently filtered through a candle filter made from sintered Sic. Gas filtration temperature is normally around 650 "C. The gasifier is equipped with a continuous fuel feeding system for biomass fuels. A more detailed description of the gasifier can be found elsewhere[5,6].
Gasification During the gasification experiments in the Varnamo 18 MWh, pressurised circulating fluidised bed (PCFB), IGCC biomass power plant, the operational pressure and temperature were around 17 bar and 950"C, respectively. The filtering unit operated at a temperature between 300 and 400°C A variety of commercially available
539
renewable fuels such as wood chips, bark, forest residue, straw and different types of RDF were gasified solely or in mixtures with wood chips or forest residue. A sampling pipe connected to the main gas pipe after the filtering unit was used for the collection of tars. The pipe system between the sampling point and the sampling equipment was equipped with electrical heating and hereby the temperature of the gas could be kept at the same level as the temperature in filtering unit. The tars were collected in a series of wash bottles filled with either toluene or dichloromethane. In a large-scale gasification system the operational parameters such as pressure and temperature can only vary within a certain range. Because of this the plant has to operate at temperatures between 900 and 1000°C and at pressures between 15 and 22 bar (e). In studies of the effect of temperature and pressure on tar formation, the results from Varnamo gasifier is compared with the results obtained from the LU-gasifier. Gasification experiments in LU-gasifier were conducted at 12 bar and at temperatures within the range between 800 and 930OC.Although the dimensions and the design of the two gasifiers are quite different from each other, the observed tendencies caused by different parameters were comparable. Characterization and quantifcation of tars and light aromatics Content of the light aromatic compounds such as benzene, toluene and indene in the gas was determined on-line with both the spectrophotometric and chromatographic methods. The constituents of the tars collected from the gas or extracted from the fly ash were identified by GC-MS analysis. The presence of the compounds having large quantities was confiied by doping the samples with pure compounds. Doping of the sample was also used in identification of the structural isomers having retention times close to each other. The quantitative amount of the tars was determined based on the contents of the 22 most abundant PAHs in the gas. The procedures of methods used for calibration and optimising the GC-MS is described elsewhere [9]. Content of the heavy fraction of the tars that could not be analysed with GC-MS, was determined by gravimetric methods.
RESULTS AND DISCUSSIONS CLASSIFICATION OF TARS
Our experiences show that the composition of the tar is dependent on several parameters such as gasification temperature, fuel properties and ER-value. The tars from tests at low ER-value and low temperature in the LU-gasifier (ER= 0.25, temp=-80O0C) a total number of 65 different compounds could be recognised. The number of tar species was lower for the experiments performed at high ER and high temperature. The tars extracted from fly ash contained the lowest number of tar species. Typical chromatograms of different tars are shown in Figure 1. Different species of tars could be classified within four different categories of PAHs: a) ordinary, b) alkyl-substituted, c) oxygenated, arid d) nitrogen containing. Among these, the compounds containing oxygen and nitrogen must be treated with a special concern for several reasons. PAHs containing oxygen, make the base structure for the extremely h a d l chemicals classified as dioxins. As known, dioxins are 540
chlorinated derivatives of oxygenated PAHs. Presence of oxygenated PAHs together with chlorine or chlorine sources such as HC1 are a necessary condition to form dioxins. Nitrogen containing PAHs on the other hand can act as NOx precursors in combustion of the producer gas.
II
12
14.15
I
39.49
Figure 2 (a)
Figure 2 (b)
54 1
16. 17
’*”’
2 0 . 2 1 22
Figure 2 (d) Figure 2. Typical chromatogram of PAHs formed during gasification: a)LU, wood chips; ER = 0.25, temp=-830°C b) LU, wood chips; ER = 0.38, temps-830"C c) Viirnamo, wood chips, E R 4 . 3 , temp=-950°C, and d) fly ash from V&mmo. 12) Pyrene, Key to Figure 2 1) Naphthalene, 13) 2,3-Benzoflueren, 2) 2-Methyl-naphthalene, 14) Chrysene, 3) 1-Methyl-naphthalene, 15) Benzo[a]anthracene, 16) Benzo[b]fluoranthene, 4) Biphenyl, 5 ) Acenaphthylene, 17) Benzo[k] fluoranthene, 6) Acenaphthene, 18) Benzo[a]pyrene, 7) Dibenzofuran, 19) Perylene, 8) Fluorene, 20) Indeno( 1,2,3-cd)pyrene, 9) Phenanthrene, 2 1) Dibenz(a,h)anthracene, 10) Anthracene, 22) Benzo(g,h,i)perylene. 11) Fluoranthene,
TAR QUANTIFICATION Independently from the origin of the tar, the components presented in Table 1 were the most abundant constituents in the tars. The quantitative measurements were based on the contents of these compounds in the tar. Table 1. The main tar constitutes in the gas and fly ash.
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The amount of the tar in the gas and fly ash varies depending on process parameters. Gasification at high temperature and high ER results in a lower content of tars in the product gas. For fly ash the fuel properties and the gas filtering temperature seems to be the two crucial parameters. Table 2 shows the ranges for the measured tar contents in the gas and fly ash. Table 2. Contents of benzene and PAHs in the product gas and fly ash. IRDF
]Bark
]Wood chiDslStraw
lsalix
I
*) Molecular weight.
The amount of the tars in the fly ash from bark tested in the V b a m o gasifier is in general higher than in fly ash fiom wood chips. It seems to be a relationship between the fixed carbon content of the fuel and the tars in the fly ash. Fuels with a high fixed carbon such as bark result in a higher amount of tars in fly ash than the fuels with low fixed carbon. However there is an indication that the fuel particle size have big influence on the tar content of the fly ash. The total amount of the unburned carbon in the fly ash fiom both bark and wood chips in Vamamo plant is in the same range and at the same level as in the fly ash from combustion plants. There is a very clear difference between the compositions of the tars in the gas and the tars extracted from fly ash. Whilst naphthalene is the main constituent of the tars in the gas, heavy PAHs such as phenanthrene, flouranthene and pyrene are more abundant in the tars from the fly ash. However the proportion of the tars condensed on the fly ash particles is very small and only equal to a few parts per thousands of the total tar in the system. In LU-gasifier the difference between the amounts of the tars in fly ash from different fuels becomes negligible. The explanation is that the gas filtering temperature at LU-gasifier is performed at comparatively high (> 65OOC) temperature. The high filtering temperature prevents the tars fiom condensing on the fly ash particles.
543
Effect of thefuel and temperature on tar composition
In tars collected from the product gas, naphthalene is the dominant compound followed by phenanthrene, acenaphthylene and pyrene. Table 4 shows a comparison between the tars fiom two different fuels and two different operational conditions in the Varnamo gasifier. The presented values are an average of two to four measurements in each case. Table 4. Comparison between the tar from different fuels tested in Vlmamo.
*) T1: Normal gasificationtemperature, T2: 30°C below T1 gasification temperature. As can be seen in the table there is a big difference between both the amount and the composition of the tar fiom barWforest residue-mixture (60/40%wt.) and from willow. At similar condition, gasification of the bark-forest residue mixture results in a higher yield of tar than the gasification of willow does, showing the importance of the fuel characteristics in formation of the tars. Nevertheless the mechanisms that decides the tar formation are still unclear and it is not possible to draw a general conclusion between the amount of the produced tar and the fuel inherent properties such as volatility and fixed carbon content. On the other hand there is a clear temperature effect on both the amount and the composition of the tar: with increasing temperature the total amount of the tar decreases and the proportion of the heavy compounds in the tar declines. Oxygenated and nitrogen-containing compounds
The oxygenated compound in the tar can be divided into two different groups: one consisting of single ring compounds and the other one consisting of polycyclic 544
compounds. Single ring oxygenated compounds are mentioned under the name phenolic compounds in the literature. Fonnation of phenolic compounds is associated with the low pyrolysis and gasification temperature. The phenol peak could be identified in GC-MS chromatograms of tars from tests in LU-gasifier at low temperature and low ER-values. At higher temperature these compounds undergo degradation reactions and dissociate into species with lower molecular weight. Several measurements from the Vamamo gasifier showed that the contents of the phenolic compounds in the product gas was zero. The presence of oxygenated polyaromatic compounds in the product gas is important because they can act as dioxin sources. Aside for dibenzofiuan, the analysis instrument was not calibrated for quantification of other oxygenated polyaromatic compounds. In the gas from LU-gasifier, dibenzofuran stands for between 0.5 and 1.5 %wt. of the total PAHs. The corresponding value for the Varnamo gasifier is considerably less and in most cases dibenzofuran does not exist in the gas. However to be able to study the effect of different parameters on behaviour of these compounds the ratio between the peak area for oxygenated compound and peak area of the dibenzofuran were calculated. These calculated values are presented in Table 5. Table 5. Comparative values for the peak area of oxygenated aromatics.
1
I I
1
LU-bark
LU- woodchips
I
ER-value Temperatur ("C) dibenzofiuan
0.25 10.25 10.39 10.33
1.oa 1.00I 1.00I 1.00I 1.oc I other OJ vgenated tar compounds 3,4-methoxybenzaldehyde 0.00 0.00 phenol 0.99 0.00 o-cresol 0.00 0.00 0.00 p-cresol I 0.02 0.00 2-ethylphenol I 0.00 0.001 0.00 0.001 3,5-xylenol I 0.00 2.3 -dimethvluheno1 I 0.00 0.03) 0.001 0.001 0.OC dlhvdrobenzofuran I 0.00 3-ethvl-5-methvluhenol I 0.00 xanthene 2,2-diphenylacetamid total I I I I I I I Fuel-N (%wt.) 0.55 0.27
4 0.00 0.00 0.00 0.00 0.00 0.001
33
~
*, **Wood chips respectively bark gasification in the Vamamo plant. 545
As seen in the table the total content of oxygenated aromatics decreases by increasing temperature. At a relatively narrow range of temperature changes, more oxygenated compounds are present in the gas at lower ER-values than it does at high ER-values. The samples from Vamamo, despite the low ER, are almost free from the oxygenated compounds. The higher gasification temperature in Viirnamo plant can explain this. If we assume that the response of GC-MS detector is the same for dibenzofuran and other oxygenated compounds, the content of the latter mentioned in the gas from LU-gasifier can be as high as the 2%wt. of the total. The observed tendencies of ER-value and temperature effect indicate that probably, the oxygenated compounds are products of biomass primary dissociation to smaller fragments. At higher temperature these fragments are converted to lighter compounds and permanent gases in a series of chemical reactions. Within the examined ranges at LU, the changes in the temperature and ER-value do not considerably affect the content of the nitrogen containing compounds. However a high inherent amount of the fuel nitrogen seems to enhance the formation of these compounds, hence a higher value was measured for the bark. Absence of nitrogen containing components was observed in the gas from Viirnamo plant. This is probably due to the very large operational temperature difference between the LU and Varnamo gasifiers. COMBUSTION BEHAVIOUR OF THE LIGHT AROMATICS AND TARS IN GAS TURBINE The product gas from V h a m o gasifier is, after filtration, combusted in a gas turbine to produce electricity. The tar related gas filtering problems reported by other sources were not observed at Varnamo IGCC plant during the test runs. During the regular inspections of the hot gas filter units, gas piping, turbine combustion chamber and gas turbine blades no indications of deposit formation caused by tar or other problematic materials such as soot could be found. This shows that the constituents of the tars are at gaseous state from the gasifier to the combustion chamber. The content of the volatile organic compounds (VOC) in the exhaust from the gas turbine was measured with an on-line instrument at different occasions. The instrument used was a Bernard Atomic 3005 with a detection limit equal to one ppm propane. Despite the high content of the aromatics and PAHs in the fuel gas (650010000 and 1000-4000mg/Nm3,respectively) the emission of VOC from the plant was below the detection limit for the analysis instrument. Dioxin measurements in flue gas and fly ash have been done for fuels with high content of chlorine. No dioxin has been detected in these measurements.
SUMMARY Tars from two different pressurised fluidised bed gasifiers were qualitatively and quantitatively analysed. Utilising the same kind of the fuels, the total amount of the tar in the gas fi-om the small scale LU-gasifier could be as high as ten times above the corresponding value for the Varnamo-IGCC plant. The content of the tar captured on fly ash particles showed to be dependent on the gas filtering temperature, probably the fixed carbon content of the fuel and the particle size of the fuel. High gas filtering
546
temperatures decrease the content of tars in the fly ash. The total amount of the unburned carbon in the fly ash was comparable to corresponding values from the combustion processes. The ER-value and the temperature affected the composition of the tars in both gasifiers. The content of the oxygen containing aromatics and poly-aromatics decreased by both increased ER-value and increased temperature at LU-gasifier. The contents of the nitrogen containing tar components seemed to be affected by the fb-N in the feedstock. Except dibenzofuran, no other oxygen or nitrogen containing tar components could be detected in the gas from V i a m o IGCC-plant. Probably due to a high gasification temperature such compounds undergo degradation reactions inside the gasifier. There was no indication to deposit formation caused by tars and the amount of VOC in the exhaust from the gas turbine is below the detection limit for the analysis methods. ACKNOWLEDGEMENTS The Varnamo IGCC research and development project has been carried out during 1996-2000. The work has partly been performed in collaboration between Sydkraft, Foster Wheeler, Electricitt de France and Elkrafi. It has also been partly financed by Elforsk AB, Sweden, The Swedish National Energy A b s t r a t i o n and the European Commission. REFERENCES Mark, A., Paisley, P.E., In: Developments in thennochemical biomass conversion. V. 2. Bridgwater, A.V. Boocock, D.G.B. London: Blackie Academic and Professional. 1997. p. 1209-1223. Stassen, H.E.M, Venendaal, R., Knoef, H.A.M. (1993) UNDP/WB small-scale biomass gasifier monitoring report, Vol. 1 Findings. Hasler, P. h., Buechler, R., Nussbaumer, T. h., In : VTT Symp. (1999), 192 , 37 1-382. Cahill, P.; Nieminen, M.; Dutton, M.; Rasmussen, G.; Kangasmaa, K., High Temp. Gas Clean., [Pap. Int. Symp. Exhib. Gas Clean. High Temp.], 3rd (1996), 782-793. Editor(s): S c h d t , Eberhard. Publisher: Institut h e r Mechanische Verfahrenstechmk und Mechanik der Universitaet Karlsruhe (TH), Karlsruhe, Germany. Hallgren, A.; Bjerle, I. and Chambert, L. A. PCFB gasification of biomass. In proceedings to the 206th American Chemical Society National Meeting, Fuel Chemistry Division, Chicago IL, 1993. Hellgren, R. Thermochemical conversion of biomass: A techmcal feasibility study concerning wood pyrolysis at high temperatures and biomass gasification at elevated pressures. Licentiate Thesis. Department of Chemical Engineering 11, Lund University, Lund, Sweden, 1995. StAhl, K.; Neergaard, M. In: Sipila, K.; Korhonen, M. (Eds): Power production from Biomass 111, Gasification and Pyrolysis R&DnD for Industry, Espoo: VTT, 1999; pp 73-78.
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(8)
(9)
Varnamo Demonstration Plant, A demonstration plant for biofuel-fired combined heat and power generation based on pressurised gasification, Published by Skogs Boktryckeri AB, 1998, Trelleborg, Sweden. Padban, N.; Odenbrand, I., Energy h Fules. 1999,13 (5), 1067-1073.
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Final Report: Varnamo Demonstration Programme Krister Stfill, Magnus Neergaard2 and Jonna Nieminen3 Sydkrajl AB, SE-205 09 Malmo, Sweden Sycon Energikonsult AB, SE-205 09 Malmo, Sweden Foster Wheeler Energia Oy, P. 0.Box 201, FI- 78201 Varkaus, Finland
'
ABSTRACT: Sydkraft AEi has built the world's first complete IGCC Power Plant, which utilises wood as fuel. The plant is located in Vamamo, Sweden, and the technology used in the power plant is based on gasification in a pressurised circulating fluidised bed gasifier. The gasification technology is developed in co-operation between Sydkraft AB and Foster Wheeler Energy International Inc. The plant at Vamamo produces about 6 MW electricity to the grid, as well as 9 MW heat to the district heating system in the city of Vamamo, fiom a total fuel input equivalent to 18 MW. The Varnamo plant is an important step forward in developing hghly efficient and environmentally acceptable technologies based on biomass fuels. Experiences gained in this co-generation plant will be utilised in design of new and larger plants which will generate twice as much electricity as a conventional steam cycle plant with the same amount of heat demand. The start-up phase was completed during spring 1996 and following that a demonstration programme was launched, which will continue until June 2000. During this period advantages and possible limitations of the new technology are evaluated. Specific areas of interest includes environmental issues, fuel flexibility and production costs in future facilities in addition to the technical development and improvements of the plant. The accumulated operating experience amounts to about 8500 hours of gasification runs and about 3600 hours of operation as a fully integrated plant as per the end of 1999. The test runs have been very successful and the plant has been operated on different wood fuels as well as straw and RDF. One of the last tests included operation on 100% straw, whch was accomplished without disturbances or problems. INTRODUCTION Increasingly heavy demands are expected on future power plants in terms of efficiency, impact on the environment, fuel flexibility, power prodtiction costs etc. Biomass fuels are in most countries a domestic source of fuel and are often found as waste products in different lunds of industries e.g. agriculture, forestry, pulp549
and paper. Further, a lot of biomass waste such as e.g. packaging material is presently landfilled and in the future there will probably be more stringent requirements to recycle garbage, thereby reducing landfilling. Integrated gasification combined cycles (IGCC) have been developed and demonstrated for power generation using fossil fuels as feedstock. The main features are the possibility of cleaning the gas fiom impurities, such as particulates, sulphur, etc. under pressure before the gas enters the combustor of the gas turbine, and also the relatively h g h electrical efficiency. Higher efficiencies also means relatively lower emissions. On the basis of these considerations, Foster Wheeler Energy International, Inc. and Sydkraft AE3 have been developing the pressurised IGCC for biomass fuels since 1991. In June 1991, Sydkraft took the final decision to build a co-generation plant at Vamamo, Sweden, to demonstrate the technology. The plant generates 6 MW of electricity and 9 MW of heat for district heating. The Varnamo Demonstration Plant is the first of its lund in the world. The plant is aimed at demonstrating the complete integration of a gasification plant and a combined cycle plant, fuelled by biomass. The basic idea is to demonstrate the technology rather than to run a fully optimised plant. Flexible and conservative solutions were chosen for the plant layout and design, to ensure the success of the project and to make the plant suitable for R & D activities. The Demonstration Programme is now almost finished and the results are being evaluated and reports produced. Studies concerning fbture plants based on the experience gained at Vamamo are also being performed and these will present expected investment costs as well as operating costs for the next generation of demonstration and commercial plants. An aerial view of the Vamamo plant is shown in Figure 1.
Figure 1: Aerial view of the V h a m o Demonstration Plant 550
THE PROCESS The wood fuel is dned in a separate fuel preparation plant, using a flue gas dryer, to a moisture content of 520%. A simplified process diagram and a cross section of the gasification plant are shown in Figure 2 and 3.
Figure 2: Process diagram The dried and crushed wood fuel is pressurised in a lock-hopper system to a level which basically is determined by the pressure ratio of the gas t&bine,-and is fed by screw feeders into the gasifier a few meters above the bottom. The operating temperature of the gasifier is 950 - 1000°C and the pressure is approximately 18 bar (g). The gasifier is a circulating fluidized bed and consists of the gasifier itself, cyclone and cyclone return leg. The three parts are totally refractory lined. The fuel is dned, gasified and pyrolyzed immediately on entering the gasifier. The gas transports the bed material and the remaining char towards the cyclone. In the cyclone, most of the solids are separated from the gas and are returned to the bottom of the gasifier through the return leg. The recirculated solids contain some char, which is burned in the bottom zone where air is introduced into the gasifier. The combustion maintains the required temperature in the gasifier. After the cyclone, the gas produced flows to a gas cooler and a hot gas filter. The gas cooler is of a fire tube design and cools the gas to a temperature of 350 400OC.After cooling the gas enters the candle filter vessel where the particulate cleanup occurs. Ash is discharged from the candle filter, as well as from the bottom of the gasifier, and is in the meantime cooled and depressurised. The gasifier is air-blown. Thus about 10% of the air is extracted from the gas turbine compressor, M e r compressed in a booster compressor, and finally injected into the bottom of the gasifier.
55 1
The gas generated is burned in the combustion chambers and expands through the gas turbine, generating 4.2 MW of electricity. The gas turbine is a single-shaft industrial unit. The fuel supply system, fuel injectors and the combustors have been redesigned to suit the low calorific value gas (5 MJ/nm3). The hot flue gas from the gas turbine is ducted to the heat recovery steam generator (HRSG), where the steam generated, along with steam from the gas cooler, is super-heated and then supplied to a steam turbine (40 bar, 455"C), generating 1.8 MWe.
Figure 3: Cross section of the Varnamo gasification plant The plant is equipped with a flare on the roof of the gasification building, which is used during start-up procedure and when testing less well known conditions, in order to protect the gas turbine. In Table 1 and 2 the technical data as well as the suppliers are listed.
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Table I : Technical data of the V h a m o Plant Powerheat generation Fuel input Fuel Net electrical efficiency (LCV) Total net eficiency (LCV) Gasification pressure / temperature Lower calorific value of Product Gas Steam pressure I temperature Plant owner
6 MWe 1 9 MWth 18 MW fuel (85%ds) Wood chips. (Several other fuels have been tested with good results.) 32% 83% 18 bar (g) / 950°C 5 MJ / m3n 40 bar (a) 455°C Sydkraft AB
Table 2: Suppliers of the Vamamo Plant Suppliers: Engineering Gasifier Gas cooler Ceramic hot gas filter Metallic hot gas filter Gas turbine Booster compressor Heat recovery steam generator Steam turbine Plant control system
SYCON / Foster Wheeler Foster Wheeler Foster Wheeler Schumacher GmbH Mott Corporation Al3B Alstom Gas Turbines Ingersoll-Rand Foster Wheeler Turbinenfabrik Nadrowski GmbH Honeywell
DEMONSTRATION / DEVELOPMENT PROGRAMME An extensive demonstratioddevelopment programme has been carried out during
1996-2000. The work has partly been performed in collaboration between Sydkraft, Foster Wheeler, Electricite de France and Elkraft. It has also been partly financed by Elforsk AB, Sweden, The Swedish National Energy Administration and the European Commission. The overall aim of the demonstration programme was to verify the status and future potential of the biomass IGCC concept, utilising the Bioflow technology from a technical and economical point of view. In order to achieve this it has been important to identify and verify the status of different parameters e.g. operability, maintainability and availability. Of particular interest to the success of the gasification technology is to verify the quality of the gas produced in the gasifier as well as the operation of the gas turbine. In parallel to the test runs an extensive work to estimate investment, operational and maintenance cost for future plants, based on experiences from the Vamamo plant have been performed.
553
EXPERIENCE GAINED DURING TEST OPERATION GENERAL Commissioning of the plant started in late 1992 by start-up of the fuel preparation plant. Commissioning of the combined cycle was completed on liquid fuel during March 1993. The first gasification test on wood chips at low pressure was performed in June 1993, and combustible gas was produced and burned in the flare. It should be remembered that at the time for commissioning of the gasifier, no experience existed from any biomass gasifiers at this pressure level. Accordingly, during tests with different bed materials, temperatures and pressure levels, deposits sometimes occurred. Deposits and fouling have verified the importance of carefully controlling the process as well as ensuring a suitable design of components. During the Demonstration Programme, magnesite (MgO) has been used as bed material in the Varnamo gasifier, and this has proved very successful. As magnesite is more expensive than dolomite tests were carried out to check feasibility of re-circulation of bottom ash and these proved very successful. However, we still believe that it will be usefbl to continue testing different bed materials or mixtures of bed materials to further optimise the gasification process and achieve the best result i.e. minimum of deposits, cost and best possible gas quality. On the other hand, deposits can also be handled with a suitable design of the gasifier and the down-stream components. Apart from this, already during the early design stage, in particular two areas were of great concern, namely the gas clean-up and the gas quality Concerning the hot gas filtration one of the ideas behmd this is of course to allow gaseous tars to pass through the filter and other tars to stick to the filter cake and not pass into the fine pore structure of the filter itself. As the amount of benzene and tars is not insignificant from gas heating value point of view, this is very important to achieve. The diagram below (Figure 4) shows a typical operating curve for the hot gas filters and clearly indicates that no continuous increase in pressure drop is taking place. IGCC Varnamo
I
80 12.05
12.10
12.15
12.x)
12.25
12.30
12.35
12.40
12.45
12.50
I
Figure 4: Pressure drop in a hot gas filter cleaned by nitrogen pulsing
554
12.55
GAS QUALITY During the commissioning as well as the demonstration programme the gas quality has been checked regularly. The gas quality regarding hydrogen content turned out to be slightly lower than predicted, but the heating value has been maintained by an increase in methane.
Table 2: A typical range of dry gas composition, % volume
co
H2
CH4
16-19%
9.5-12%
5.8-7.5%
c02 14.4-17.5%
N2 48-52%
The percentage in Table 2 is by volume and gas heating values in the range of 5.0 - 6.3 MJ/m3nhave been recorded. Different operating conditions in the gasifier as well as a change of fie1 produce different amounts of light tars and benzene as can be seen in table 3 below. Bark tends to produce less benzene and tars than ordinary wood chips.
Table 3: Benzene and light tars in the gas Fuel Bark 60% and forest res. 40% Pine chips
Benzene mg/m3n Light tars, mg/m3n 5000 - 6300 1500 - 2200 7000 - 9000 2500 - 3700
During almost all tests it has been general practice to measure not only the main gas constituents, but also benzene and light tars. An interesting correlation between the amount of light tars and benzene in the product gas has been noticed, which can be seen in Figure 5 below. A correlation also exists between methane and benzene, see Figure 6. Light tars and Benzene in product gas from tested fuels
Benzene (mglnm3 dry gas)
Figure 5: Amount of light tars vs benzene in product gas
555
Benzene och Methane in product gas from tested fuels
4.50
5.00
5.50
6.00
6.50
7.00
7.50
8.00
Methane K (dry gas)
Figure 6:
Measured amount of benzene vs methane in the product gas
Due to the relatively low combustion temperatures in the gas turbine combustors when burning product gas thermal NO, is very low. Total NO, emissions can however be higher compared to operation on liquid fuel with steam injection due to the conversion of fuel bound nitrogen into NO,. From Figure 7 below the influence of the amount of fuel bound nitrogen is evident. The recorded levels of alkalines have been below 0.1 ppm wt.
140,OO
120,oo 100.00 80.00
60.00 40.00
20.00
0,oo
20 -CFBOl FQOOl NOconcenttation in turbine exhaust CFBOlFQ002: NOPconcentration in turbine exhaust
-
Figure 7: Gas turbine exhaust emissions (ppm) 556
1
HOT GAS FILTER PERFORMANCE Originally a hot gas filter of the ceramic type was installed. This filter consisted of ceramic filter candles arranged in six groups with separate back-pulsing. The ceramic filter showed good filtration efficiency, with stable pressure drop. However, after more than 1200 hours of trouble free operation suddenly two ceramic candles broke. The difficulty to detect a relatively small failure in a hot gas filter was then noticed in practice even though no serious damage was caused. The broken candles as well as other candles from the filter were analysed but the reason for the breakdown was never traced by the supplier. The complete set of candles was changed to a new design of ceramic candles and t h s was installed in the plant. After less than 350 operating hours one of the new type of candles broke. The break-down was established by the supplier to be caused by mechanical fatigue since micro cracking was found in all tested elements and a chemical attack was excluded. To find out possible reasons for the mechanical fatigue, measurements inside the pipe (noise) as well as in pipework and steel structure have verified the vibrations to be very low. Finally to avoid any risk of fatigue, grids supporting the candles have been installed. However, since the installation of this grid a tendency of bridging has been observed, which we did not experience earlier. To protect the gas turbine in case of a hot gas filter break-down a metallic police filter has been installed down-stream of the main filter. Rather soon after this installation was completed another failure of the main filter occurred. From figure 8 can be seen that the pressure drop over the police filter increases rapidly, whereas it is hard to spot the broken candle from the pressure drop across the main filter. In the figure 8 the upper curve represents the police filter pressure drop, whereas the lower curve represents the pressure drop across the main hot gas filter.
15.15.00
15.25.00
I -RMAIOCPW,dp hd-
15.36.00
lime filter
Figure 8:
15.45.00
15.55.00
-~ ~ 9 1 o c w odp6 ,pdicefilter
Hot gas filter failure
During the summer 1998 it was decided to install metal filter candles instead of the ceramic candles in the main hot gas filter. The metal filter candles are installed in the original filter vessels but with a new tube sheet and back-pulsing arrangement. The metal filter has, like the ceramics, shown very good filtration efficiency, with stable pressure drop. This filter has now been in operation for more than 2500 557
hours without any filter breakage or other damage during operation. Investigations carried out after the end of the last test indicate that there is no degradation of these elements although they have been exposed to gas and ash not only from wood chips but also from RDF and straw. GAS TURBINE EXPERIENCE
The gas turbine installed in the plant is an almost standard Typhoon from ABB Alstom Gas Turbines in Lincoln, England. Modified components are the combustors, the burners and the addition of an air bleed from the compressor. Furthermore, has a special design gas control module been developed to control the product gas, steam and nitrogen to the unit. In September 1995, the commissioning of the modified gas turbine on liquid fuel was completed, and the first test runs on product gas were performed in October. In order to minimise the risks, the first test runs were very short and product gas was introduced gradually with a corresponding reduction in liquid fuel, which eventually resulted in operation solely on product gas. Already prior to being supplied to V h a m o , the special combustors and burners were tested in a rig in England utilising synthetic gas. Combustion has always been reliable in the turbine whether operating on gas fuel or liquid. The relatively low heating value of the gas (about 1/10 of natural gas) proved to be of no problem to the gas turbine and a stable flame has always been established even when the heating value has been lower than normal. Not even initially it was necessary to maintain a pilot flame of liquid fuel and thus has all operation for 3600 hours been on 100% gas and no liquid. This is valid within the full operating range from 40% to full load. Internal inspection of the turbine and combustors is camed out after every test run and we can now say that the hot gas clean-up is operating most satisfactory and no damage to the delicate parts of the turbine etc has been observed. A thorough combustion of the hydrocarbons has always been registered with results between 1 and 4 ppm, whereas a slightly high figure of CO has been observed with figures up to and sometimes even above 200 ppm on part load. Work is however going on to improve this. As has been mentioned before (Figure 7) levels of NOx around 150 ppm has been recorded when operating on gas produced from biomass with high nitrogen content (like bark) while the lower nitrogen content of hardwood considerably reduces the NOx down to a mere 50 ppm. Recent tests on 100% straw, which is a nitrogen rich fuel (about 1% of dry fbel) have produced a NOx figure of slightly above 250 ppm. The development of new combustors will reduce the formation of NOx and further development of selective catalytic oxidation (SCO) of NH3 and HCN will most likely further reduce the emissions. The conventional SCR technology, the cost of which has lately been reduced, can of course also be used. In the early days of the project the possibility of a small gas turbine to burn a LBTU gas was a great issue. In 1995 it was proved in Varnamo that with the gas normally produced by the gasifier i.e. a heating value of 4.5 MJkg this can be accomplished without the use of a pilot flame. As during all tests, flame stability has not been a problem we had to deliberately and gradually reduce the gas quality in a test to find out when flame-out would OCCLX. In figure 9 below this is shown. Since the engine was running at about 50% load and steam was used to dilute the gas it is likely
558
that operation could have been maintained also at a lower value than the one achieved ( 0 . 8 MJkg) during less extreme conditions.
(3)
Gas LCV (MJkg) Gas Turbine output (MW) Gas Turbine exit temperature (“C)
Figure 9: Test of gas turbine flame-out FUEL FLEXIBILITY
The fuel feeding system with the lock-hopper, pressurised fuel silo and screw feeders is primarily designed for wood chips and fuels with equivalent densityheating value per m3. During commissioning and the first years of testing, forest residue and wood chips was the fuel used in general. A number of different fuels have however been tested in the plant during the last couple of years. As fuels with considerably lower density are of interest from gasification point of view these low density fuels has to be pelletized due to the installed type of fuel feeding system. Accordingly we have used straw, willow, bark and sawdust pellets. For testing purposes it is also beneficial to be able to mix different types of fuels to a predetermined ratio and this is of course considerably easier with the fuels in pelletised form. The following fuels have been tested:
0
0
Wood chips Forest residue (bark, branches etc.) Saw dust and bark pellets Willow (salix) Straw RDF
All these fuels have proved to be good and easy to gasify without causing deposits or sinter in the systems. The gas produced when operating on wood chips made from hard wood contains more benzene and tars as shown in table 3 above than
559
forest residue containing bark. Bark has actually proved to be an excellent fuel and even feed rates up to 100% bark is easily gasified and the gas is very good for filtration and gas turbine operation. The rather high levels of alkalines in willow (salix) has not caused any problems in any part of the system and the amount of sintered material in the bottom ash was very small, Tests have been performed with willow in pelletised form and from small amounts up to 100%. The only clear and negative effect observed was a reduction in heating value of the gas. The change was however not so great and the gas turbine operation was not affected. Straw has always been considered a very difficult fuel to budgasify due to its high levels of alkaline and great amount of ash in the fuel. Also the chlorine level is very high in comparison to wood fuels. Tests have been carried out with straw as a mixture with bark, but lately tests with 100 % straw has been performed. About 200 tons of straw has been gasified without any problems or sintering, and a gas was produced with a hydrogen content slightly higher than normal, which proved excellent for gas turbine operation. The tested RDF was in a pelletised form. The size of these pellets is though bigger than the size used for most other fuels (18 mm vs 8 mm). The pellets are produced from waste paper, plastics, cardboard, alumina etc. Tests with up to 50% RDF have been carried out and make up fuel has been sawdust and bark. Encouraging results have been achieved including gas turbine operation on the gas produced. PLANT CONTROL EXPEMENCE
During normal operation the flare valve is closed and the pressure in the gasifier is controlled through a by-pass valve at the booster compressor. Quite naturally is the gasifier a “slow system” concerning both pressure and temperature control. The output control of the gas turbine is completely different and it responds more or less instantaneously. Operation in the fully integrated mode made the pressure, temperature and gas quality in the system vary a bit when the gas turbine suddenly compensated for a small change in either parameter. This feature of the Bio-IGCC was apparent already during the design stage of the project, and the gas turbine supplier was asked to operate the unit with the control valve “fully open”. This way of operation should have several benefits (lower pressure in the gasifier, reduced requirement of auxiliary electric power etc.). Finally it should have a positive effect on the conditions in the gasifier. A test simulating this has been performed and the results are evident fiom the figure below. Gradually all parameters in the gasifer stabilise completely and since the process runs smoother the variation in output fiom the gas turbine is also very small.
560
(1) (2) (3) (4)
Figure 10:
Gas turbine valve position (YO) Gas flow (kg/s) Gasifier pressure (bar) Gas LCV (MJ/kd
Plant output control (First GT control 3 Gasifier control)
ECONOMY AND MARKET OPPORTUNITIES
In Europe as in most parts of the world all kinds of fossil fuels are relatively cheap. Fuels based on biomass can accordingly not compete unless some kind of support or subsidies are given fiom the government or alternatively fossil fuels are taxed more than today. Since these instruments not yet have been hlly utilised, comparisons will here be made only between technologies burning the same kind of fuel i.e. biomass. As the technology used in V&namo never before have been tried the cost of the plant was very high. A considerable amount of valuable experience has been gained, which can be used in the next generation of demonstration plants. Some demo plants must be built to further develop the technology and reduce the cost of the plants, and these demo plants will require financial support as they will not be commercially viable. Particular areas for development include fuel system, inert gas system and emissions (NOx). After this development and the technology has matured we have calculated the investment cost to be around 1500 EUROkWe for a 55 MWe plant producing back pressure steam. For a smaller plant of 15 MWe the cost is estimated to 2300 EUROkWe. The cost of the produced electricity from the 55 MWe IGCC plant is shown in figure 11 below, where a comparison is made with a competing technology namely the CFB boiler. As the comparison is made for a back pressure installation both plants produce 65 MW steam, but the IGCC will produce 55 MWe whereas the CFB will be limited to 20 MWe. In cases where the high amount of electricity for a fixed heat demand is valuable the IGCC comes out very favourable, and as can be seen
561
from the diagram the IGCC is competitive already at around 33 EUROMWh and increasingly so for higher cost of electricity. IGCC 55 MWe165 MW process heat (mature technology) vs. CFBlST 20 MWe165 MW heat Heat generation cost vs. Electricityprlce
-
60.00 50.00 40.00 p 30.00
g
20.00
- * t - -CFBIST
2
10.00
J
0.00 -10.00 -20.00 -30.00 Electricity (EURiMWh)
Figure 11: Cost of electricity versus value of produced heat (back pressure steam).
In Figure 12 below, the comparison is made for condensing plants with 70 MWe output. The cost of electricity for a certain fuel price is shown for the biomass IGCC unit as well as for a CFB using the same fuel. The high electrical efficiency of course makes the IGCC come out very well at high fuel prices. However, as is evident fiom the diagram it is competitive already at a very low fuel cost.
-
IGCC vs. CFBlST 70 MWe Only Power Application Generation cost vs. Fuel price 80.00 70.00
= 60.00
3
5 50.00 P
f
- - * - -CFB
40.00
30.00
ii 20.00
10.00 6.0
7.2
8.4
9.6 10.8 12.0 13.2 14.4 15.6 16.8 18.0 19.2 Fuel prlce (EURIMWh)
Figure 12: Cost of electricity as a function of fuel cost
562
It is likely that the use of RES (Renewable Energy Sources) will be promoted in the future and the EU has announced a very ambitious program to increase the use of RES until 2010, amounting to 12% of the energy supply. The biomass IGCC as well as the CFB will benefit from ths, and political actions will further strengthen their situation. In figure 11 above was shown that from around 33 EUROMWh the IGCC has an advantage over the CFB. However, also from other aspects the IGCC may become successful as either alternatives are hard to find, or very expensive. This applies to retrofit of natural gas combined cycles (i.e. when a certain amount of renewable fuels should be used in existing power plants) as well as when “difficult fuels” should be used such as straw, bagasse and RDF.
CONCLUSIONS FROM THE DEMONSTRATION PROGRAMME The difficulties encountered initially in the Varnamo Project were overcome after a couple of years of intense commissioning and testing. The Demonstration Program, started up during 1996, has been very successhl and has proven that the pressurised biomass IGCC technology works. The best proof of this is the acheved number of operating hours. Accordingly the complete plant has been in operation with the gas turbine operating solely on product gas produced by the gasifier in excess of 3600 hours. Huge experience has been gained from gasifier operation for more than 8500 hours. Results acheved can thus be summarised as: High pressure gasification technology works. Gas produced can be burnt in a gas turbine under stable conditions. Hot gas filtration is efficient and reliable , Technology is capable of gasifying “difficult fuels”. No harmful effects identified on gas turbine or other components. NOx-emission slightly h g h at present for some fuels, but solutions available . Emissions of HC very low and emissions of dioxins below detection level also for chlorine rich fuels. The IGCC is competitive with conventional biomass technology for condensing applications at today’s fuel cost. The biomass gasification technology is highly suitable for retrofit to existing NGCC.
REFERENCES (1) (2)
Stihl, K., Neergaard, M., Experiences from the Biomass Fuelled IGCC Plant at Varnamo, presented at Biomass for Energy and Industry, 10” European Conference and Technology Exhibition, Wurzburg, Germany, June 1998. Stihl, K., Neergaard, M., Stratton, P., Nieminen, J., IGCC power plant for biomass utilisation Varnamo, Sweden, presented at Developments in Thermochemical Biomass Conversion, Banff, Canada, May 1996.
563
Examining the Thermal Behaviour of Biomass Ash by Various Analytical Techniques S . Arvelakis,' H. Gehnnann,2M. Beckmann, and E.G. Koukios' Bioresource Technology Unit, Dept. of Chemical Engineering, National Technical University of Athens, Zografou Campus, GR-15700 Athens, Greece. Clausthaler Umwelttechnik Institut GmbN, Leibnizstrasse 21 -23, D38678 Clausthal-ZellerJeld, Germany.
'
ABSTRACT: Thermal analysis techniques, such as DTNTGA and sintering tests, in combination with SEM-EDX analysis, were tested in this paper as sources of valuable information towards the ultimate goal of predicting the high temperature behaviour of the biomass ash during thermochemical conversion of various solid biofuels. The critical role of alkali metals and other biomass ash constituents (Si, C1, S) was indicated, along with the potential of leaching as a promising biomass pretreatment for optimal conversion performance. INTRODUCTION & BACKGROUND
Solid biofuels of agricultural origin (e.g., wheat straw, corn cob, olive-oil residue) represent a large part of the available biomass resources. Nevertheless, their use as feedstocks of thermochemical conversion systems (combustion, gasification, etc.) has been associated to serious operation problems. In particular, the inorganic content (ash) of these materials tends to sinter and melt at the temperature range of thermochemical reactors, resulting in deposition (Baxter, 1993; Miles et al., 1995; Baxter et al., 1996), fouling (Jenkins et al. 1997), agglomeration (Nordin and Ohman, 1996), corrosion, erosion, and other problems (Benson et al., 1996; Moilanen et al., 1996), thus limiting the efficiency and availability of the conversion system, and leading to uneconomical operation of the biomass conversion units. Therefore, analysing and understanding the thermal behavior of the ash of these biofuels is of great importance for providing the necessary information for ensuring the smooth operation of biomass combustion and gasification reactors (Mann et al., 1991; Folkedahl et al, 1993; Bakker et al., 1997; Benson, 1997). According to the recent literature, the principal ash-forming inorganic constituents of agricultural biomass are alkali metals, and silicon. The former are in the form of inorganic salts, dissolved in the inherent moisture, chemically linked to carboxylic or other functional groups, or as complex ions and chemisorbed material. 5 64
The latter take the form of hydrated silica or deposits on the cell walls. Other inorganic components include C1, P, S, and alkali earth materials, i.e., Mg and Ca, which are present either on the cell walls or as crystalline Ca or Mg oxalates in the cytoplasm. Alkali metals, especially K, exhibit high mobility and tend to react with silica, by breaking the Si-0-Si bond to form low melting point silicates - a reaction talung place at temperatures far below 900 C -,as well as with S to produce alkali sulphates. Chlorine acts as a facilitator of these reactions by increasing the mobility of potassium, which is usually present as KC1. Potassium chloride is among the most stable at high temperature, gas-phase alkali-containing chemical species, while the amount of Cl in the fuel often dictates the amount of the alkali potentially vaporised during combustion or gasification. Calcium has also the tendency to react with S to form sulphates, but the lower mobility of calcium in combination with the low S content of most biomass types do limit the significance of this type of problem. The alkali silicates and sulphates so produced have melting points as low as 700 C, and tend to deposit on the reactor walls or on the heat exchange surfaces, when conventional grate-fired systems are used; whereas, in the case of fluidised bed reactors, they significantly contribute to bed sintering and defluidisation of the inert material through the development of a sticky deposit on the surface of the bed particles. Th~sthermal behaviour of the biomass ash results in a number of technical problems, such as lowering of the heat transfer coefficient value, restricting the gas flow through the reactor due to increased deposits, and an eventual decline of the conversion efficiency. Large-scale fouling, deposition and/or agglomeration can lead to unscheduled plant shut-downs and subsequent deterioration of the plant economics. The object of this paper is the application of thermal analysis techniques, such as DTNTGA, and sintering tests performed in a muMe furnace, in combination with SEM-EDX analysis techniques, to deepen ow understanding of the high-temperature of the ash of various solid biofuels.
MATERIALS & METHODS
In this work, we have used ash samples form three different agricultural biomass fuels, wheat straw, olive-oil residue, and corn cob. Experiments were performed on ash from untreated biomass feedstocks, as well as from biomass samples pretreated according to two different pretreatment methods, i.e., fractionation (Arvelakis et al., 1996) and leaching (Jenluns et al., 1996; Turn et al., 1997). Biomass fractionation consisted in the mechanical separation of each raw material, following its milling in a hammer mill, in two fractions according to particle sue, and study of the thermal behaviour of the coarse fraction, i.e., the one with average particle size greater than 1 mm, which in all cases accounted for 50-80% (w/w. dry basis) of the initial biomass. Leaching consisted in the treatment of the same materials with water at room temperature in order to reduce their ash content and/or change the ash chemistry. The experimental methodology of the two pretreatment techniques used in t hs work is described in detail elsewhere (Arvelakis et al., 1996). This study has focused on the following 9 samples (symbols in parenthesis): (1) Neat wheat straw (WS). (2) Coarse fraction of wheat straw (WSF). (3) Leached coarse fraction of wheat straw (WSL). (4) Neat olive-oil residue (OR). 565
Coarse fraction of olive-oil residue (OW). Leached coarse fraction of olive-oil residue (Om). Neat corn cob (CC). Coarse fraction of corn cob (CCF). Leached coarse fraction of corn cob (CCL). Two different thermal methods were used for the analysis of the hightemperature behavior of the various ash samples, i.e., simultaneous DTNTGA using a NETSCH STA 429 therrnal analysis instrument, and sinterhg tests performed in a simple muffle furnace, both in combination with a SEM/EDX technique for the elemental analysis of the thermally treated ash samples, using a JEOL 6300 scanning microscope. In the case of the DTA/TGA method, the ash samples were placed in a kaolin sample holder and gradually heated at a lO-C/min rate up to the level of 1000 C. The STA instrument was connected to a PC to generate plots of weight loss and temperature vs. time. For the sintering tests, the ash samples were placed into 25-mL porcelain crucibles and gradually heated from 600 to 1000 C, at 50-100 C intervals, and for periods of 1 h each time, while the physical state and weight of the samples were monitored at the end of each time interval. The thermally treated ash samples generated in both cases were further examined using a SEWEDX elemental analysis technique to determine surface morphology and chemical composition of the heated ashes, thus permitting a more distinct view of the transformations caused by the complex physicochemical interactions among the main ash-forming inorganic elements.
(5) (6) (7) (8) (9)
Table I: Elemental analyses of various biomass ash samples (% w/w, dry ash basis). Ash
K20 NazO
ws
14.1 13.1 8.3 32.2 36.7 5.4 17.3 38.6
WSF WSL OR OW ORL
cc
CCF
4.3 3.3 2.2 8.9 3.9 3.1 1.5 1.1
CaO
MgO
SiOz
14.4 4.4 11.5 21.3
4.3 0.9 3.2 7.9 3.3 4.6 2.2 1.4
39.2 38.5 48.0 32.6 34.3 22.1 23.1 18.5
11.8 20.2 8.1 0.95
A1203 SO3 3.9 1.7 1.7 6.0 2.5 3.4 4.6 0.32
C1
5.3 5.9 1.4 5.0 0.7 2.9 6.1
2.7 3.0 0.1 1.4 1.4
ND
ND
0.0 1.6
RESULTS & DISCUSSION The composition of the nine ash samples is presented in Table 1, whereas Table 2 summarises the results of characterisation of the corresponding biomass fuels by proximate and ultimate analysis, as well as measurement of their higher heating ( H H V ) values. From these Tables, we can see that leached biomass samples contain significantly less ash, and particularly less alkali metals, chlorine, and sulphur, in comparison to the non-leached ones. On the other hand, the use of fractionation can also significantly decrease the amount of ash in the coarse fractions but, at the same time, it leads to substantially increased relative amounts of alkali metals, chlorine and sulphur contained in the pretreated samples. 566
7.5 5.0 77.2 17.7 0.6 478 6.0 0.3 0.8 39.5
6.5 7.5 76.0 16.5 0.8
43.7 5.1 0.4 0.4 42.0 18.9
Moisture
Ash
Volatiles
Fixed carbon
N
C
H
S
c1
0
HHV (kJ/kg)
19.7
WSF
ws
Characteristic
20.2
42.3
0.2
0.0
4.4
48.6
0.3
14.6
81.8
3.6
15.2
L
ws
19.8
38.8
37.0 19.8
0.2
0.3
5.8
46.2
37.9
4560
0.2
0.2
20.8
0.0
5.4
6.1 0.2
46.1
0.4
17.6
53.0
1.1
1.o 51.3
15.1
80.4
2.0
1.6 83.3
9.0
cc
6.6
ORL
19.5
77.9
2.6
8.8
ORF
0.2
0.3
5.9
50.7
1.4
23.4
72.0
4.6
5.5
OR
19.4
18.7
44.6
0.1
0.2 44.7
0.0
5.9
6.2
0.0
48.6
0.0
13.5
86.5
0.8
5.4
CCL
46.8
0.3
16.6
81.5
1.8
10.1
CCF
Table 2 Results of proximate and ultimate analysis of the biomass fuels used in this work (% wlw, dry basis).
D T f l G A TESTS The DTNTGA curves clearly show a melting behavior of wheat straw ash in the range 750-900 C, while leaching seems to only marginal improve the thermal behaviour of the particular ash sample. A minimum of mass loss is observed for all three straw ash samples; this can be explained as a result of the rapid sintering-melting reactions among ash constituents, which prevent the vaporisation of the volatile elements via the creation of molten silicate layers. Inspection of the samples at the end of their heating process revealed that, in all cases, the straw ash was transformed from an initially loose, granular material into a hard-structured one. This effect was more profound for the nonleached straw ash samples. The neat and the fractionated olive residue ash samples performed in a similar to wheat straw way during the DTA/T.GA tests. On the contrary, the leached olive residue sample exhibited a totally different behaviour: its DTA line revealed a significant endothermic phenomenon occurring at the range of 750-850 C, which can be attributed to the dissociation of alkali carbonates, whereas no melting behaviour was observed up to the temperature level of 1000 C. Inspection at the end of the test confirmed that the heated material had kept its initial, granular form. Similarly to the straw and olive residue samples, the untreated corn cob ash started to melt around 800 C, whereas its fractionated sample at a much lower level, i.e., 600 C, and the leached one at a much higher one, i.e., 800 C.
Table 3 Sintering tests of the various biomass ash samples. Initial Condition (600 C) loose, grey loose, grey-black loose, whiteorange loose, red loose, white-red loose
Interim Condition (750 C) sintered, grey partially fused, light grey partially sintered, white sintered, light green heavy sintered, green loose
Interim Weight Loss (%, dry basis) 6.9
cc
loose, red
loose, red
10.4
CCF
loose, white-grey loose, grey
fused,
23.7
dark blue loose, light grey
14.0
Samples
ws WSF WSL OR
ORF ORL
CCL
4.2
9.5 11.6 ND ND
568
Final Condition (1 000 C) fused, grey fused, white-yellow sintered, white
Total Weight Loss (“h,dry basis) 11.0
fused, dark green fused, green loose
21.2
light sintered, dark red fused, blue-purple partially sintered, light grey
14.7
8.3 11.3
25.2 7.0
33.8 14.0
ASH SINTERING TESTS The results of the ash sintering tests are presented in Table 3. These observations are in good agreement with those of the DTA/TGA tests. Thus, all straw ash samples exhibited similar eutectic behaviour, starting at ca. 750 C, with the ash of the leached samples showing a lower tendency for melting, and all three samples being finally transformed from a loose-grain material into a sintered or fused, hard-structured one. This transformation reached an end at ca. 1000 C, where the ash of the non-leached samples was transformed into a very hard material, whereas the ash of the leached sample sintered. In all cases, a mass loss of approximately 10 % was observed during the heating process. The mass-loss rate of the non-leached samples appeared to be constant during all stages of the heating process, whle that of the leached ones reached its highest value below 750 C. The olive residue and corn cob ash samples performed in the same way as straw, although the mass loss during the thermal treatment was found to be significantly lower with thls test, and the melting behaviour appeared to be more pronounced in the case of the coarse fraction. On the contrary, the appearance of the leached olive residue ash sample ash remained intact after this heating process.
SEM-EDX TESTS The objective of the analysis of the thermally modified ash samples from the above two treatments by the SEM-EDX technique was to identify the main elements participating in the formation of the transformed ash structure and possibly obtain a better understanding of the phenomenon, by a correlation of the SEM-EDX results (presented in Tables 4-6) with the elemental analysis of the biomass ash (Table 2). Both thermal treatments appeared to have similar effects on the straw ash surface chemistry. The main inorganic elements constituting the surface structure of the heated samples were in all cases, including the leached samples, K and Si, followed by Ca. Other components of the modified ash structure included small amounts of S and C1, and trace amounts of Na, Mg, Fe and P. This composition is a strong indication for the formation of a complex K-Ca silicate by a complex series of reactions taking place among different inorganic ash constituents at elevated temperatures. The results of the SEM-EDX elemental analysis of the heated straw ash samples (Tables 4-6) are in good agreement with the chemical composition of the initial straw ashes, as presented in Table 1.The main elements participating in the ash structure (Si, K, Ca, C1, S) are those identified by several authors (see Introduction) as the main reason for problem-causing thermal behaviour during biomass conversion at elevated temperatures. According to the SEM-EDX analysis of the heated neat olive residue ash samples, the main elements participating in the hard, f b e d structure are alkali metals (up to 50 % K), as well as Si, Ca, S and C1. On the other hand, Ca and Si constitute the main structure of the leached olive residue ash sample after the heating process. In the case of corn cob, the results of the sintering tests (Table 3), in combination with those from the SEM-EDX analysis of the thermally treated ash samples, are in full agreement with the results obtained by the DTA-TGA thermal analysis method, clearly showing that meltingkntering effects begin at around 750 C or even at lower temperatures.
5 69
Table 4 SEM-EDX elemental analysis of thermally treated (DTA/TGA) biomass ash. Samples WS WSF . -.
WSL OR ORF --ORL CC CCF CCL
K20
Na20
CaO
48.3 81.0 56.7 46.3 71.0 9.1 27.1 77.2 57.3
3.0 4.0 0.2 0.0 0.0 0.0 0.0 0.2 0.2
20.4 17.1 9.2 28.4 20.1 35.8 34.0 12.2 12.2
MgO Si02 3.1 13.3 9.0 1.2 2.3 14.4 8.3 3.1 1.2 2.0 22.0 22.0 6.0 13.1 4.4 2.3 2.3 14.4
FezOp 3.4 0.7 0.5 5.4 0.5 9.9 6.8 0.5 0.5
A1203 2.2 0.7 0.0 2.2 0.9 0.2 3.4 0.0 0.0
SO3 2.6 0.5 7.2 2.6 0.5 0.0 2.9 3.2 8.2
P205
3.7 2.0 4.5 3.7 2.0 0.9 6.6 2.5 4.5
C1 0.0 1.9 0.4 0.0 1.7 0.0 0.0 0.4 0.4
Table 5 SEM-EDX elemental analysis of the sintered biomass ash samples at 750 C. Samples
K
Na
Ca
Mg
Si
A1
Fe
P
SO3
C1
WS WSF WSL
33.7 46.9 27.9 55.6 68.2 3.73 49.9 69.0 42.6
0.0 0.0 0.2 0.0 0.9 0.7 0.5 0.0 1.4
24.5 14.5 1.6 25.0 20.9 68.5 20.5 1.3 21.5
1.1 1.1 3.2 0.9 1.7 4.7 0.8 0.7 4.1
26.5 15.6 37.3 3.8 1.6 4.52 14.4 14.0 17.5
0.1 0.7 11.5 0.0 0.7 2.0 0.9 0.1 0.6
2.1 1.1 13.3 4.5 0.0 6.7 0.7 0.5 0.3
2.8 2.4 0.18 3.1 2.5 3.1 0.6 5.7 3.4
3.5 8.5 3.0 5.5 2.1 4.8 11.3 1.2 6.9
5.9 9.2 0.8 0.9 1.5 0.3 0.8 7.6 1.4
OR OW ORL CC CCF CCL
Table 6 SEM-EDX elemental analysis of the sintered biomass ash samples at 1000 C. Samples WS WSF WSL OR ORF ORL CC CCF CCL ~
~~
K20 16.4 16.4 1.01 26.2 ~22.4 0.3 10.6 69.8 0.6
Na20 3.9 6.9 0.6 4.8 1.0 0.0 4.8 5.2 0.4
CaO 13.1 9.5 33.8 26.3 14.9 59.2 30.7 ND 49.5
MgO 4.4 3.9 6.4 12.7 0.6 7.0 12.4 1.8 17.4
Si02 60.4 60.7 55.9 13.6 46.4 17.1 21.4 22.7 17.5
570
A1203
ND 0.0 0.0 7.6 12.5 3.6 5.1 ND 0.7
Fe20
P20
3
5
1.7 0.3 1.1 5.1 1.6 9.3 4.3 0.3 0.6
0.2 ND ND 3.6 ND ND 6.7 ND 3.5
SO3 ND 2.3 1.4 ND 0.8 3.1 5.5 ND 9.5
Cl ND 0.0 0.0 0.1 0.0 0.4 0.2 0.3 0.3
CONCLUDING REMARKS The application of the two methods for the study of the thermal behaviour of the biomass ash gave comparable results. Alkali metals, especially K, in combination with Si, S, and C1, were shown to be the main constituents of the quasi-eutectic state of the particular ash types examined. The use of biomass leachmg as a pretreatment has, in general, a positive effect on the thermal behaviour of biomass ash, the degree of which significantly depends on the biomass concerned. Specifically, the use of leaching proved to substantially improve the thermal behaviour of the olive residue ash, whereas similar action was rather limited in the case of corn cob, and marginal in the case of wheat straw. These effects can be explained by the composition of the initial biomass ash, as leaching has a more pronounced effect on alkali metals and other chemical components promoting ash melting behaviour.
ACKNOWLEDGEMENT The research collaboration of the two team was facilitated through a post-graduate research grant to S . Arvelakis by the European Commission.
REFERENCES 1. Arvelakis, S., Taralas, G., Koullas, D. P., & Koukios, E. G., (1997) Upgrading Agricultural Residues as Feedstocks of Electricity Generation by Gasification. Proceedings 3rdBiomass Conference of the Americas, Montreal, p.607. 2. Bakker, R. R., Jenkins, B. M., William, R. B., Carlson, W., Duffy, J., Baxter, L. L., & Tiangco, V. M. (1997), Boiler Performance and Furnace Deposition During a Full Scale Test With Leached Biomass. Proceedings 3rdBiomass Conference of the Americas, Montreal, p.497. 3. Baxter, L. L. (1993), Ash Deposition During Biomass and Coal Combustion: A Mechanistic Approach, Biomass & Bioenergy, 4(2),85. 4. Baxter, L. L., Jenkins, B. M., Miles, T. R., Miles, T. R., Jr., Milne, T., Dayton, D., Bryers, R. W. & Oden, L. L. (1996), Inorganic Material Deposition in Biomass Boilers. Proceedings 9" European Bioenergy Conference, Copenhagen, 2, 1114. 5. Benson, S. A. (1998), Ash Formation and Behavior in Utility Boilers. Newsletter published by Microbeam Technologies Inc. 6. Benson, S., Steadman, E. N., Zygarlicke, C. J. & Erickson, T. A. (1996), Ash Formation, Deposition, Corrosion, and Erosion in Conventional Boilers. Applications of Advanced Technology to Ash-Related Problems in Boilers, Plenum Press, New York, p. 1. 7. Folkedahl, B. C., Steadman, E. N., Brekke, D. W. & Zygarlicke, C. J. (1993), Inorganic Phase Characterization of Coal Combustion Products Using Advanced SEM Techniques. Proceedings of the Engineering Foundation Conference, Sollhull, England, June 20-25, p.399. 8. Jenkins, B. M., Bakker, R. R. & Wei, J. B. (1996), On the Properties of Washed Straw. Biomass & Bioenergy, 10(4), 177.
57 1
9. Jenkins, B. M., Bakker, R. R., Williams, R. B., Baxter, L. L., Turn, S. Q., Thy, P., S h e , M., Lesher, C., Sclippa, G. & Kinoshita, C. (1997), Measurements of the Fouling and Slagging Characteristics of Banagrass (Pennisetum Purpureum) Following Aqueous Extraction of Inorganic Constituents. Proceedings 3'd Biomass Conference of the Americas, Montreal, p. 705. 10.Mann, M. D. & Galbreath, K. C. (1991), The Role of Ash Chemistry and Operatin Parameters on Ash Agglomeration and Deposition in FBC Systems. Proceedings 5 Engineering Foundation Conference on Inorganic Transformations and Ash Deposition During Combustion, Palm Coast, p.773. 11.Miles, T. R., P. E., Miles, T. R., Jr., Baxter, L. L., Bryers, R. W., Jenkins, B. M. & Olen, L. L. (1995), Alkali Deposits Found in Biomass Power Plants, Volumes I, 11, Summary Report, NREL Subcontract TZ-2- 11226-1. 12.Moilanen, A., Nieminen, M., Sipila, K. & Kurkela, E. (1996), Ash Behaviour in Thermal Fluidised-Bed Conversion Processes of Woody and Herbaceous Biomass. Proceedings 9" European Bioenergy Conference, Copenhagen, 2,1227. 13.Nordin, A. & Ohman, M. (1996), Agglomeration and Defluidization in FBC of Biomass Fuels-Mechanisms and Measures for Prevention. Applications of Advanced Technology to Ash-Related Problems in Boilers, Plenum Press, New York, p.353. 14.Turn, S. Q., Kinoshita, C. M. & Ishimura, D. M. (1997), Removal of Inorganic Constituents of Biomass Feedstocks by Mechanical Dewatering and Leachmg. Biomass & Bioenergy, 12(4), 241.
B
572
Measuring and modelling the gas residence time distribution in biomass furnaces S. Biollaz*, Th. Nussbaumer Laboratory of Thermodynamics in Emerging Technologies, Swiss Federal Institute of Technology, ETH Zurich, CH-8092 Zurich, Switzerland C . H. Onder Measurement and Control Laboratory, Swiss Federal Institute of Technology, ETH Zurich, CH-8092 Zurich, Switzerland
ABSTRACT: Measuring and modelling the gas residence time distribution (RTD) in furnaces is difficult but reveals important i nformation concerning the mixing quality. The present study has aimed at a development of a measuring system and the application within various secondary combustion chambers (S CC) of biomass furnaces. Important process parameters, the so called three T’s (time, turbulence, temperature) were determined with the developed measuring system. With certain model assumptions for the reactor dynamics, the degree of mixing and the mean residence time are calculated from the measured RTD. The determination of the fluctuating process parameters is possible with a time resolution of 15 seconds. The carbon monoxide oxidation was chosen as a case study reaction. With physicochemical models and the determined process parameter mentioned above the CO oxidation was simulated and compared to the measured CO emission. While the measurement system does yield a rough description of the process, the determination of the efficiency of the mixing on a molecular level remains difficult. The significant differences of CO measured versus CO simulated is explained by this fact. Nevertheless, the results of the experiments and the simulation confirms earlier experiences of the need of an early and complete mixing of the combustion air and a high temperature for low CO emissions. Additionally, the applicability of the proposed measurement and data processing strategy has been shown.
INTRODUCTION Emissions of carbon monoxide (CO) or nitric oxides (NO,) are significantly effected by the operation of furnaces. The level of carbon monoxide emissions for example is a good indicator for the completeness of the combustion process. To build combustion chambers for biomass furnaces with a complete combustion or low-CO emissions, respectively, an improved understanding of the physicochemical aspects of the oxidation of CO is mandatory. This study is a contribution towards this direction. As Corresponding author: Paul Scherrer Institut, Villigen, Switzerland,
[email protected]
573
the present study also deals with methodological aspects of measuring and modelling the gas residence time distribution (RTD) in furnaces the methodological results can be transferred to fundamentally comparable problems i.e. NO, modelling, mixing quality of gasification air in biomass gasifiers, etc. It is well known that reaction only takes place if the reactants are mixed on a molecular level. Historically, combustion research was focussed on the conversion of carbon to COz and heat. The demanded degree of mixing of the reactant streams was therefore moderate. In view of the goal of low emissions the state of mixedness becomes a decisive factor. In general the so called three T's (time, turbulence, temperature) determine the emission level. In this study the oxidation of CO is chosen as a case study reaction. Figure 1 shows typical characteristics of four different biomass furnaces for the CO emissions as a hnction of the excess air ratio h. Furnace ,,d"reaches much lower CO emissions in the range of 20 mg/m3, than furnace ,,a" with 2000 mg/rn3".In addition the CO minimum in furnace ,,d"is reached with a much lower excess air, i.e. hop,of 1.3 compared to furnace ,,a" with hop,of 2 . Furnace ,,b"and ,,c" are intermediate systems. The measured characteristics for all furnaces can be explained with the assumption that below hop,there is locally a lack of oxygen. As there is overall an excess air the mixing quality of combustion air and combustible gases must be improved for a further reduction of the CO emissions in the secondary combustion chamber (SCC). This finding is confirmed comparing furnace ,,c" and ,,a' where furnace ,,d" has an improved mixing of combus tion air. 100000 Img/m31
co
10000
1000
I00
10 0
1
2
3
Excess air ratio
4
(-1
5
k
Fig. I Emissions of carbon monoxide (CO at 11 Vo1.-% 0,) in function of the excess air ration h [I]. Legend: a Manual wood fung, poor standard b Manual wood fuing, high standard c Automatic wood fuing, high standard (1990) d Automatic wood f h g , high standard with good mixing of combustion air (since 1995)
The operating conditions in the combustion chamber can be described by the process parameters which are gas temperature, excess air, mean residence time, and the
574
state of mixedness. With this approach various furnaces can be compared easily and model calculations become possible. Figure 2 shows the result of such a model calculation. The CO decomposition (CO/COo) is calculated as a function of temperature (T), mean residence time ( I ) , and the state of mixedness which is described by the Bodenstein number (Bo). The model is as follows:
co -= COO
4.9 Bo
(l+q)2e
I-q)
2
-(I-ql2e
Bo --.( Itq)
2
It is assumed that the combustion air and the combustible gases are evenly mixed before entering the SCC. In table I the kinetic data are compiled for Esq. 3. These kinetic data are valid for the temperature range from 567 up to 2087 "C.
Table I Kinetic data for the modelling of CO decomposition [2] parameter 0 2 reaction order HzO reaction order activation energy reaction rate constant
abbreviation c1
dimension [-I
values 0.5 0.5 125'500 1.3 1014
[-I
P
[J/mol]
EA
[ cm3/rnol /s
From Figure 2 it is clearly visible that once the reactants are mixed on a molecular level and if the temperature is sufficiently high (T > 800 "C)the reaction time is quite short, i.e. below 1 s mean residence time to get a high CO decomposition (lo-").
10-2
10"
1o
-~
10'~ 10.6
10.'
lo-* 1 0q9
lo-'" 0.0001
0.001
0.01
0.1
1
10
mean residence timet [s]
Fig. 2 Model calculation for the CO decomposition. Model parameters: Mean residence time, state of mixedness (described with Bodenstein number Bo) and temperature. 0 2 concentration 2 Vol% and H 2 0 concentration 15 Vol%. Kinetic data see Table 1
575
For the understanding and interpretation of mixing processes or turbulence the residence time distribution (RTD) contains valuable information. Besides the RTD for flue gases there are two additional factors which influence the reactor behaviour [3], namely the degree of segregation and the earliness of mixing. These three parameters describe the state of mixedness. Considering separate reactant streams the time required for mixing is usually much longer than the time required for the reaction. Thus the chemical kinetics become negligible. The RTD can be determined experimentally and is typically used in chemical reactor engineering for the determination of the reactor behaviour. It uses the injection of a tracer as an excitation of the reactor in steady state operation. At the outlet of the reactor the transient response of the tracer concentration is measured. This method has been used for the experimental investigation of furnaces by various authors [4-81. The authors cited used for their investigations simplified experimental setups or did not have to consider the dynamics of the measurement system. In this study a measurement system is presented which is applied under real conditions on biomass furnaces and the measurement dynamics for the RTD determination are fully taken into account.
EXPERIMENTAL SYSTEM The motivation for the experimental fixmace on the one hand was to have a test rig for the validation of the to be developed measuring system. On the other hand the motivation was to obtain model-based guidelines for the design, operation, and improvement of biomass furnaces. As a case study reaction the CO oxidation was chosen and the modelling concerned the CO oxidation in the secondary combustion chamber (SCC) [4]. The experimental furnace was to yield measurement data which could be used to validate various physicochemical models. Therefore it was necessary to vary the process parameters such as gas mixing, mean residence time, and temperature over a wide range in the SCC. Figure 3 shows a schematic diagram of the experimental furnace. It is a commercially available 40 kW under stoker burner for wood chips. It corresponds to a regular furnace for heating purposes. er
U
U
PCC: primary combustion chamber
W'
-
-
---
-1-
-
-
-2I
I
SCC: secondary combustion chamber (after burner)
Fig. 3 Automatic wood firing as experimental biomass furnace
576
In the section between the primary combustion chamber (PCC) and the heat exchanger, four different types of SCC were installed. Their RTDs should be close to those of PSR, PFR, and an intermediate type. Two examples of the SCC investigated are shown in Figure 4. The SCC of type I has a length-to-diameter ratio of 8.5. It should therefore behave similarly to a PFR. With the frequent change of the flow direction, additional axial mixing must be expected during flow. Type I1 has a much smaller length-to-diameter ratio of 2.5. Therefore the RTD of this type is not expected to be explainable with either PFR or PSR. Additionally, the mixing of the combustible gases with the secondary air is driven only by the impulse of the secondary air and is not supported by geometric effects as in type I. Due to the design of the SCC investigated it is not possible to distinguish between the mixing types, i.e. mixing of air with combustible gases and self-mixing. Additional sampling points for the measuring system between SCCi, and SCC,, would yield the needed information. Nevertheless, it is obvious that the different furnace types will have different effects on the two mixing types.
sec
primary air
I 44 I primary air
primaryair
wood
I 44 Iprimaiyair wood
Fig. 4 Secondary combustion chamber of types I (left) and I1 (right) The measurement system developed, as shown in Figure 5, is introduced into the furnace on level SCC,,, so that the ceramic tip is in the gas flow. The key elements of the measurement equipment are a suction pyrometer for the temperature measurement, a wide-range oxygen sensor (BOSCH LSU4) for the fast oxygen measurement, and a thermal conductivity detector (TCD) to measure the transient response of the Helium tracer concentration. For the determination of the transient response of the TCD measurement system itself Helium tracer was injected just at the inlet of the suction pyrometer. In Figure 7 a typical transient response of the TCD measurement system is shown (curve with index SCCi, ,.). The Helium tracer can be injected into the primary as well as into the secondary air flow (see Figure 4). Near the entrance of air into the combustion chamber the tracer is thoroughly mixed with the air. With the solenoid valve it is possible to obtain a reproducible tracer injection. The maximum helium concentration in the flue gas is 2%.
577
furnace
TCD: Thermal CcndwtivilyDcrector LSU:Lambda Sondc Univcr~sl
Fig. 5 Measurement system: suction pyrometer, LSU4 and TCD sensors.
METHODOLOGY FUNDAMENTALS OF RESIDENCE TIME DISTRIBUTION (RTD) In this study the degree of turbulence is described with the residence time distribution (RTD). It is a widely used approach in chemical engineering to describe the state of mixedness in reactors in this way. The two extremes of the state of mixedness are represented by the plug flow reactor (PFR, no mixing) and by the perfectly stirred reactor (PSR, perfectly mixed). The reactant flow in the PFR is neither macro nor micro mixed, whereas in the PSR mixing occurs down to the molecular level, thus both macro and micro mixing take place (see Figure 6). A variety of real flows can be characterised by series, parallel or loop connections of PFR and PSR. Additionally there exist other models such as the dispersion model (dispersed plug flow) which allows to model mixing conditions between the two extremes of PFR and PSR. Various system excitations are available to experimentally determine the RTD or, more general, the reactor dynamics. Commonly used excitations are impulse and step signal, periodic and random signal. The basic idea is to excite the system and determine how it reacts to this excitation. The transfer function G(s) describes the linear reactor dynamics and thus is independent of the stimulus [9,10]. In frequency domain a series connection of two reactors is described by the multiplication of their transfer functions, whereas a parallel connection is represented by a sum of the corresponding transfer functions. The use of the transfer function G(s) for the evaluation and interpretation of the state of mixedness is therefore advantageous, The transfer functions of PFR and PSR are as follows [lo]:
G,(s)
= e-*IPFR
578
(4)
In both models the only model parameter used is the mean residence time tpFR and t PSR . Figure 6 shows the reactor dynamics of the PFR and the PSR in the normalised time and frequency domain (dimensionless time 8 = t/ , dimensionless frequency fe = 1/2n;8). In the time domain the step response F(8) as well as the impulse response E(8) (which is the RTD) can be discussed. This type of data presentation is normally used in chemical engineering application. But the same data can also be presented in the frequency domain, the so called Bode plot. This type of presentation allows to identify effects which are not visible in the classical used plot in the time domain. The Bode plot consists of the magnitude IGj we)/ and the phase arg{ G('jw8)). magnitude of frequeoze response
step response I
10'
I
I
I
I
dispersion model h
f
0.4
0.2
I 0
I
I
dispersion mode\
1"
2
dimensionless time t3
3
IOU
10.'
[-I
lo'
dimensionless frequency
impuls response residence time distribution (RTD)
fa [-I
phase offrequenze response IORi
4 7 5n
I
I
-5n dispersion model
PFR
-10K
0
I
2
3
-1511 0
dimensionless time 8 [-I
2
4
6
8
1
0
dimensionless frequency fe [-I
Fig. 6 Reactor dynamics of plug flow reactor (PFR), perfectly stirred reactor (PSR) and dispersed plug flow (dispersion model; parameter value Bo = 8.8) An alternative model for real flows is the dispersion model with the model parameters Bodenstein number (Bo) and mean residence time I . The Bodenstein number which is defined as Bo = uL/D characterises the degree of backmixing during flow. The parameter D is called the axial dispersion coefficient, u is a velocity and L a length. The RTD of the dispersed plug flow model ranges from PFR at one extreme (Bo = =) to PSR at the other (Bo = 0). The transfer function for the dispersion model with closed-closed boundaries is [lo]:
4t
l+s.(7) Bo Figure 6 shows the example of the reactor dynamics of the dispersion model calculated with a Bodenstein number of 8.8 q=
579
SIGNAL PROCESSING The goal of the signal processing is the parameter determination of the state of mixedness (i.e. Bodenstein number, mean residence time, etc.). The identification of the model parameters is done by using an off-line identification scheme. The identification can take place either in the time or in the frequency domain. Since the convolution in the time domain becomes a multiplication in the frequency domain, it is advantageous to execute the identification in the frequency domain to save computational costs. Additionally, the presentation of the results in the frequency domain offers new insights. Independently of the system to be investigated and the measurement strategy used the input and output signals of the system are needed. The frequency response is determined by performing a Fast Fourier Transformation (FFT) on both RTD signals. In this study these are the signals corresponding to position SCCk and SCC,,, (see Figure 4). In order to calculate the output of a combustion chamber model based on the measured input, the FFT of the input signal is first multiplied with the model frequency response of the combustion chamber (see Esq. 8). GSCCo,
(') = GSCC,fi ('1
'
GModeluf Combustion Chamber
('1
(8)
The modelled output signal is then calculated by performing the inverse FFT on the modelled output spectrum. The model parameters of the model of the combustion camber are found by fitting the modelled output signal with the output signal measured. The fitting is effected by minimising the least square error.
RESULTS AND DISCUSSION Figure 7 shows a typical result of the procedure explained in the last chapter. The meas. are the signal used for the system excitation is a step signal. SCCi,meas. and SCCOut, measured input and output signals where SCCi,meas. is in fact the transient response of the measurement system. Taking these two signals, the reactor dynamics of the SCC can easily be determined. For the SCC of type I1 an example of a RTD modelled is shown in Figure 7. The model used is the dispersion model (see Esq. 6). The values of the model parameters determined are a Bodenstein number of 8.8 and a mean residence time of 0.6 s. It clearly shows that the model for the RTD explains the frequency response measurement up to a frequency of 2 Hz. At the frequency of 2 Hz the signal-to-noise ratio of 100 is reached. Any mixing processes which affect the transfer fbnction above this frequency cannot be identified. The presentation of the result in the frequency domain offers an additional advantage. The reactor behaviour modelled (Comb. Chamber,d.) can be directly compared with the behaviour calculated (Comb. Chambercd,.). The calculated frequency response of the reactor is the ratio of the measured output and input spectra. This comparison is not possible in the time domain due to the measurement noise at high frequencies. This noise is amplified by the compensation of the measurement dynamics and thus no useful presentation of data is possible in the time domain. Having the model parameter determined the reactor dynamics can be calculated and be plotted in the time and frequency domain (see example with Bo = 8.8 in Figure 6 ) .
580
lo2 I
1.O
-
0.8 A 0.6
Comb. Ch?mbercalc,
1
Comb. Chamber
1o'
h
0.4
1o 2
0.2 0.0
0
1
2
3 time t [s]
4
1o4
5
10'
1O0
1.~ 0'
frequency f [Hz]
s c!
7
lOrr 5n Comb. Chambercalc,
1
i
Y
0
1
2
3
4
0
5
time t [s] -
-
SCCi, meas.
see,,,
2
4
6
8
1
0
frequencyf[Hz]
SCCout, meas. mod,
- Comb. Chambercalc.= SCCout,meas.
.. - Comb. Chambermod,
'
SCCin, meas.
Fig. 7 Measured and modelled RTD for SCC type 11. Presentation of the data in the time domain (left) and in the frequency domain (left). Dispersion model was used for SCC modelling (see Esq. 6). Model parameter determined: Bo = 8.8 [-I, = 0.6 [s] Legend: Measured input signal, transient response of measurement system SCCm,meas. SCCOUI,meas. Measured output signal Modelled output signal with Esq. 8 sccout,mod Modelled reactor behaviour of combustion chamber Comb. Chambercal,, Calculated reactor behaviour of combustion chamber Comb. Chambermd.
In Figure 8 and 9 two examples are shown for the continuous measurements and determination of process parameters such as temperature, excess air, state of mixedness and mean residence time. The measurement of temperature, excess air, and helium concentration takes place at the outlet of the SCC (SCCou,). The CO emissions and the excess air are measured separately in the chimney of the biomass furnace. This measurement equipment works simultaneously with the RTD measurement, but with a higher time resolution. The Helium injection occurs via secondary air flow to determine the state of mixedness in the SCC (see Figure 4). Square waves with a period of 30 s are used to excite the system during 5 minutes. Therefore 10 step responses in both directions result, such that every 15 seconds new measurement data of the state of mixedness are provided. Each step response is evaluated individually as shown in Figure 7 and thus the model parameters are determined off-line for every 15 sec. This corresponds to the time resolution of the RTD measurement system.
581
10000
lww 1
. . . . . ..:.. . . . . . . .:. . . . . . . . . . . ;. . . . . . . . . 1000
100
1;
::be. I # . . i ‘“‘. # t i I000
20
9SO
I
............................
. . .. . . . . .
. ...#..
...
.
. . . . .
.....................
0 4s
0 3s 00
I0
20
30
40
I .o
50
time [minutes]
1.5
i* “:!“ad% ””
2.0
2.5
excess air h [-I
3.0
Fig. 8 Continuous measurement of CO emissions and process parameters (T, h, Bo, i ) as a function of time for the SCC of types 11. Correlation of CO emissions, temperature and NO, as function of excess air h.
For the SCC type I1 Figure 8 shows a large variation of the CO emissions within a very short time (factor of 10 in magnitude). Since the excess air as well as the temperature are quite constant, the measured CO variation could be expected to be explainable with variations in the state of mixedness, i.e. Bodenstein number and mean residence time. But the measurement shown here does not allow a conclusion in this direction. For the SCC type I the continuous measurements are shown in Figure 9. All the process parameters measured are comparable to the measurements in Figure 8, besides the determined Bodenstein number. For SCC type I the Bodenstein number is 5 whereas for the SCC type I1 the Bodenstein number is 9, both averaged over 5 minutes. SCC Type I has a smaller Bo and therefore a more intensive axial mixing. Reminding the CO decomposition shown in Figure 2 the remaining CO concentration for both SCC should be much lower for the temperature and mean residence time measured. The main difference between SCC I and I1 is probably the earliness of mixing and the intensity of that mixing which would explain the smaller Bo for SCC I. This difference might be the cause for the different CO emissions measured. Unfortunately the determination of the efficiency of the mixing on a molecular level remains difficult. In this study just a few typical examples of a much more detailed study are presented [4]. In that study measurements with various types of SCCs show the influence of the state of mixedness on the CO emissions, in particular the earliness of mixing and therefor confirms the general trends presented here.
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loo00
950
30
I5
1000
0 55
0 45
0 35 00
I0
10
30
40
50
10
15
20
25
3.0
excess air h [-]
time [minutes]
Fig. 9 Continuous measurement of CO emissions and process parameters (T, A, Bo, ) as a function of time for the SCC of types I. Correlation of CO emissions, temperature and NO, as function of excess air h.
CONCLUSION The measuring system developed was successfully applied within various secondary combustion chambers (SCC) of biomass furnaces to determine important process parameters. With certain model assumptions for the reactor dynamics, the degree of mixing and the mean residence time are calculated from the measured residence time distribution (RTD). The determination of the fluctuating process parameters is possible with a time resolution of 15 seconds. While the measurement system does yield a rough description of the process, the determination of the efficiency of the mixing on a molecular level remains difficult. Due to the lack o f satisfactory data on the mixing performance, such as the mixing time of the reactants, the earliness of mixing, and the mixing on a molecular level (degree of segregation), the CO oxidation measured cannot be completely explained. The calculations shown in Figure 2 concerning the CO decomposition show the importance of a complete mixing of reactant streams. Otherwise unreacted CO from the fuel bed would dominate the CO emissions measured. Therefore detailed information on these mixing effects is still needed. To determine the degree of segregation, i.e. the bypass stream of unreacted CO, a much faster and much more sensitive sensor is required. Therefore, in this application, it would be helpful to have an independent method for the determination of the degree of segregation.
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Nevertheless, the results of the experiments and the simulation confirms earlier experiences of the need of an early and complete mixing of the combustion air and a high temperature for low CO emissions. Additionally, the applicability of the proposed measurement and data processing strategy has been shown. The methodological results can be transferred to comparable macro mixing problems.
ACKNOWLEDGMENTS This work was funded by the Swiss Federal Office of Energy and was carried out at the Laboratory of Thermodynamics in Emerging Technologies (formerly Laboratory for Energy Systems) in the Bioenergy Group.
REFERENCES Nussbaumer Th. (1994) Wood Combustion, International Conference on Advances in Thermochemical Biomass Conversion, A.V. Brigwater (Ed.), Blackie Academic &Professional, pp. 575-589, 1994, ISBN 0 7514 0171 4 2. Howard J.B., Williams G.C., Fine D.H.. (1973) Kinetics of Carbon Monoxide Oxidation in Postflame Gases. 14IhSymposium (Int.) on Combustion, Combustion Institute Pittsburgh 3. Levenspiel 0. (1972) Chemical reaction engineering. 2nd ed., John Wiley & Sons, New York 4. Biollaz S. (1 997) Messen des Yemeilzeitspektrums in der Nachbrennkammer von Holzfeueungen zur Modellierung der Kohlenmonoxid-Oxidation.Ph.D. Thesis. (ETH) Zurich 5. Seher A, Dorn I. H., Engelhard H., Gruber W. and Maichel K.W. (1988) Radioaktive Indikatormethode in der chemischen Verfahrenstechnik. Chem.-Ing.Tech. 60, Nr. 9, S. 691-698 6. Malek Ch. (1993) Zur Bildung von Stickstoffoxid bei einer Staubfeuerung unter gleichzeitiger Berticksichtigung des Ausbrandes. Ph.D. thesis, TU Clausthal 7. Nasserzadeh V., Swithenbank J., Lawrence D., Garrod N. and Jones. B. (1995) Measuring gas-residence times in large municipal incinerators by means of pseudo-random binary signal tracer technique. Journal of the Institute of Energy, September 1995,68, pp. 106-120 8. Couturier G., Kallner P. and Berger N. (1996) Ignition and Flame Stability Limits for 50 kW Pulverized Wood Flames. In: First European Conference on Small Burner Technology and Heating Equipment, 25-26 September 1996, Zurich, Switzerland, pp. 155-163 9. Oppenheim A.V. and Willsky A S . with I.T. Young (1983) Signals and Systems. Prentice-Hall, Englewood Cliffs, NJ 10. Wen C.Y. and Fan L.T. (1975) Models forjlow systems and chemical reactors. Marcel Dekker, Inc. New York 1.
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A general model for the investigation of packed bed combustion with respect to wood C. Bruch, Institute of Energy Technology, Swiss Federal Institute of Technology, Zurich, Switzerland B. Peters, Institute for Nuclear and Energy Technology, Karlsruhe Research Centre, Germany T. Nussbaumer, VerenumResearch, Zurich, Switzerland
ABSTRACT In the current contribution a general numerical model for combustion in a packed bed is developed and applied to the thermal conversion of a packed bed of wood particles. The simulation method consists of two submodels, which represent the flow of the gas in the void space of the bed and a finite number of fuel particles for which heating, drying, pyrolysis and heterogeneous combustion take place. Thus, kinetic data excluding heat and mass transfer rates is applicable from single particle experiments. The two submodels are coupled by heat and mass transfer. Due to restrictions of memory and computation time, a one dimensional model is chosen for the description of a single particle within the bed. The gas phase is discretized by a finite volume approach on a Cartesian mesh. Because of the representation of the solid phase by different particles, a new approach based on the definition of neighbour particles is chosen to calculate the interaction between &heparticles. By employing an object oriented approach, the numerical treatment of particle processes for a finite number of particles is facilitated. Furthermore, it allows tracking of &heconversion history of each particle and easy extension of the model for further material and kinetic data.
INTRODUCTION Usually, wood chips in automatic furnaces bum in packed bed arrangements, such a! grate and understoker furnaces. Thus, the combustion process takes place in two zones: the packed bed, where the solid fuel is converted into the raw gas and the gas phase. representing the combustion of the raw gas with secondary air. Despite the improvements of the past years, wood furnaces still offer a potential for optimisation with respect to lower emissions and higher efficiencies. However, this requires a deeper knowledge of the processes taking place in a furnace. With increasing computer power, numerical methods, such as Computational Fluid Dynamics (CFD) became a useful tool for the optimisation of combustion processes in conjunction with experiments.
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To investigate flow conditions in the furnace chamber of wood furnaces CFD codes have been applied in the past [ 1,2]. It could be shown, that mixing conditions between raw gas and secondary air still need to be improved to reduce the excess air ratio during the combustion of wood. While for investigation of the gas phase in the furnace chamber commercial CFD codes are available, no general model for the description of the solid fuel conversion in the packed bed exists. However, for optimisation of the flow conditions with respect to reduced emissions, detailed information about the species distribution in the gas phase are necessary as boundary conditions in a CFD code. Moreover the packed bed itself offers a potential for optimisation, e.g. regarding the primary air distribution.
So far boundary conditions for gas phase calculations are taken from measurements or empirical correlation, limiting the application only to specific cases. Therefore the aim of the current project is to develop a numerical model, which predicts the conversion of the solid fuel in the packed bed. The model should take different operating parameters and main fuel properties, such as size and humidity of the fuel particles, into account.
MODELLING PACKED BED COMBUSTION In the recent years different numerical models for the conversion of wood in a packed bed have been presented, e. g. [3-61. Existing models mostly describe the bed as a porous media by an Eulerian approach, with the conservation equations for the solid and the gas phase solved with the same mesh. This approach implies that heat and mass transfer can only be taken into account according to the dimensions of the bed but not within the particles itself. Temperature and species distributions are assumed to be homogenous over the fuel particles. Thus, the influence of the particle dimensions on the conversion process can only be captured by simplified assumptions or macrokinetic data. For general use, the application of macrokinetic data has to be excluded. As has been shown by [7] for the pyrolysis of wood, the influence of the heat transfer in the particle cannot be neglected if particle diameters are in the range of several milimeters. Thus, the assumption of homogenous temperature distributions over the particle is not valid and a different set of macrokinetic data for each particle diameter would be needed to cover the influence. In understoker and grate furnaces the wood particles used are typically sized between 5 and 25 mm, which requires the heat and mass transfer depending on the size of the particles to be taken into account to avoid macrokinetic data.
SEPARATION OF THE PROBLEM The challenge is, to develop a model for the packed bed which is not only resolving heat and mass transfer within the bed,but also transport processes over each particle. Thus, a particle model is combined with the description of the flow through the porous bed. A Langrangian formulation for the solid phase in the bed is chosen here, where each particle is monitored during its conversion process. For implementation of the model an object oriented approach based on the programming language C++ has been chosen. The authors developed the software library TOSCA (tools of object oriented software for continuum mechanics application) [8] which provides all classes required for description of continuum mechanics problems, as e.g. numerical solvers, discretization methods and data bases 586
for property and kinetic data, thus excluding the need for coupling with any other CFD code. Within the code each fuel particle represents a separated object, allowing the single particle model to be run separately from the flow field model. Therefore calculated data for the degradation process of a particle in the bed can be compared to experimental data from thermogravimetric analysis of large wood particles. To reduce memory requirements for the large number of particles, these are grouped together, as shown in Fig.1, in a particle pile, which represents a fuel combined with a certain reaction scheme. All information which is common to a certain group of particles as e.g. material, reaction scheme or ash content is only stored once, whereas information which is specific to each particle as e.g. temperature distribution or particle size is stored separately. Each particle is identified by a number which is used by overlying management classes to arrange particles in a packed bed dependant on a given porosity of the bed. Packed Bed
Particle Pile, Material 1 - - r p p a r t r
Te
L
LParticle n
Particle Pile, Material 2
-e
Particle n+ 1
Fig.I Arrangement of discrete particles to a packed bed.
GOVERNING EQUATION Due to the high number of particles in a packed bed, the model for a single particle hits to be simple. Here, an one dimensional approach has been chosen as compromise between accuracy and computing time. The change of a scalar I$ in time within the particle is influenced by diffusion, convection and source terms. Thus, the energy and species distribution over the particle can be described by the general transport equation
representing an infinite plate, an infinite cylinder and a sphere for n = 0, 1,2. Within the fuel particle, values for the different scalars are averaged over gas and solid phase, which is denoted by ( ) in equation (1). For the species equation effective is calculated as mean value of molar and Knudsen diffusion coefficient. diffusivity reff For the energy equation a mean value over all species including the contribution of radiation is used to calculate thermal conductivity. While distributions of species and energy over the particle radius are changing during conversion, all property values are recalculated after each time step. The flow velocity within the porous particle is estimated by the continuity equation assuming constant outflow of the produced gases from the particle. The one dimensional conservation equations are discretized in space by a finite volume approach and a backward &fferentiation formula is applied to time integration.
587
During conversion of the solid fuel, the rate of homogeneous and heterogeneous reactions which take place in the particle is calculated by a standard Arrhenius approach
ii = k,,,i exp(--). Eai
R.T
fl~,,~ k=l
Thus, all chemical reactions are described by the same formal approach. For heterogeneous reactions the concentration of the solid phase can be replaced by the specific surface, allowing the description of intrinsic rates. Dependant on the ash content shrinking of the particle during thermal degradation is taken into account. Here it is assumed that the ash remains in the centre of the particle leaving an inert particle behind at the end of the degradation process. Within the model, a certain number of reactions is combined to a reaction scheme which describes the conversion of the solid fuel. For wood no general conversion model exists, thus the kinetic data is separated from the particle model and managed by a database, allowing an easy exchange of the underlying data. The particles are assembled to a packed bed by assuming a constant porosity in the bed and coupled to the surrounding gas phase by heat and mass transfer. Therefore, a second model describes the gas phase in the porous bed. For the description of the flow, the bed is modelled as a porous media where the void space is formed by the particles. It is assumed that primary air is flowing continuously through the bed and transport due to diffusion can be neglected compared to convection. The mass conservation in the gas phase is then only dependant on the divergence of the flow field and the source term due to mass exchange with the porous particles and reactions in the void space,
To calculate flow velocity within the bed, the momentum equation is solved. The pressure loss is calculated via the Darcy equation, implying low flow velocities and a linear dependence of the pressure loss on the flow velocity. Thus the momentum equation is given by
Similar to the species transport, diffusion of heat compared to the convective transport is negligible. Thus, the energy balance is given by , , +wth adt ( P G ~ , ~ ) + ~ ( p G=~-vpvC G ~ v+ ~G ) ~ +zCh
where Gmn, Gch and G,,, are source terms due to interaction with the solid phase and reactions in the gas phase.
588
The system of equations is discretized in space by a finite volume approach, while for the time integration an implicit Euler method is used. Particle and flow model are solved consecutively, which implies that conditions in the bed change slowly Compared to the integration step. To reduce the required computation time the flow model is solved here only for one dimension even if the software library TOSCA provides also classes for a two dimensional approach. In the packed bed calculations presented here, the particles are assembled in an one dimensional pile with a given porosity and a certain number of neighbour particles. Thus, each cell of the discretized flow field is filled up, until the given porosity is reached. It is assumed that the porosity of the bed remains constant during conversion. Therefore, in case of depletion of the solid, particles move downward within the mesh of the flow field, resulting in a decreasing bed height.
CONVERSION OF WOOD During the conversion of wet wood, drying, pyrolysis, char combustion and gasification occur in the particle. Depending on the boundary conditions, such as heating rate or particle diameter, the different processes might overlap during the conversion. However, for wood particles the conversion process not only depends on the boundary conditions which are present at the particles surfaces, but also on the history each particle has experienced before. Distribution of energy and species within the particle change with varying boundary conditions, thus no general variable exists through which the state of conversion can be characterised. In models treating the solid phase in the bed as a porous media, these influences are difficult to include, because of the assumption of homogeneous conditions over the particle. In the present contribution mass loss history of each particle is monitored and the influence of the boundary conditions is accounted for.
For modelling the drying process in wood, different approaches have been proposed in the past. In numerical models the description of the mass loss of water by a heterogeneous reaction, as introduced by [9], is advantageous with respect to numerical stability. However, the macrokinetic data used in this approach is specific to certain boundary conditions, which excludes this model from a general use. Thus, a different approach is chosen here. Under the assumption that drying occurs at a constant evaporation temperature, the drying rate can be determined by the energy equation in the particle assuming that drying is dependant only on heat transfer in the particle. If the temperature at a certain position reaches the evaporation temperature, all additional energy flowing into the particle will be consumed by the drying process as long as water remains in the particle and the temperature does not drop below the evaporation temperature. This results in a steep drying front which is moving into the particle. The velocity of the drying front is limited by the heat transfer in the porous particle. For spherical particles the velocity of the drying front varies over the particle radius due to the changing surface to volume ratio. For an infinite plate the velocity of the drying front remains nearly constant over the radius allowing to define a mean value of the front velocity.
589
16
x
14 -
-e
12 10-
Y
g 8-
>-
6-
42 450
U=
30%,50%,70%
600
750
T-
900
[KI
Fig.2 Velocity of the drying front in an infinite plate of wood for different temperatures and water contents. Applying the single particle model to an infinite plate of wood undergoing a drying process, as shown in Fig. 2, it can be shown that the mean velocity of the drying front depends on the ambient temperature and the initial moisture content. To focus on the drying process neither pyrolysis nor char combustion is included for this case. With increasing water content of the wood particle the velocity of the drying front decreases because more water mass needs to be evaporated at a given position. The velocity of the drying front increases with temperature, however at high temperatures the temperature difference over the particle dimension is higher and with it the heat loss from the particle due to radiation, which results in a reduced slope of the curves. For pyrolysis of wood, models of different complexity have been proposed. Most of the models are based on schemes of competitive and consecutive reactions, through which the wood is converted into light gases, tars and char. Kinetic data for the different reactions vary over a wide range and no general model for the description of wood pyrolysis exists. At a certain temperature, denoted as isokinetic temperature, the different kinetic data lead to comparable pyrolysis rates, as shown by [lo]. However, for particles sized between 5 and 25 mm the pyrolysis is determined by heat transfer over the particles. Even if in the outer part of the particle the temperature is close to the isokinetic temperature, the pyrolysis rate for the different kinetic data varies due to the non homogeneous temperature distribution in the particle. As shown in Fig. 3 for different pyrolysis kinetics and a spherical particle which is exposed to a radiation flux fiom a blackbody close to the isokinetic temperature, the time for pyrolysis may vary by a factor of 3. The difference between the different kinetics increases with increasing diameter, because of the higher temperature difference over the particle. Due to the sensitivity to the kinetic data, the database and the computation model are separated to investigate the influence of different models and data on the conversion process.
590
0
20 25 [mm] Pa Fig.3 Time for pyrolysis for a spherical wood particle and different pyrolysis models. 5
10 15 particle diameter d
The porous char which is generated during pyrolysis is converted in the bed either by combustion with oxygen, or gasification reactions with carbondioxide or vapour. The kinetic rate for the heterogeneous reactions is affected by transport processes of the gaseous reactants to and products from the surface. Thus, within the particle model the kmetic rate of a heterogeneous reaction may be represented as intrinsic rate by replacing the species concentration in equation (2) by the specific surface. For wood, the kinetic data and the evolution of the specific surface during heterogeneous reactions is uncertain.
0
0.4 0.6 0.8 1 conversion [-I Fig.4 Calculated values for the particle radius as function of the fuel conversion for a spherical particle and O2 or C02as reactant. 0.2
Depending on the boundary conditions, the regime in which the char conversion takes place, might vary between shrinking core and reacting core behaviour, as illustrated in Fig.4 for conversion of a char particle with oxygen and carbondioxide based on the single particle model. Kinetic data for char reactions is taken from [14]. 591
Under certain boundary conditions, a simplified approach, such as the shrinking core model, does hold, as shown in Fig. 4 for the conversion with oxygen. However, for a general model this simplification is excluded to capture the cases in which differences from the ideal regime occur. Due to the calculation of the mass transfer within the particle model, no assumption has to be made on the regime under which the char is converted. Thus the whole range between shrinking and reacting core regime in each particle is described by the model. As can be seen from the given examples, a discretization of the particles is necessary for a general approach, in which the need for macrokinetic or empirical data is reduced. Therefore the separation of the packed bed into different particles as realised in the presented model seems necessary to capture the rate limiting parameters within the bed.
HEAT AND MASS TRANSFER IN A PACKED BED Due to the Langrangian formulation applied to the solid phase, the use of an effective thermal conductivity as usually applied to porous media is not necessary. In a packed bed heat is transported between solid particles by radiation and conduction. For materials with low thermal conductivity, such as wood, conduction contributes only to a minor extent to the overall heat transport. Furthermore, heat transfer due to convection between the primary air flow through the porous bed and the solid has to be taken into account. Heat transfer due to radiation and conduction between the particles is modelled by the exchange of heat between a particle and its neighbours. The definition of the neighbours depends on the assembly of the particles on the flow field mesh. As the surface temperature for each particle in the bed is known, Fourier’s law can be used to determine the heat flux due to conduction exchanged between two particles at different temperatures. The contact area between the particles is related to the particle radius and a contact angle which is set by the user. Each particle exchanges energy due to radiation with its neighbours. To take into account reflection between the particles it is assumed that each particle contributes its radiative loss to a global radiation field. The radiation field is related to the flow model, and the cell each particle is contributing its radiative heat loss to, is determined by the position of the neighbour particles. It is assumed that no radiation due to short ray travelling distances is absorbed in the gas phase between the particles. In order to assess transport mechanisms due to convection various correlation for heat and mass transfer coefficients in a packed bed have been derived. For the present application the transfer coefficient in the bed is related to the transfer coefficient of a single particle in a gas flow according to [15]. Due to the outflow of the gases during pyrolysis and char conversion the calculated transfer coefficient is decreased, thus Stefan correction is included to calculate the transfer coefficient at a finite flow over the boundary. In order to investigate convective heat and mass transfer, a laboratory reactor has been set up at the Research Centre Karlsruhe [ 161, see F i g 5 A hot gas flow is used to heat up the packed bed, which is located in a furnace pot suspended from a weighting cell. To minimise heat loss from the packed bed, the reactor wall is heated. Therefore radial temperature gradients are assumed to be negligible. Thermocouples are used to determine the temperature distribution over the height. The maximum height of the packed bed in the reactor is 210 mm.
592
Arrangement of Thermocouples in the bed
Weighting Ce
TC 1 TC 2 Hot Gas Flow
TC 3 TC 4 TC 5 TC 6 TC 7 TC 8
2 cm
[&
1 cm, -
TC TC 10
Fig.5 Laboratory reactor for packed bed experiments [ 161. Comparisons of calculated and measured temperature distributions in the bed are shown in Fig.6.
300
E 250
e
g 200
c
a
150
10
E J 100
Slate, dPa= 12,6 mm
50
0 0
1000
3000 4000 5000 6000 time [s] Fig. 6 Comparison of measured and calculated temperature distributions in an inert packed bed.
2000
To focus solely on the heat transfer, slate has been used as an inert material with a low thermal conductivity, comparable to wood. Temperatures displayed refer to the gas
593
temperature in the void space of the bed. Due to the use of thermocouples for the temperature measurement, the influence of the solid on the measured value cannot be excluded in the experiments. Agreement between measured and calculated temperature distributions is satisfactory. As can be seen from Fig.6 the inlet temperature of the gas in the experiment at the beginning is below 300°C which has been chosen as boundary condition in the calculations. Moreover the final temperature measured at different locations in the bed reaches or even exceeds the gas temperature. This indicates that a certain amount of heat is transported from the heating into the bed, where adiabatic conditions have been chosen for the calculation. However, the model predicts satisfactory heat transfer due to convection within in the bed. For drying of wet material, experimental and calculated mass loss curves of the water for different gas temperatures are shown in Fig.7. To avoid interaction with the pyrolysis process, a gas temperature of 150°C has been chosen for the drying of wood. At this temperature the drying velocity is low, thus wood in the experiments contained only a small amount of water. Because pyrolysis of wood starts at temperatures of about 200"C,wet slate particles have been used for drying experiments at higher temperatures.
0
5000
10000
time [s]
time [s]
Fig.7 Drying of slate and wood particles in a packed bed for different water contents and gas temperatures.
From the experiments two sections with different drying velocities can be distinguished. The high drying velocity at the beginning is attributed to the drying of the particle surfaces. After surfaces of the different particles in the bed are dry, the drying velocity is determined by the heat transfer into the particle. Results from the packed bed model indicate a nearly constant drying velocity for the whole period. Thus, the heat transfer through the packed bed is slower than the heat transfer into the particle. This results in a drying front which moves through the bed, where the particles upstream of this front remain at their virgin water content. This can be explained by two reasons: on the one hand the heat transfer coefficient is slightly overestimated as could be seen from comparisons with experimental results [ 161, on the other hand it is reasonable that drying at the particle surfaces start at temperatures below lOO"C, which has been chosen here as constant evaporation temperature. However, taking into account that no empirical data has been used within the model and property data has 594
been taken from literature, the agreement between experimental and calculated is satisfactory. It has to be kept in mind that the main emphasis here is not the exact description of the drying process but a general description of the combustion process of wet fuel particles. Thus, the drying model has to predict the overall drying time sufficiently accurate, what can be achieved with the simple drying model implemented here.
REACTING BED OF WOOD PARTICLES To model a packed bed of wood particles pyrolysis and char conversion schemes can be selected from the database. Homogenous reactions within the void space are modelled by describing each volume cell in the numerical grid of the flow model as a continuous stirred reactor. Due to the lack of reliable kinetic data for the conversion of gaseous species under packed bed conditions, only the conversion of hydrogen and carbon monoxide is currently taken into account. For the combustion of hydrogen an infinite rate is assumed whereas the conversion of carbon monoxide is calculated according to [ 171. Distributions in monitor particle
1$-. 1, w 0
7 .-0, r
0.8
0 0 0
'0
n 0.6 m
z!Jj.- 0.4 1
.-0, m
$ 0.2 a,
n
o
0 0
q = l O O kW/rnzoo var=0.2 rn/s 0
dpa=0.015rn
-
0
d
t
'
hair
Fig.8 Example for a bed of reacting wood particles. For a simple one step pyrolysis and char combustion with oxygen, an example for the conversion of wood particles assembled to a one dimensional pile is given in Fig.8.
595
The bed of particles is exposed to a heat flux of 100 kW/m2from above, while primary air with a flow velocity of 0.2 d s is supplied at the bottom of the pile. Under the given boundary conditions the heat transfer through the bed is slow, which results in a narrow reaction front moving through the bed. Thus, the different particles are exposed to comparable boundary conditions, where the distribution of temperature, velocity and wood mass of a monitor particle in the upper of the bed is shown in Fig.8. Due to the narrow reaction front, the bed height is decreasing linearly with conversion. A monitored particle shows, that the temperature distribution during pyrolysis is non homogeneous but remains at a constant value during char combustion. Due to the gas production flow velocities in the particle during pyrolysis are in the range of c d s , preventing the oxygen in the bulk flow from reaching the particle surface before pyrolysis is completed. For char combustion under conditions where oxygen is present in the packed bed, the uncertainty of the kinetic data does not affect the conversion rate, due to the diffusion limited process.
SUMMARY A numerical model is presented to describe the thermal conversion of solid fuels in a packed bed. For wood particles it can be shown, that a discretization of the particle dimensions is necessary to resolve the influence of heat and mass transfer on the conversion of the solid. Therefore, the packed bed is described as a finite number of particles interacting with the surrounding gas phase by heat and mass transfer. Thus, the entire process of a packed bed is viewed as the sum of single particle processes in conjunction with the interaction of the gas flow in the void space of a packed bed. Within the present model, neighbour particles exchange heat due to conduction and radiation with each other. A comparison between experiments and predicted temperature profiles for heating and drying of wood and slate particles in a packed bed showed satisfactory agreement. This indicates, that heat and mass transfer processes in a packed bed are determined sufficiently accurate and thus, encourages the simulation of thermal conversion of wood particles. This process is determined by a narrow reaction front which propagates through the bed. Although further experiments are desirable, the approach based on the division of the solid phase into discrete particles appears to capture the rate limiting parameters and offers a deeper insight into the conversion of a packed bed.
Acknowledgements This study is funded by the Swiss Federal Office of Energy. The experiments have been performed by Dr. E. Schroder at the Research Centre Karlsruhe.
596
NOMENCLATURE gas concentration [mol/m3] CV specific heat [Jkg KI Ea activation energy [kJ/mol] ko frequency factor [l/Sl K permeability [m21 n shape factor (0 for plate, 1 for cylinder, 2 for sphere) pressure P r radius i reaction rate gas constant R time t T temperature gas velocity in the particle U mass of water/ mass of dry fuel U gas velocity in the bed VG species mass fraction Yi diffusion coefficient r porosity EP thermal conductivity h dynamic viscosity P density P 0 source term average value in porous media C
0
Indices ch con eff G th tot 00
chemical reactions conduction effective Gas thermal total value in the bulk flow
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7. Maschio G., Lucchesi A., Koufopanos C. (1994) Study of kinetic and transfer phenomena in the pyrolysis of biomass particles. In: Advances in thermochemical biomass conversion, 746 - 759. 8. Peters B. (1995) A detailed model for devolatilization and combustion of waste material in packed beds. 3rd European Conference on Industrial Furnaces and Boilers. Porto, 86-104. 9. Chan W. R., Kelbon M., Krieger B. B. (1985) Modelling and experimental verification of physical and chemical processes during pyrolysis of a large biomass particle. Fuel, 64, 1505-1513. 10. Grgnli M. (1996) A theoretical and experimental study of the thermal conversion of biomass. PhD thesis, NTNU, Trondheim. 11. Balci S., Dogu T., Yucel H. (1993) Pyrolysis kinetics of lignocellulosic materials. Ind. Eng. Chem. Res. 32,2573 - 2579. 12. Kung H. C. (1972) A mathematical model of wood pyrolysis. Combustion and flame, 18, 185-195. 13. Di Blasi C., Russo G. (1994) Modelling of transport phenomena and kinetics of biomass particles. In: Advances in thermochemical biomass conversion, 906 - 921. 14. Kulasekaran S. et al. (1998) Combustion of porous char particle in an incipiently fluidized bed, Fuel, 77, 1549-1560. 15. Gnielinski V. (1982) Berechnung des W2rme- und Stoffaustauschs in durchstrijmten ruhenden Schuttungen. Ve~ahrenstechnik,16( 1). 16. Schrijder E. (1999) Bestimmung des Druckverlustes und des Warmeiiberganges von gasdurchstromten Feststoffschiittungen in der PANTHA Anlage. Report No.6373 Research Centre Karlsruhe. 17. Howard J., Williams G., Fine D. (1973) Kinetics ofcarbon monoxide oxidation in postjlarne gases. 14" Symposium (Internat.) on Combustion. The Combustion Institute, Pittsburgh, 975-986
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Reactor Network Modeling of a Biomass Dedicated Swirling Combustor and a Fluidized Bed Gasifier R. El Asri, A. A. Konnov and J. De Ruyck Vrije Universiteit Brussel, Department of Mechanical Engineering, Pleinlaan 2, 1050 Brussels- Belgium
ABSTRACT: Reactor network models of two different and representative biomass conversion applications were developed and tested. The models were integrated in the commercial process simulator ASPEN PLUS. The first modeling case is a biomassdedicated swirling bed combustor. CFD modeling and experimental temperature profiles in the burner were used to develop a reactor network. A detailed reaction scheme was used for combustion modeling of volatiles. The sensitivity analysis runs led to the localization of key parameters that affect greatly the combustion quality. Modifications of the combustor design resulting in significant drop of the CO concentration in flue gases were proposed. The second modeling case is a biomass bubbling fluidized bed gasifier providing fuel gas for an indirect fired gas turbine plant. The two-phase theory was adopted for the hydrodynamic flow modeling inside the gasifier. Physical, mechanical and chemical processes were described through available models from the literature. A series of plug flow reactors were accordingly built where kinetics of both heterogeneous and homogeneous reactions were implemented. These reactors represent both bubble and dense phase with mass transfer between the two phases. A series of sensitivity studies was run to determine the parameters that affect the quality of the emitted gas fuel. The results are compared to both experimental tests performed at the University plant and data available in the open literature. No CFD calculations were performed for the gasifier modeling. The paper is an overview of the work undertaken in the framework of the Ph.D. thesis of R. El Asri where many details can be found, and which will be submitted in the near future.
INTRODUCTION The thermochemical conversion of biomass involves a series of different processes where hydrodynamic, mechanical, physical, and chemical phenomena are to be considered. It is therefore almost impossible to deal with the whole process without making drastic assumptions in order to reduce the high number of parameters in play. Unfortunately, these simplifications greatly affect the reliability of the modeling to predict the behavior of the thermochemical conversion of biomass, and its capability to contribute to the understanding of biomass thermal processing. Reactor network modeling appears as an attractive method to deal with the complex modeling of biomass conversion. The integration of such models in existing 599
commercial software allows for obtaining an accessible and reliable modeling engineering tool particularly when the CFD modeling is used as a background to provide the non-chemical phenomena information and flow interchange, whereas detailed chemistry is next performed in a post-processing step. Efforts made up to date about the development of t h s kind of modeling are very limited [l, 2, 31, particularly for biomass applications where more elaborated work is needed. Therefore, the objective of t h s work is to contribute to this research line by elaborating reactor network modeling of two different and representative applications, the bubbling fluidized bed gasification and a swirling bed combustion of different biomass fuels. Flash pyrolysis investigations were performed in the framework of the modeling work to ensure reliable and necessary model initial inputs. Dependence between the operating temperature -which is ranked among the most controlling parameters of the process- and the fast pyrolysis yields was investigated through a series of experiments. In cooperation with the university of Zaragoza (Spain) the flash pyrolysis of seven different biomasses was carried out at temperatures ranging between 600 and 900 C [ 5 ] . Adequate mathematical fittings were drawn and then implemented as main inputs in the model.
CYCLONIC BIOMASS COMBUSTOR The modeling of a small biomass combustor was performed in the framework of the EC contract JOR3-CT97-0184. The objective of this project was to improve the design and to establish the adequate operating conditions of a novel biomass combustion system suitable for small heat applications. The burner is of cyclonic type with a swirling bed of biomass particles. The equipment is particularly suited for small granulomeh-y biomass materials such as nut shells, small wood chips, grape and olive pips. A schematic of the biomass burner is presented in Figure 1. The fuel is originally fed at the middle of the conical upper part of the combustor. The biomass particles are heated up and entrained by the swirling flow. Devolatilization and combustion mainly take place in the bottom of the combustor. Light burning particles are entrained by the vortex flow to the upper section of the combustor. Combustion of the volatile gases is extended therefore to the major part of the combustor volume. Three air supplies control the cyclonic flow suspending the solid particles. In practice, the variation of the air flows is used to find conditions with acceptable level of the major pollutants, CO and NOx, while keeping an overall desired stoichometry. The considered biomass feedstocks are almond shells and wood chips. REACTOR NETWORK FLOWSHEET
The splitting of the cyclonic combustor into reactors is shown in Figure 1. The corresponding reactor network accepted in the combustion modeling is shown in Fig. 2. The reactor network structure is largely based on the current understanding of the processes in the combustor issuing from the limited experimental observations and from the CFD calculations [4]. The key assumption is a separation of the biomass pyrolysis and subsequent char and gas oxidation. Thus, the initial part of the network represents flash pyrolysis of the biomass. The combustion of the pyrolysis products starts in the reactors Dense bed 1 and Dense bed 2. These two reactors model the cylindrical zone at the bottom part of the combustor. The block Flame 1 represents the core of this bottom part. The blocks Flame 2 - Flame 4 cover the upper
600
conical part of the combustor. The blocks Cold zone 1 and 2 model the relatively cold zone
along the walls in the upper part. The primary air feeds both the Dense bed 1 and the Flame 1. The secondary air feeds the Dense bed 2. The tertiary air goes to the first part of the flame in the freeboard (block Flame 2). The products from the block Dense bed 1 are split (block Splitter 2) to feed the Flame land 2 and the Cold zone 1. A mass exchange is assumed to take place between the Cold zone 1 and Flame 3, and between the Cold zone 2 and Flame 4. A good mixing is assumed in the dense bed and in the flame. Therefore, corresponding reactors are well-stirred ones. The blocks Cold zone 1 and 2 are assumed to be plug-flow reactors. O E T S
IUY' U
Fig. I : Schematic of the cyclonic combustor and its splitting used in the reactor network The gas residence time in the reactors depends on the gas flow rate and the reactor size. The char residence time is modelled according to the char particle size. To model this behaviour the Separator blocks connected to each reactor return a fiaction of the char issuing from them back to the reactor.
PYROLYSIS SUB-MODEL The rough feedstock is split into four portions (block Splitter 1) to feed four pyrolysis blocks where the thermal decomposition takes place. This splitting is made to simulate the heating up history of the biomass particles. Indeed, the low thermal conductivity of the biomass makes the heating up of the particles relatively slow (about 40 s for a 3-mm particle diameter). During the heating up, the particles undergo the pyrolysis process at different temperatures of the heating history. The temperature of the pyrolysis reactors (blocks Pyrolysis 1 to Pyrolysis 4) is assumed to range from 600 to 900 C with a step of 100 C. The biomass splitting ratio as well as the temperature of the block Pyrolysis 4
60 1
ire set as parameters. The flash pyrolysis blocks yield tar, char and gas. Tar is assumed o be solely composed of benzene, while char is assumed to be pure carbon. Tar and :har velds as well as dry gas composition are based on the bench scale fluidized bed lash pyrolysis experiments [5]. The corresponding fitting equations are implemented in he pyrolysis reactors. The pyrolysis products are mixed (block Mixer 1) and divided into two portions to 'eed the so-called dense bed. The reactors Dense bed 1, Dense bed 2, and Flame 1 :lame 4 include extended chemical mechanism describing the combustion of char, tar ind pyrolysis gases
SEPARATOR 10
SPLlrrER 1
PYROLYSLS4
BIOMAS
SEPARATOR 1
DENSE BED 1
Fig.2: The reactor network of the cyclonic combustor. HETEREGENEOUS REACTIONS The approach of combustion modeling used in the present work is similar to that of Chelliah et al. [6]. In particular, the gaseous combustion is described by the detailed reaction mechanism, while the char combustion is described by the semi-global heterogeneous mechanism. The following reactions were originally taken into account for solid carbon: C+O.502=CO C + H20 = CO + H2
c + co*= 2 co
(1)
(2) (3)
C+2H*=CH4 (4) The rate constants of the reactions (1)-(4) increase from the bottom to the top of the combustor. This reflects the modeled specific area rise when the particle size decreases. In the absence of the direct measurements of the oxidation rate for the char formed from almond shells, different rate constants from the literature [6,7,8] have been attempted. 602
Test calculations have been performed for the single char particle combustion and compared with the experimental data of Van Der Honing [9]. The rate constant derived by Winter et al. [7] for the char formed from sewage sludge is currently adopted in the mechanism.
GASEOUS COMBUSTION MECHANISMS The C/H/N/O gaseous combustion mechanism originates from a home-made detailed reaction mechanism for small hydrocarbons' combustion described in [ 101. It includes 959 reactions for the following species: H, H2, 0, 02,OH, H02, H20, H202, CO, COz, HCO, CH3, CH4, C2H6, CH20, C2H5, CH2, CH30, CH2OH, CH, CzH2, C2H4, C2H3, CH30H, CH3HC0, C2H, CHICO, HCCO, CH3C0, CH3O2, CH302H, CHZHCO, CN, N, NH, HCN, NO, HCNO, C, HOCN, HNCO, NCO, N20, N H 2 , HNO, NO2, C2N2, NNH, NH3, N2H2, HONO, NO3, m 0 3 , N2H3, N2H4, c m , H C m 7C6H5, C3H3, C6H6, C3H4, C6H50,N2 and Ar. Tar (represented by benzene) oxidation chemistry is also included in the mechanism. T h s detailed reaction mechanism has been thoroughly validated in a wide range of conditions typical for combustion of different hydrocarbon fuels [ 111.
NETWORK MODELING RESULTS The influence of the maximum temperature in the combustor on the composition of the flue gases is shown on Figures 3 and 4. The temperature increase leads to better and faster combustion of the char and CO (Figure 3). It is clear that the temperature in the combustor cannot be too high to avoid ash melting and increase in the NOx formation. The concentration of NO is usually within 100 - 200 ppm limits (Figure 4) in reasonable agreement with experimental observations. Assuming different pyrolysis product distributions for almond shells according to the measurements of El Asri et al. [5], Font et al. [12, 131, or Parodi et al. [14], different composition of the flue gases, in particular CO contents, have been predicted. It was found that the amount of char formed in the pyrolysis stage is of great importance. The fraction of fine particles that escape to the conical part of the reactor is of relevance for the CO production, but it doesn't affect the NO formation. The variation of the air splitting ratios between different blocks is of minor importance. Particular efforts in the reactor modeling have been paid to the identification of the major sources of CO formation. Numerous sensitivity and parametric runs reveal that carbon monoxide formed in the flash pyrolysis of biomass particles is burnt out completely under typical conditions in the combustor leading to ppm levels of CO. However, slow oxidation of the char particles continues to generate additional CO. Two sources of the CO have thus been identified in the model. The first one is a slow char burning in the cold near-wall region (Cold zones 1 and 2 in Figure 1). The CO concentration may reach a few percent in thls zone. Another source is the combustion of the light biomass particles starting from the fuel injection point. Moving injection point to the bottom significantly reduces the predicted CO levels (Figure 5).
603
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,4
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Fig.4:Variation of the O2 (circles), C 0 2 (triangles), H20 (squares) and NO (crosses) concentration in the flue gases with the maximum temperature in the combustor
604
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Fig.5: CO at the exit of the reactor as a function of the fraction of fines that escape to the conical part.
COMPARISON WITH EXPERIMENTS In the experimental in-flame study performed elsewhere [ 151 the distribution of CO was qualitatively confirmed. Combustion of sieved almond shells without “dust” fraction significantly reduces CO formation. High levels of CO (up to 2 - 5 %) were observed close to the junction between cylindrical and conical parts of it, confirming possible leak of CO through the cold near-wall region. Strong non-axisymetry in the combustor flow makes detailed quantitative comparisons difficult however. Therefore, to improve the performance of the combustor the following design modifications were recommended: 1. It is necessary to reduce the amount of the biomass and char particles in the upper part of the combustor and in the cold zone. 2. It is highly necessary to keep sufficient oxygen concentration across the combustor through better mixing in the circumferential direction.
To realize these requirements the biomass feeding point was lowered; a contraction ring was placed between the dense bed and the upper part of the combustor. These modifications significantly reduce the formation of CO under similar temperatures and feeding rate conditions without impact on NO formation (patent pending).
605
BUBBLING FLUIDIZED BED GASIFIER MODELING W B GASIFICATION PLANT The W B pilot plant consists of a feeding system (a hopper equipped with a rotary valve and a conveyor screw), the fluidized bed reactor, preheating burner, cyclone and a control system. The bubbling fluidized bed gasifier has a capacity of 400 kgh and consists of a bed (0.8 m diameted0.6 m height) with an extended freeboard section (1.2 m diameter12 m height). More technical details about the gasifier can be found in [ 161. HYDRODYNAMIC FLOW BEHAVIOR The two-phase flow theory is adopted in the model; it consists in two phases, a dense phase and a bubble phase separated by a film through which the mass transfer occurs. Gases move upward in both bubble and dense phase with plug flow, which proves to be adequate to describe the flow in a bubbling fluidized bed gasifier [17]. The expansion of the bed height is determined by bubbles behavior and bed temperature. The residence time of gases in the fluidized bed is merely estimated with the gas flow rate and the cross sectional area of the reactor bed. The char is recirculated in each compartment to adjust solids residence times. A small fraction of char, found from the elutriation model [181, is allowed to escape to the above stage. FUEL PYROLYSIS AND GASIFICATION REACTIONS
In this model, the process of fuel pyrolysis is assumed to be achieved instantaneously once the fuel is introduced into the bed and then dispersed throughout the reactor due to the fluidization effect. The subdivision of the reactor into many compartments was adopted (apart from other reasons) to allow this homogeneous distribution of the fuel from the bottom to the top of the gasifier to be taken into account. Pyrolysis yelds are assumed from [5]. Reaction rates from Van Den Aarsen [I71 are used in t h ~ smodel for char-carbon dioxide and char-steam reactions kinetics respectively; these two relations are modified to account for the empirical surface rate reaction expression. OXDATION, WATER GAS SHIFT AND TAR CRAChYNG REACTIONS Many authors suggested that all the oxygen fed to the reactor is immediately consumed in a thin zone just above the air distributor [19,20]. Oxygen is however generally observed at the exit of fluidized bed gasifiers [5,19]; oxidation is therefore considered throughout the bed. Winter kinetics [7] for wood char combustion were adopted, while a multiple-step overall kinetic mechanism for the oxidation of hydrocarbons from Hautman [2 I] was used. By accounting for the kinetics of the cited reactions as well as the water gas shift reaction kinetics, the model accounts naturally for the depletion of inlet oxygen and hence a combustion zone more marked near the d e t gas distributor. To allow water gas shift equilibrium to be naturally reached, adequate kinetics from Karim [22] are used in the model. Under most of the wood fluidized bed gasification conditions, benzene turns out to be the dominant element among tar species [23]; a simplification is therefore made in the model by considering benzene and naphthalene to be the main tar compounds. A first order reaction rate from Van Den Aarsen [ 171 was adopted in the model. Since the bubble phase is assumed to be solid free, the char generated in the bubbles through tar 606
cracking is driven to the dense phase via mass transfer blocks to respect the two-phase theory principle used in the model.
REACTOR NETWORK FLOWSHEET Figure 6 shows the biomass fluidized reactor model network integrated in the process simulator.
PYROLYSIS 2
Fig.6: Aspen+ bubblingjluidized bed network model The rough feedstock is split into four portions to feed four pyrolysis blocks where the thermal decomposition takes place. T h s splitting is again made to take in account the heating up history of the biomass particles. During the heating up, the particles undergo the pyrolysis process at different temperatures of the heating history. The pyrolysis reactors’ temperature is assumed to range from 600 C to the final steady temperature of the gasifier. The fluidized bed gasifier itself is divided into five compartments to simulate the actual behavior of pyrolysis yields, which diffise instantaneously throughout the bed after the fuel is introduced to the reactor. This subdivision of the reactor was also selected to allow mass transfer between bubble and emulsion phases to be function of the vertical location in the bed. Each compartment is itself divided into two separate chambers for the bubble and the dense phases with plug flow reactor as model for the two chambers. The mass transfer between the chambers is ensured via mass transfer blocks where Fortran routines were implemented to account for the mass transfer rate variation with bed height and gaseous components concentration gradient [18]. A plug flow reactor also simulates the freeboard zone
607
where the water gas shift reaction is assumed to reach equilibrium. For all the described reactors, temperature, pressure and volume are the main thermodynamic inputs. Suitable kinetics for plug flow reactors is introduced using parallel Fortran blocks.
HEAT DUTYANALYSIS Since temperature is set as a constant parameter during the sensitivity runs,a heat duty analysis is necessary to determine which are the realistic situations in which the operation of the gasifier is nearly adiabatic. Figure 7 shows the heat duty of the sawdust-fueled fluidized bed gasifier as a function of the air factor and the temperature in the case kinetics are taken in consideration. The air factor was varied from 0.2 to 0.5 whde the temperature ranged from 600 to 900 C; positive heat duties correspond to exothermic situations. Three nearly adiabatic situations can be distinguished: (T=700 C, ER=0.3), (T=800 C, ER=0.35) and (T=900 C, ER=0.4). The first optimum situation corresponds to the run performed in the wood-fueled fluidized bed gasifier described above.
Fig. 7: Heat duty of the sawdust-fueled fluidized bed gasifier as a function of the air factor and the temperature in the case kinetics are taken in consideration COMPARISON BETWEEN INVESTIGATIONS
MODEL
RUNS
AND
EXPERIMENTAL
Comparison was made between the results issued from the sawdust-case model and some key experimental results reported in the open literature [16, 19,20,24,25,26,27]. Equilibrium runs are also illustrated to see how realistic the equilibrium approach can be. Figures 8, 9 and 10 compare the reported experimental results and the model results at both the equilibrium and with kinetics for major combustible gas components and tar versus bed temperature. A reasonable accordance can be seen between the kinetics-
608
model results and the experimental ones, showing that the model represents an acceptable average of the real situation, particularly for carbon monoxide and hydrogen. Results from Arauzo [24] investigation deviate from the average values reported by the other experiments, which can be explained by the different feedstock that has been used (bagasse). The tar content in the gas predicted by the model shows a less sharp decrease compared to the experiments from above 750 C; this can be due to the direct combustion of tar at high temperatures, whch is not considered in the present model due to lack of kinetic data. The equilibrium model overestimates the carbon monoxide and the hydrogen content, and underestimates the carbon dioxide concentration. Figures 11 and 12 show the comparison made between the reported experimental results and the model results at both the equilibrium and with kinetics for major combustible gas components versus the equivalence ratio. Still, the model presents a good agreement with the experimental data, particularly for hydrogen and carbon monoxide. The equilibrium model overestimates the carbon monoxide and the hydrogen content, and underestimates the carbon dioxide concentration. The higher the air factor the closer the kmetic-controlled model and the equilibrium model. Tars predicted by the model get overestimated while the air factor was increased; this divergence may come from the direct combustion of tar not considered in the model.
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Fig.12:H2 at the exit of the gasifier as a function of the air factor (Temperature = 750C)
CONCLUSIONS The reactor network modeling provides a fair understanding of the biomass thermoconversion in the complex flows. In case of a swirling bed combustor, design modifications required for pollutant abatement were recommended from this reactor network analysis. These changes in design based on the modeling predictions proved to be efficient in CO reduction. In case of a fluidized bed gasifier the model is able to predict the product gas composition with reasonable accuracy, malnly in terms of hydrogen and carbon monoxide. The model diverges from experimental investigations for the outlet tar prediction at hgh temperatures. Besides tar craclung kinetics that are already implemented in the present modeling, adequate and simplified tar oxidation kinetics should be developed to be integrated in the model. ACKNOWLEDGEMENTS The assistance of G. Colson (VUB) in the ASPEN PLUS modeling is gratefully acknowledged. This work is performed in the framework of the EC project JOR3CT97-0184 and VLIET bis 970385. REFERENCES Benedetto D., Pasini S., Falcitelli M., La Marca C., and Tognotti L., “NOX Emission Prediction From 3-D Complete Modeling to Reactor Network Analysis”, Proceedings of the Mediterranean Combustion Symposium, Turkey, June 20-25, pp.432-443, 1999. 2. Carvalho M. G., Azevedo J. L. T., and Nogueira M., “Model Based Control of Industrial Combustion Equipment”, Mediterranean Combustion Symposium, Turkey, 1999. 3. Sotudeh-Gharebaagh R., Legros R., Chaouki J., and Paris J., “Simulation of Circulating Fluidized Bed Reactors Using ASPEN PLUS”, Fuel, Vol. 77, No. 4, pp. 327-337, 1998. 4. Konnov A.A., El Asri R., De Ruyck J., Courquet J., Simonin O., and Allard G., “Progress in the Small Biomass Combustor Modeling Within the EC Co-Funded Programme”, 5th International Conference on Technologies and Combustion for a Clean Environment, vol. 11, pp. 965-971, 1999. 5 . El Asri R., De Ruyck J., Arauzo J., and Gea G., “Fluidized Bed Flash Pyrolysis of Biomass”, Proceeding of the 2”d Olle Lindstrom Symposium on Renewable Energy - Bioenergy, Sweden, pp. 121-127, 1999. 6. Chelliah, H. K., Makino, A., Kato I., Araki, N. and Law C. K., “Modeling of a Graphite Oxidation in a Stagnation-Point Flow Field Using Detailed Homogeneous and Semiglobal Heterogeneous Mechanisms with Comparison to Expedments”, Combustion and Flame, v. 104, pp. 469 - 480,1996. 7. Winter, F., Prah. M. E. and Hofbauer, H. “Temperartures in a Fuel Particle Burning in a Fluidized Bed: The Effect of Drying, Devolatilization, and Char Combustion”, Combustion and Flame, v. 108, pp. 302 314, 1997. 8. Blackham, A. U., Smoot, L.. D. and Yousefi, L., “Rates of Oxidation of Millemeter-Sized Char Particles: Simple Experiments”, Fuel, v. 73, pp. 602 - 612, 1994. 1.
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612
9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24. 25. 26. 27.
Van der Honing, G. “Volatile and Char Combustion in Large Scale Fluidized Bed Coal Combustors”, Ph.D dissertation, Twente University, Netherlands, 1991. Konnov A. A., Detailed Reaction Mechanism for Small Hydrocarbons’ Combustion, Release 0.4, http ://homepages.vub.ac.be/-akonnov/, 1998. Konnov, A.A. “Development and validation of a detailed reaction mechanism for the combustion of small hydrocarbons”, 28-th Symposium (Int.) on Combustion, Edinburgh, 2000, Abstr. Symp. Pap. p. 317. Font, R., Marcilla, A., Verdu, E., Devesa, J. Ind. Eng. Chem. Prod. Res. Dev., v. 25, pp. 491-496, 1986. Font, R., Marcilla, A., Devesa, J., Verdu, E. Ind. Eng. Chem. Prod. Res. Dev., v. 27, pp. 1143-1149, 1988. Parodi, E. and Zanella, E. Energy Biomass Wastes, v. 13, pp. 879 - 895, 1990. De Ruyck J., Konnov A. A., and El Asri R., “Pollutants Reduction in Small Biomass Combustion Systems”, EC project JOR3-CT97-0184, Twelve Months Progress Report, 1999. Maniatis K., “ Fluidized Bed Gasification of Biomass “, Ph.D dissertation , Aston University, 1986. Van Den Aarsen F.G., 1985, “Fluidized Bed Wood Gasifier Performance and Modeling”, Ph.D dissertation. Twente University of Technology, Enschede, The Netherlands, 1985. El Asri R., De Ruyck J., and Verelst H., “Wood Bubbling Fluidized Bed Gasifier Modeling Through ASPEN+”, Proceeding of Flowers’97, pp.657-668, 1997. Jiang H. and Van Morey R., 1992, “A numerical Model of Fluidized Bed Biomass Gasifier”, Biomass and Bioenergy, Vol. 3, No. 6, pp.431-447, 1992. Weimer Alan W. and Clough D.H., “Modeling a Low Pressure Steam-Oxygen Fluidized Bed Coal Gasifymg Reactor”, Chemical Engineering Science Vol. 36, pp. 549-567, Pergamon Press Ltd., 1981. Hautman, D. J., Dryer, F. L., Schug, K. P., and Glassman, I. “A Multiple-Step Overall Kinetic Mechanism for the Oxidation of Hydrocarbons”, Combustion Science and Technology, Vol. 25, pp. 219 - 235, 1981. KarimG.A., J. Inst. Fuel, 53(219), 1974. Kinoshita C.M., Wang Y., Zhou J.,”Tar Formation Under Different Biomass Gasification Conditions”, Journal of Analyhcal and Applied Pyrolysis, 29, pp 169181,1994. Esperanza E., Aleman Y . , Arauzo J., and Gea G., “Fluidized Bed Gasification of Sugar Cane Bagasse. Influence on Gas Composition”, Proceedings of the 2”d Olle Lindstrom Symposium on Renewable Energy- Bioenergy- Sweden, June 1999. Narvaez I., Orio A., Aznar M. P., and Corella J., “Biomass Gasification with Air in an Atmospheric Bubbling Fluidized Bed. Effect of Six Operational Variables on the Quality of the Product Raw Gas”, Ind. Ing. Chem. Res., July- 1996. Bilodeau J.F, Proulx P. and Chomet E., “A Mathematical Model of Fluidized Bed Biomass Gasification”, Canada, 1994. Kurkela E., and Stahlberg P., “Air Gasification of Peat, Wood and Brown coal in Pressurized Fluidized Bed Reactor. Carbon Conversion, Gas Yield and Tar Formation”, Fuel Process. Techno. 3 1, pp. 1-21, 1992.
613
New Test Method to Determine Efficiency and Emissions of Slow Heat Release Appliances Burning Solid Fuel C.K. Gaegauf, Y. Macquat Center of Appropriate Technology, CH-4438 Langenbruck, Switzerland
ABSTRACT Presently there are no standards for type tests including test methods and test procedures to determine heat output, efficiency and emissions of slow hear release appliances burning wood. Slow heat release appliances are designed to supply heat after the burn cycle by applying mass heat storage. Nowadays type tests of slow heat release appliances are done according to test standards for continuous burning appl iances (e.g. DIN 18891). There is quite a discrepancy between the test procedure for actual type tests and the real world conditions of such appliances when operated. The new test methods and procedures proposed follow as much as possible the draft European and international standards for solid fuel burning appliances CEN/pr 13240 and ISOiDIS 13336 respectively. The test method combines the flue loss methods to determine the appliance efficiency and gaseous emissions, such as carbon monoxide. A calorimeter room is applied to determine the heat released over a test cycle. There is a dilution tunnel used for the measurement of particulate matter. The new test method has been examined with three slow heat release appliances burning wood. Criteria are proposed for easy identification of the appliance performance by the consumer. 1
1.1
OBJECTIVES
INTRODUCTION
Slow heat release appliances are designed to supply heat after the burn cycle by applying mass heat storage. These appliances are characterized by the heat accumulation capacity in the construction material. The heat is stored during the burn cycle where as heat is released with a time lag to the room over an extended time period. Typical examples of slow heat release appliances are soapstone stoves or tiled stoves (Kachelofen). Slow heat release appliances are widely used in the Alp regions, but also in Scandinavian countries and Finland. They are efficient, easy to operate and are appreciated by their users because of the continuous heat release. These appliances are suitable as the sole heating system for houses and also very appropriate for low energy houses.
614
1.2
TEST STANDARDS
The existing type test standards are all adapted for continuous burning appliances. Due to the lack of adequate type tests for slow heat release appliances, they are temporary type tested with test standards for continuous burning appliances such as DIN 18891. Consequently the slow heat release appliances are overheated, due to the consecutive bum cycles as required by these standards. The corresponding test results do not properly represent the performance of the appliances operating under praxis conditions. Type test standards (CENITC 295) for residential solid fuel burning appliances with < 50 kW are presently under evaluation in the European Community (EC). The draft standards of the Cornit6 de Normalisation (CEN) cover the type tests for determination of efficiency and emissions of solid fuel burning appliances, such as cookers, single room heaters, inserts and boilers. Outside the EC, the International Organization for Standardization (ISO) has worked out type test standards for solid fuel burning appliances. The I S 0 standards cover type tests (ISO/TC116 SC 3) for determination of heat output, emissions and efficiency for individual heating appliances. 1.3
COMPARISON BETWEEN THE DRAFT TEST STANDARDS CENflC 295 AND ISOflCll 6 SC 3
The CEN/TC 295 draft standard prEN 13240 [l] is based on measurements of efficiency and flue gas emissions at a nominal burning rate. The emission factors are based on concentration measurements of the pollutants in the flue gas. The efficiency is calculated indirectly by the flue loss method taking into account the thermal flue gas losses (sensible heat) and the chemical losses (combustible gases, here as carbon monoxide, CO). The ISO/TC 116 SC 3 draft standard ISO/DIS 13336 [2] is based on measurement of the Total Suspended Particles (TSP) in a dilution tunnel, where the flue gas is diluted with ambient air to a constant flow (Constant Flow Sampling, CFS). As an option, cirbon monoxide emissions can also be measured in the dilution tunnel. The heat output and efficiency are directly measured with a calorimeter room, an insulated cabin which is cooled with ambient air. The ISO/DIS standard requires tests of three different bum rates with settings at the minimum, medium and maximum burn rate. Within the scope of a research project [3], the prEN and ISO/DIS draft type test standards, especially the methods for determination of the efficiency and the measuring of the emissions, were compared with a continuous burning appliance by the Center of Appropriate Technology, Langenbruck. The efficiencies evaluated by both methods vary only marginal at all tested bum rates. The differences are within the accuracy range of the said test methods. The same is true for emission measuring methods. The measurements show that the test method with direct determination (calorimeter room) and the indirect determination (flue gas) of the efficiency can be considered as equivalent (Fig. 1) The same can be concluded for the measurements of emissions in the flue gas and the dilution tunnel (Fig. 2). Provided that the test procedures are identical for both standards, equivalent results for efficiencies and emissions are obtained.
615
The knowledge and the results of these research activities were an important basis for the evaluation of new type test standards for slow heat release appliances.
Fig. 1: Comparison of efficiencies measured in accordance with ISO/DIS and CEN/prEN test standards with a continuous burning appliances.
2
2.1
Fig. 2: Comparison of the emission factors determined in accordance with ISODIS and CEN/prEN test standards based on heat output with a continuous burning appliances.
REQUIREMENTS FOR A NEW TYPE TEST STANDARD PHILOSOPHY
The test procedure shall represent the real world operation of slow heat release appliances burning wood and take into account the transient parameters of a batch wise burning combustion process. It shall filfill typical test laboratory requirements such as reproducibility of the results, laboratory independence, measuring accuracy and simple test procedures. The test procedure shall support the wood appliance industry for their high quality products and should have a good relation of price and performance. The test procedure shall refer as much as possible to existing international (ISO) and European (CEN, Austrian standards) test standards. 2.2
SCOPE
The new test methods and procedures to determine heat output, efficiency and emissions of slow heat release appliances shall be the basis for new type test standards. The performance data puts manufacturers, architects, users and authorities in the position to compare and evaluate such appliances. 2.3
CONSIDERATION OF PARTICLE EMISSIONS
Not as the CEN standards yet, the I S 0 standards foresee measurement of particles in the flue gas. In view of the problematic nature of nanoparticles in breathing air, the measurement of particles gains in importance. This is reflected in the European clean air act with the PM 10 standard, setting limits on particle emissions of the size less than 10 microns. In order to reduce the particle emissions, the formation of particles shall be decreased at the source. Wood burning appliances produce higher particle loads than 616
other combustion systems. It is therefore important that the new type test standards include measurement of particle emissions. Attention shall be paid to the practical suitability of the measurements and low investment costs.
3
TEST INFRASTRUCTURE
The Center of Appropriate Technology has a test rig in its Laboratories for Sustainable Energy Systems with the possibility to test burning appliances in accordance with both draft standards CEN/prEN and ISO/DIS. The measurements can be realized simultaneously, considering both measuring concepts.
DETERMINATION OF EFFICIENCY AND HEAT OUTPUT
3.1
The slow heat release appliances are operated in an insulated and air cooled calorimeter room (Fig. 3), where the heat output can be measured directly. In parallel, the efficiency is determined indirectly with measurement of flue gas temperature and concentrations of carbon dioxide (CO2) and carbon monoxide (CO) in the flue gas, in accordance with the CEN/prEN flue loss method. Two methods are applied for determination of the heat release to the room. On the one hand the determination of heat output with the calorimeter room during the bum cycle results directly into the heat output curve. On the other hand the heat output is verified with the measurement of the surface temperature of the appliance.
Exhiur
I
Platform scale ~~~
Fig. 3:
Test rig: Calorimeter room combined with the dilution tunnel and a flue gas measuring section as part of ISO/DIS 13336 and prEN 13240 respectively test methods.
617
EMISSIONS
3.2
The gas sampling in the flue duct is designed in accordance with CEN/prEN. Flue gas temperature and concentrations of carbon dioxide (CO2) and carbon monoxide (CO) are measured. In parallel, a dilution tunnel (Fig. 3) is used in accordance with ISODIS. The dilution tunnel collects the flue gases of the appliances at the end of the flue duct, where they are mixed with ambient air. A fan in the dilution tunnel ensures a constant volumetric flow (Constant flow sampling, CFS). In addition, the particle emissions are measured in the dilution tunnel as well as in the flue duct (Fig. 4).
I
ilulion tunnel /flue duct flow meter
'filbr holder heaier
Fig. 4:
Particulate measurement in the dilution tunnel (arrangement based on ISODIS-draft test method 13336), respectively in the flue duct.
The total particle load as TSP (Total Suspended Particles) is determined with a measuring method following the instructions of American Standards of the US Environmental Protection Agency (US-EPA). The filter unit consists of a sampling probe directly followed by a filter and back-up filter (filter-0 50mm). The flue gas samples are taken isokineticly. In addition to the TSP measurements some fractions of nanoparticle emissions in the range of 30 to 400 nm were done with Scaning Mobility Particle Sizer (SMPS), and particles in the range of 400 nm to 20 pm with an optical Aerosol Spectrometer. The sampling point in the dilution tunnel was close to the gas sample point.
618
4
TESTPROGRAM
4.1 TEST
The test method has been evaluated on 3 typical slow heat release appliances with following objectives:
-
Determination of efficiencies by flue loss method vs. calorimeter room. Measurements of emissions in the flue duct vs. dilution tunnel. Determination of heat output with burn rate vs. calorimeter room. Determination of heat output curves with the calorimeter room vs. measurement of the appliance surface temperature. Evaluation of suitable criteria for the definition of the end of a burn cycles and a test cycle respectively.
With each of the test appliances 2 tests were performed including one conditioning burn cycle and 3 test cycles. The test cycles were chosen as specified by the manufacturers. The measurements of the emissions in ths flue duct and in the dilution tunnel started with the ignition of the kindling and ended, when COZcontent in the flue gas became lower than 2 ~01%.The measurement of the calorimeter room and the surface temperature measurements were performed during the entire heat release period. 4.2 TEST APPLIANCES As typical representatives for slow heat release appliances a soapstone stove and a tiled
stove were used for the tests. The third type of slow heat release appliance was equipped with an integrated heat exchanger to supply hot water for a central heating system. This appliance was chosen to check the feasibility of the new test method in view of the necessary recording of the heat split between heat released to the room and heat released to the central heating system.
Table I
Slow heat release appliances tested ~~~~~~
Type
~
~
~
Soapstone appliances
Tiled stove (Kachelofen)
1150
1500
7.5
10
Nominal Heat output [kW]
3
2.5
Heat cycle time [h]
7
14
Weight [kg] Fuel [kg]
619
Appliance with boiler insert 1100 20 to boiler to room
10 1.2 24
4.3 FUEL TYPES Debarked logs of beech wood were used as test fuel. The test fuel fulfil requirements regarding water content, piece size and number of pieces in accordance with the specification of ISOiDIS 13336. 5
RESULTS
5.I
EFFICIENCY
The direct (with calorimeter room) and indirect (with flue loss method) efficiencies were simultaneouslydetermined in the tests.
Table 2
Comparison of direct and indirect efficiencies for 3 different slow heat release appliances as avg. of 3 test cylces (based on lower heating value, LHV). Soapstone Tiled appliance stove
Appliance with boiler insert
Fuel
[kgl
7.5
9.5
20
Bum cycle time
[h]
1.1
0.98
3.8
Flue gas temperature
[“C]
253
212
156
Excess air ratio
[-1
3.0
3.9
2.0
Efficiency indirect (CEN/prEN)
[%I
72.3
71.7
87.9
Efficiency direct (ISO/DIS)
poi
70.2
56.4
89.0
5.2
HEAT OUTPUT CURVES
The heat output curves are a basis for definition of the heating interval and the average heat output of the appliance. Fig. 5 and Fig. 6 show the curves of the heat released measured with the calorimeter room and the appliance surface temperature method. For this method, 7 to 9 temperature propes were placed on the test stove in order to cover all characteristic parts of the appliance. The weighted average surface temperature results from the measured temperatures from each temperature probe and the corresponding surface areas.
620
Fig. 5: Heat output and average surface temperatures (weighted) of the soapstone stove over 3 consecutive bum cycles.
5.3
EMISSIONS
Gaseous emissions
The measurements of CO emissions were just done in the flue duct, because of the fact that simultaneous measurements in the dilution tunnel and the flue duct showed equivalent results in a recently published study (3). CO emissions in the flue duct as average of three burn cycles
Table 3
~~
Soapstone appliance
Tiled stove
Appliance with boiler insert
3.34
2.38
2.0
~
CO-emissions
[g/m:]
@ 13v01%02
CO-emission factor (based on fuel dry)
[gflCgfwIdyl
38.40
27.46
23.1 1
CO-emission factor
[WJinput]
2.26
1.16
1.36
[g/kWhinput]
8.14
5.80
4.90
(based on input energy)
CO-emission factor (based on input energy)
62 1
Particle emissions
The measurement of particle emissions was done over a entire burn cycle of the slow heat release appliance. The total suspended particles (TSP) were measured in the flue duct and in the dilution tunnel. The particle size distribution was just measured in the dilution tunnel. In Table 4 the emission factors are expressed as particle mass respectively particle numbers related to dry fuel. Table 4
Sample point
TSP DT”
TSP FG”
TSP FG”
Method
Flat filter [@gl
Filter cartridge [@gl
1.49
1.24 1.83
Mean value Minimum Maximum
*)
’)
6
Soapstone stove: Particle emission factors per kg dry fuel.
Flat filter
PM 0.4 DT3’ SMPS
PM 0.4 DT3’ SMPS
[@gl
[@gl
[#kl
1.29
1.68
1.49
1.09 1.58
1.55 1.72
1.68 1.12
8.7
* 10’’
3.39 * 1Ol6 12.80 * loi6
TSP DT: Total suspended particles measured in the dilution tunnel (DT) TSP FG: Total suspended particles measured in the flue gas PM 0.4 DT: Particulate matter of the fractions < 400 nm and particle number (#) with SMPS measured in the dilution tunnel (DT); Assumptions: spherical shape and density of 1500 kg/m3
DISCUSSION
Efficiencies
The efficiencies of the slow heat release appliances are shown in Table 2. The results of the directly and indirectly determined efficiencies of the soapstone stove and the appliance with the boiler insert are very similar: approx. 71 % f 1.5% for the soapstone stove and approx. 88 % f0.6 % for the appliance with boiler insert. For the tiled stove the direct method resulted in an efficiency of 56.4 %. This in contrast to the indirect one, achieving some 71.7 % efficiency. The reason can be explained as follows: The supply of combustion air was closed after completion of the bum cycle (criteria: COZless than 2 vol%), when the heat release still was in progress. The damper in the flue duct thus remained open. This resulted in a substantial heat loss due to leaking air through the flue ducts in the stove. This effect shows the necessity of dampers in the flue duct for efficiency reasons. If there is no air leakage, the results achieved confirm the equivalence of the direct and indirect test methods to determine efficiencies.
622
Determination of the heat output curves with the calorimeter room versus appliance surface temperatures. Fig. 5 and Fig. 6 show the heat output curves of the appliances and the weighted surface temperature curves. A very good conformity of these curves can be observed in case of the soapstone stove. Also the curves of the tiled stove are quite conform, although the start peak could not be reproduced with the surface temperature measurements. The heat output peak is a result from convection phenomena at the noninsulated flue duct, which is not detected by the surface temperature method. The surface temperature measurement method gives some qualitative answer on the heat release characteristic. The inaccuracy will become larger, with an increasing convective part of the heat released. In our opinion this method seems thus not to be suitable for use in a type test standard, also because of the fact that the accuracy is difficult to estimate if the method is applied to various appliances. Gaseous emissions For the determination of gaseous emissions the gas sampling from the flue duct in accordance with CEN/prEN 13240 has been found reliable. In order to facilitate easy comparisons it is suggested to express the emission factors in future based on mass of dry fuel or input heat. Particle emissions The emission factors determined in the dilution tunnel with the flat filter are approx. 12 % lower than those measured in the flue gas (Table 4). The variation though of the emission factors overlaps. The differences between flat filter and filter cartridge results from the flue gas can be explained by the somewhat lower filter efficiency of the manually filled quartz wool in the cartridge. Further criteria for the justification of the TSP-measurement in the dilution tunnel are more of a practical nature. It turned out that some condensation problems arise on the filters at the flue duct because of strong formation of water vapor at the beginning of the burn cycle. These problems arise because of the characteristic burn cycle of a slow heat release appliance: In the initial phase after igniting the fuel batch, the wood logs catch fire and release the fuel water over a short time period, thus generating water vapor while the stove is still cold. Even heating of the flat filter holder could not solve this problem entirely. The SMPS-measurements give further information on particle size distribution, as an important aspect of particle emissions [6]. With some assumptions on particle geometry and density the results of the particle fractions from the SMPS measurements can be used to estimate the total particle load TSP and the particles emission factor.
623
7
CONCLUSIONS
Based on the experimental work, a first draft was worked out for a new test method to determine heat output, eficiency and emissions of slow heat release appliances. The requirements as defined in chapter 2 such as accuracy, repeatability, laboratory independence and good pncelperformance relation were considered. Following methods are the key aspects of the newly proposed type test: 1. 2.
3.
Indirect measurement of efficiency with flue loss method. Qualitative determination of the heat output curve with a simplified calorimeter room set up. The heat output curve is used for the definition of standard heating period and standard heat output. CO-emissions and other gaseous constituents in the flue gas are measured in the flue duct. Particle emissions are measured in the dilution tunnel.
The test method is designed to meet real world conditions of an appliance operating in the field. The draft test method (see Appendix 1) has been presented to the Comitk de Normalisation CEN TC 29YWGS and is accepted as basis for a continuation of the work in the standard body. On a national level the test method has been proposed as a test standard for slow heat release appliances to be certified with the Swiss Quality Label for wood heat appliances.
8
ACKNOWLEDGEMENTS
The Swiss Federal Office of Energy (BFE) funded these investigations. REFERENCES [ 13 CENIprEN 13240, Final drafi, Residential solidfuel burning appliances, Room-
[2] [3]
[4] [5] [6]
heaters, Requirements and test methods, European Committee for Standardisation ( 1999) ISOIDIS 13336, Draft International Standard (DIS), Solidjiel burning appliances - Test method for determining power output, eflciency and flue gas emissions, International Organization for Standardization(1997) C, Gaegauf, Y. Macquat, Die europdischen und internationalen Prifiormen f i r Feuerstatten, Bundesamt fiir Energie (BFE), Switzerland, Final report November 1999 P.E. Denyer, D. Wilkins, Further investigations into the use of aflue loss method for measurement of the thermal efJiciency of wood-burning roomheaters, CRE Group Ltd., Gloucestershire GB, F&AT Report no. 70 (1994) E. Karlsvik, ed. al., Round robin test of a wood stove - emissions, International Energy Agency (IEA), Report February 1995 U. Wieser and C. K. Gaegauf, Nanoparticle emissions of wood combustion processes, June 2000, Sevilla, 1st World Conference and Exhibition on Biomass for Energy and Industry,.
624
APPENDIX 1 TYPE TEST FOR SLOW HEAT RELEASE APPLIANCES 1
SCOPE
The proposed test methods is the basis for future type test standards including test methods and procedures to determine heat output, efficiency and emissions of slow heat releaseappliances. The test methods are developed to meet to real world conditions when operation an appliance in the field to take into account the transient parameters of a batch wise burning combustion process.
2
2. I 0
0
0
2.2
BASIC LAYOUT OF TEST METHODS TO DETERMINE HEAT OUTPUT, EFFICIENCY AND FLUE GAS EMISSIONS REQUIREMEIWS Slow heat release appliances are classified in defined heating-interval-classes in accordance with the duration of heat they released beyond a minimum required heat output. To be qualified as a slow heat release appliance, the appliance has to achieve a minimum duration of heat release beyond a minimum required heat output. All appliances are tested with an identical flue duct height. This is true for natural and induced draught appliances. There are requirements for efficiency and emission levels to be met. MANUFACTURER SPECIFICATIONS The manufacturer specifies the maximum allowable fuel load for a bum cycle, which is considered
as the nominal fuel load. 0
0
3
3. I
The manufacturer declares whether the appliance can be operated at a low bum rate and specifies the minimum quantity of fuel load for the low bum rate cycle. The settings of the combustion air control devices, flue gas dampers etc. during the test cycle have to be in accordance with the manufacturer's instructions.
TEST METHOD
TEST RIG
The appliance is installed in a calorimeter room (Figure I). The set-up of the calorimeter room is in accordance with the ISO/DIS 13336. The temperatures in the input and output airflow are measured. The air is set at a constant flow. There is no balance installed in the calorimeter room, since the accumulating appliances have a high mass, which would require an expensive balance.
625
Fig. I :
Slow heat release appliance in the test rig. With the temperature difference of the input and output air of the calorimeter room the qualitative heat release is determined. With flue gas analyses efficiency and CO-emissions are measured. The dilution tunnel serves to measure particulate emissions.
The flue gases are leaded in a flue duct with a fixed height. The flue gas measuring section according to CEN/prEN e.g. 13240 standards is situated in the flue duct. The effluents leaving the flue duct are diluted with ambient air and guided to the dilution tunnel where the particulates are measured. The flow in the dilution tunnel is kept at a constant flow rate and monitored.
3.2
QUALITATIVE COURSE OF THE HEAT OUTPUT CURVE
In contrast to the ISODIS-test standard, the calorimeter room is just used to optain the qualitative curve of the heat released. The qualitative heat release is based on the temperature difference of input and output air of the calorimeter room.
Fig. 2:
Course of curves of the temperature difference (dTinout) of input and output air
in the calorimeter room. The curves show qualitatively the heat output of the accumulating appliance over three test cycles.
626
3.3
EFFICIENCY
The thermal losses are calculated based on excess air and flue gas temperature. The chemical losses a r e solely based on CO-content in the flue. The calculations are in accordance with the draft standards of CEN/prEN e.g. 13240. The measuring section is identical with said CEN/prEN standards. 3.4
FLUE GAS EMISSIONS
In the flue gas carbon monoxide and particulate shall be measured. CO-data can be used from the flue loss measurements. The data needs to be standardized to a given oxygen content (e.g. I3 vol %). If required also additional flue gas components can be analyzed such as nitrogen (NOx) and hydrocarbons (HC). The particulate measurements are done in the dilution tunnel. The particulate measurements shall not be interrupted during a bum cycle.
4 4. I
TEST PROCEDURE CONDITIONING TEST CYCLE
The conditioning bum cycle shall be carried out using nominal fuel load. The nominal fuel load is the maximum allowable fuel batch according to manufacturer's instructions. The conditioning test cycle ends, when the temperature difference of input and output air of the calorimeter room is in the range of 25 % to 30 % of maximum temperature difference of the entire conditioning test cycle. 4.2
TESTCYCLEATNOMINAL BURNRATE
After the conditioning test cycle, the appliance is tilled with the nominal fuel load. The nominal fuel load adds up kindling and test fuel. The mass of kindling is 10 % of the entire fuel load, minimum 500 Grams. The bottom up fire is started by igniting the kindling. The test fuel is added, when the fire bums lively. Test fuel shall be added not more than twice during the same bum cycle. Top down fire shall be ignited from the top by igniting the kindling on the top of the entire fuel batch. The bum cycle starts with the ignition of the kindling and ends, when C02-content in the flue gas is lower than 2 Vol%. The test cycle starts with the ignition of the kindling and ends, when the temperature difference of input and output air of the calorimeter room is in the range of 25 % to 30 % of maximum temperature difference of the test cycle. At this point the appliances can be loaded with fuel for a further test cycle. There shall be three consecutive test cycles.
627
Typlcal Test Cycle 11
10
0 8
-
- 7
e
0 6
I
Es N 0
u4 3 2 1
0 32.00
34.00
30.00
38.00
42.00
40.00
44.00
46.00
46.00
olrprad tlma [h]
Fig. 3:
4.3
Typical test cycle of a slow heat release appliance. The bum cycle is showed by the C02-curve (COZFG). The heat release is qualitatively showed by the temperature difference curve (dTinout) of the calorimeter room.
TEST CYCLE AT LOW BURN RA TE If the manufacturers claims the appliance suited for low b u m rate, a bum cycle with the fuel load in accordance with the manufacturer's instruction shall be carried out. The test procedure shall be the same as defined in chapter 4.2. In the low bum rate cycle only the emissions are measured. Efficiency and heat output are not measured.
4.4
TESTFUEL
Debarked logs of beech are used as test fuel. The test fuel has to meet requirements regarding water content, piece size, perimeter and number of pieces in accordance with the specification of ISOlDIS 13336. 4.5
REQUIREMENTS OF P E R F O W C E
The required performance data on emissions are to be achieved for nominal and minimum bum rate. The required performances of efficiency and duration of heat release for the nominal bum rate are derived from the averages of the three test cycles. 5
5.1
CALCULATION
FLUE GAS EMISSIONS
The flue gas emissions are indicated as concentration based on a given value of oxygen in the exhausts (e.g. vol 13%) and as emission factors as mass per unit fuel based on dry with ash. The calculations of the gas concentration figures are based on the CENlprEN test standards e.g. 13240. The calculations for the emission factors are new and need to be defined. 5.2
EFFICIENCY
The calculations of the efficiency are based on the CEN/prEN test standards e.g. 13240.
628
5.3
HEATRE-E
The heat released is calculated based on the fuel load and the efficiency. 5.4
NOMINAL HEATING INTERVAL
The nominal heating interval N t results from the qualitative course of the heating curve as the temperature difference between the inlet and outlet airflow of the calorimeter room. The time interval from the ignition till 33% of the maximal temperature difference achieved in the test cycle is the nominal heating interval of the appliance. The nominal heating interval is calculated by averaging the data of the three test cycles. 5.5
APPLL4NCE HEATING INTERVAL CLASS
The nominal heating interval determines the appliance heating interval class. Appliance heating interval class Interval of the nominal heating Class in h interval N t 6
8 12 16
5.6
HEAT OUTPUT
The appliance average heat output can be calculated with the appliance heating interval class and the heat released. The calculation of the heat released is based on the heat input by the test fuel and the efficiency achieved. 5.7
M4XlMUM HEAT OUTPUT
The maximal heat output can be determined based on the qualitative course of the heating curve. The pertinent time indicates how quickly the maximal heat output is reached. 5.8
HEAT OUTPUT OF F L U E G M A T E R HEAT EXCHANGER
The heat output of the boiler is determined by the heat released during the bum cycle (interval from lightning till the C02-stop criteria. The total heat released is determined by the nominal heating interval. 5.9
TESTREPORT
The test report shall include following data: CO-emission, CO-emission factor, particulate emission, particulate emission factor, efficiency, fuel load, nominal heating interval, maximum heat output, time necessary to achieve maximum heat output, heat output boiler, heat release by boiler. For minimum bum rate: CO-emission, CO-emission factor, particulate emission, particulate emission factor.
For nominal bum rate:
629
Prediction of combustion characteristics for woody biomass fuels -heat output J. Li, J. Gifford, K. Senelwa New Zealand Forest Research Institute Limited, Rotorua, New Zealand R.J. Hooper Centre for Advanced Engineering, Christchurch, New Zealand A. Clemens and D. Gong CRL Energy Limited, Lower Hutt, Wellington, New Zealand
ABSTRACT: Experimental and theoretical studies were undertaken to identify factors influencing the combustion performance of different biomass fuels. The experimental studies investigated five types of woody biomass including wood residues (Pinus radiata) and short rotation forest crop (Eucalyptus nitens). The experimental studies were based on using a laboratory combustion rig. Results showed that moisture content and particle size of biomass fuel were important factors affecting the combustion and heat recovery especially if incomplete combustion occurred. Higher moisture content led to a higher content of CO and C& in combustion emission, while a fine particle size of biomass contributed to heat loss due to more unburnt carbon in the fly ash and bottom ash. Empirical correlations derived between CO emission and fuel moisture, and between unburnt C in ash and fuel fine particles may be useful for defining fuel specifications and developing combustion control systems. In the theoretical studies, a stoichiometric model was employed to describe the overall combustion process and to examine the effects of the biomass composition on a fuel's combustion behaviour in term of flue gas composition, heat output and combustion temperature under ideal conditions. Incorporation with the correlations derived from the experimental work allows prediction of heat efficiency likely to be realised in a real combustion situation for stoker combustors. Good agreement between predicted and empirical results gives confidence in applying the technique to evaluate combustion performance of various biomass fuels without have to undertake combustion trials.
INTRODUCTION Changes to biomass fuel quality as a result of changes in fuel source or from seasonal changes can result in unpredictable combustion performance in commercial boilers. This problem is compounded when the fuel supply changes significantly to a plant
630
originally designed for a specific fuel type. In New Zealand, a number of new biomass combustion plants have been installed over recent time (1) with many experiencing some initial operation problems related to fuel quality variation. These problems were the driver for experimental work undertaken with the study, part of an ongoing programme of research, to improve our understanding of matching fuel types to combustion conditions. Many studies have considered advanced control systems and technologies for thermal heat recovery. Combustion efficiency can be improved by two-stage air supply or two-stage combustion for homogenous mixing of air and combustible gases to achieve complete combustion (2, 3). High concentrations of CO and HC in the flue gas have been attributed to incorrect air settings and that the parameters for correct setting are the oxygen content in the flue gas and combustion temperature (2). Similarly a computer simulation for process control purposes showed that incomplete combustion occurred due to the impact of combustion air on the reactions of gases from the pyrolysis phase (4).A COLambda (excess air ratio) control system was shown to be able to deal with changing fuel quality or changing rate of heat release in order to have minimal CO emissions with maximal combustion efficiency (5). Overall thermal recovery can be enhanced by cogeneration and advanced technologies such as fluidised bed combustion with fuel pre-drying (6, 7). Useful heat conversion efficiency as a performance characteristic has been evaluated for small boilers using wood fuels, which adopted a combustion equation based on oxidation of lignin, cellulose and hemicelluloses including the effect of air, moisture content and the charcoal fraction in the ash (8). In respect of the current work, however, it was necessary to develop simple tools to assist prediction of combustion behaviour for variable biomass fuels based on known fuel parameters. In addition, the combustion test rig used in these studies had not previously been adequately described for biomass materials. Moreover, because of the tendency of New Zealand fuel types to form slag due to the nature of the inorganic materials present in the biomass sources, one of the more important variables is combustion temperature. The work described in this paper seeks to resolve these issues. Combustion temperature is actually very hard to measure as it varies with fuel feeding interval and the measurement position in the flame. In order to predict heat output as well as combustion temperature, the effect of fuel characteristics on combustion performance needs to be identified and a stoichiometry of the overall combustion reaction described enables calculation of overall heat and mass balances for the combustion unit.
COMBUSTION TESTS A laboratory scale (nominally 50 kW) combustion system was used to conduct the combustion tests. The over stoker combustor is mounted in a rig comprising a convective tube bank, fouling probe, stack, cyclone and associated sampling, and monitoring equipment. Detailed descriptions of the experimental apparatus are provided elsewhere (9). It should be noted that no secondary air was used in the combustion process. Combustion tests were conducted for five biomass sources commonly available in New Zealand (Table 1). These fuels come from two species: Pinus radiatu and Eucalyptus nitens (E. nitens). Pinus radiaru is an important commercial wood species in New Zealand accounting for 90% of the country’s plantation forests. E. nitens is a purpose-grown plantation species which has been proven to be suitable as a short
63 1
rotation energy crop in New Zealand. Biomass residues used in the study were derived from bark, off cuts, sawdust, and waste material from forest harvesting. The preparation of the different fuel samples is described elsewhere (10). Special reference needs to be made to sample ENT (Table 1), a steam pre-treated E. nitens stem wood feed stock. This sample was derived from experimental work to evaluate the effect of fuel pre-treatment on combustion behaviour (1 1).Stemwood chips of E. nitens were treated via a new steam treatment process to produce a clean biomass fuel. The treatment involved steam hydrolysis, hemicellulose extraction and materials drying.
Table I Fuel types and combustion conditions. Fuel type __ RPB
Radiata pine bark
RPO
Radiata pine off-cuts with sawdust Radiata pine forest residues E. nitens stem wood E. nitens steam treated
RPFR ENS ENT
*
Particle size range 2% > 32mm 14%< 3mm with
dirt 5% > 32mm 20% < 3mm 5% > 32mm 2% < 3mm 4%>32mm 9%< 3mm 0%> 32mm 18%< 3 mm with
Quantity for -burn, k g 58.3
Feeding rate, g/An 234
Excess air* % 50
65.4
233
45
71.4
265
94
68.5
21 1
101
53.2
179
61
fibre Excess air is the excess portion of dry air volume used as measured by oxygen
content of flue gas divided by the stoichiometric air volume, E =
21% - 1. 21%-O,%
RESULTS AND DISCUSSION FUEL PROPERTIES The fuel characteristics are provided in Table 2. There are significant variations in ash content and net calorific value due to the fuel nature and moisture content. Fuel chemical compositions are in the typical ranges for softwood (Pinus radiata) and hardwood (E. nitens) (12).
FATORS INFLUENCING COMBUSTION COMPLETION The quantities of flue gas produced, flue gas composition, ash distribution and unburnt carbon in ash were determined for the combustion of the five biomass fuels (Table 3). Key observations from the combustion tests were: 1) Fuel moisture content is an important factor influencing the extent of combustion as higher levels of CO and CHq in the flue gas were related to the higher moisture content of the fuels regardless of fuel type; 2) Fuel particle size and excess air also affect the extent of combustion as more unburnt carbon in the bottom and fly (cyclone) ash was derived from combustion of the off-cuts with sawdust (RF’O) and E. nitens steam treated (ENT), both of which involved more fine (or fibre) particles and lower excess air.
632
Table 2 Fuel characteristics. RPB
Rpo
RPFR
ENS
ENT
Proximate analysis, % wet basis Moisture 38.90 Ash 3.54 Volatile 39.70 Fixed Carbon 17.86
36.70 0.34 52.20 10.76
52.90 3.15 36.81 7.33
5 1.40 0.69 39.13 8.78
11.80 0.50 74.99 12.68
Ultimate analysis, 96 oven dry basis C 53.86 H 5.38 0 34.68 N 0.25 S 0.03
51.01 6.11 42.12 0.21 0.02
49.41 5.98 37.81 0.1 1 0.0 1
49.59 5.82 42.96 0.2 1 0.02
49.30 5.97 43.11 0.20 0.04
Net calorific value, MJkg wet basis 12.98
12.84
8.99
8.37
17.35
Fuel type
It should be noted that the combustion tests were conducted with only primary air supply. Attempts were made to keep the feed rate and excess air as similar as possible between trial burns, however, the operation conditions were very hard to keep fixed, as the excess air was controlled manually and these biomass fuels do not have good flow characteristics due to their inconstant particle shape, uneven particle size distribution and high moisture content. Many studies related to incomplete combustion have focused on processes with two-stage air supply and have identified that excess air is a key factor influencing the formation of CO and CH (2,4,5, 13). It has also been shown that each wood furnace has a typical correlation between the CO-emissions and the excess air ratio regardless of fuel type (13). Based on a similar approach, a statistical analysis was conducted to determine whether correlations of CO and CH4 emissions could be established with a number of variables including; fuel moisture content, fuel particle size distribution, fuel feed rate and excess air ratio. With the limited number of observations, the best model for CO was derived as a function of fuel moisture content (Fig. 1) for the combustor used for these studies. There was no evidence for correlation of the other variables. The measured CH4 contents also had an increasing trend with higher moisture contents of the fuels, but could not to be fitted in any correlation equation. The role of moisture in the combustion mechanism is highly complex. The moisture of fuel is well known to increase the energy required to heat fuel particles to pyrolysis temperature, increase the thermal conductivity of fuel particles, and drive pyrolysis processes toward char formation (14). Furthermore, it may also reduce the rate of char oxidation by decreasing the temperature of char oxidation and by smothering reactive sites. In gas phase combustion, steam formed from evaporation of the fuel moisture may remove energy which would otherwise be available for degradation of volatiles into radicals for subsequent oxidation (14). The steam from fuel moisture in the pyrolysis gas may cause localised deficient mixing with oxygen, as there is no secondary air supply. The high CO and CH4 emission due to higher fuel moisture is a result of the combination of these different effects.
633
0.50 ?n
Ic $ 3 .a E
I
0.40
E
0.30
3 ;0.20 8 8 0.10 0.00
LdLI . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . .
-*40
20
0
- - - - - -
60
Moisture content, % wet basis
Fig. 1 Carbon monoxide emissions as a function of fuel moisture content for stoker combustor with primary air only for all fuel types.
In contrast to the CO and CH4 emission results, unburnt C in the ash was highest for the fuels of RF'O and ENT, the fuels with the lowest moisture content. These two fuels have more fine and fibre particles respectively with relatively lower excess air in the combustion test. Unburnt C in fly ash was due to the fine particles being carried over with the combustion gases and being carbonised rather than burnt. Adams (15) has found from his fixed bed modelling that particles of fuels and tramp solids smaller than about 5 mm were immediately entrained to the upward flow of combustion gas and never reach the grate. It is reasonable to expect that these particles contribute strongly to the unburnt C in the fly ash as well as the larger quantity of the fly ash. For bottom ash, particles larger than 5 mm may compact together without sufficient contact with oxygen to enable complete combustion. The lower excess air or lower airflow rate would exacerbate this condition and increase the unburnt C in bottom ash. If the fuel fineness is described by the proportions of material with particle size smaller than 3 mm, a correlation can be established between the unburnt C in the total ash and the fuel particle size distribution as shown in Fig. 2.
VI
40
y = 2 6 6 6 . 7 -~3~5 7 . 6 9 ~ + 11.241 .........
0 0%
5%
10%
15%
20%
25%
Proportion of fuel particles less than 3 mm
Fig. 2 Unburnt C in total ash as a function of the proportion of fuel particles less than 3 mm for all fuel types.
634
Table 3 Combustion Test Results.
RPB
RPO
PRFR
ENS
ENT
79.3 13.4 7.2 0.0879
79.5 13.8 6.7 0.0185 O.oo00
163
165
78.6 10.6 10.5 0.245 0.0029 143
79.4 12.5 8.1 0.0005
O.OOO4
78.9 10.5 10.2 0.4015 0.0439 116
Bottom Ash, % Fraction in total ash Unbumt Carbon
50.6 3.00
14.5 14.14
79.8 1.78
82.9 0.70
43.0 11.48
Fly Ash, % Fraction in total ash Unbumt Carbon
49.4 17.00
85.5 50.86
20.2 16.19
17.1 13.28
57.0 54.25
136
103
55
7
120
14
1
16
4
2
51
10
282
141
0
829
778
542
704
1132
1866
1944
2065
2065
1627
2695
2721
2607
2769
2759
98.56
99.12
96.25
98.23
99.30
77.70
77.92
67.07
65.12
83.40
Fuel
Flue gas, % dry mole N2
CO2 0 2
co CH4 Exhaust temperature, "C
Heat loss, Mkg wet fuel Heat loss on unbumt carbon in ash, QL1 Heat loss from ash due to temperature, & Heat loss from combustible gas in fluegas, Qu Heat loss from dry flue
O.oo00 156
gas, h - d f g
Heat loss from moisture in flue gas, Qa.,,, Total heat loss from fluegas, QIA
Efficiency, % Combustion efficiency, rll
Heat recovery, q 2
THERMAL EFFICIENCY There are two ways for determining the energy conversion efficiency for combustion plant (8). Using the direct method the useful energy, related to steam production, is measured. For the indirect method, useful energy is determined by measuring the difference between heat inputs and heat losses. In this study, the indirect method was used to evaluate thermal efficiency. Energy losses include all energy fluxes to the environment and the loss of reaction heat from incomplete combustion, accounted for as follows (Table 3): 1) QL,, heat loss due to unburnt carbon in the bottom ash and fly ash, Mkg fuel as fired;
635
Qu, heat loss due to ash discharged from the combustion unit at an elevated temperature of about 600°C for bottom ash and 200°C for fly ash, kJkg fuel as fired; Qu,heat loss due to incompletely burned gases in the flue gas, such as methane and carbon monoxide, kJkg fuel as fired; Qu,heat loss due to flue gas exit at a temperature of around 15OOC into the atmosphere, kJ/kg fuel as fired. Heat losses by radiation and convection from the furnace outside walls were not considered because these losses can be reduced to a minimum with good insulation. From these measurements, combustion efficiency and potential heat recovery were calculated by comparison with the net calorific value (kJkg fuel as fired) of the fuel as described bellow:
Combustion efficiency ql = 1 -
QL1
+
QL3
x 100%
Qcv Heat recovery q2=
2QJi -
1
i
x 100%
Q cv A DESCRIPTION OF THE OVERALL COMBUSTION REACTION Biomass combustion is a complicated process involving drying of solid fuel, pyrolysis and oxidation of pyrolysis gases and carbon. Many modelling studies to describe combustion performance have been based on these three stages (4, 5 , 14, 15). In order to further investigate the effects of biomass characteristics on the observed combustion performance, a simple stoichiometric model for wet biomass combustion was employed to describe the overall reaction process (Eqn. 3) (15). An empirical mole of biomass is arbitrarily normalised to one atom of carbon by the formula CIHiOi.
CO2 + (n, - n)Oz +
(g
4 N* +
[; + ),, k
Where: i is the fuel composition coefficients for elements of hydrogen when normalised to an empirical mole of biomass, i= 12X~/Xc; j is the fuel composition coefficients for elements of oxygen when normalised to an empirical mole of biomass, j =
3 4
- XdX,; 2 3
k is the number of moles of water per empirical mole of biomass, k = - XM/XC; n is the number of moles of 0 2 required for stoichiometric combustion, n = l+i/4-j/2; n, is the actual moles of O2 used, n, = n(l + E), E is the excess air ratio (i.e. extra air over the stoichiometric air requirement for completed combustion).
636
Xc, XH,Xo. XM are the C, H, 0 and moisture contents respectively of a fuel, based on oven dry and free of ash, S and N. In the above combustion reaction, the following assumptions are made: 1) Biomass is completely combusted to C02 and H2O; 2) Sulfur and nitrogen contents of the biomass are small and do not significantly affect oxidation reactions; 3) Ash content of the biomass does not significantly affect oxidation reactions. In Eqn 3 the gaseous products of combustion are limited to C02, 0 2 , N2 and H 2 0 . The CO and CH4 contents in the flue gas can be derived from the empirical correlations derived earlier and when combined with the stoichiometric equation, these can be used to calculate heat losses due to incomplete combustion. So doing allows the performance of different biomass fuel combustion to be described for the tested combustion rig in terms of flue gas composition, combustion temperature and useful heat output. In addition the method can also be extended to estimate the likely combustion temperature for a particular fuel. In an ideal combustion system, heat released from biomass combustion is completely converted to gaseous products' enthalpy (Eqns 4 to 6) which, in turn, determines the combustion temperature. Specific heat of each gas component is a function of temperature and can be derived as shown below (17). By combining the equations, combustion heat output (Qgasprducll) can thus be described as a function of combustion temperature. Qgasprmiucts = Qdrygas + Qwatergas, kJkg fuel (4) Where: Qdrygas
= Mdrygas [ ( c x ) coz
Qwatergas
(CX) 0 2 + (CX) N , I(T - 251, kJkg fuel
= M H 2 0 [C H,O (T - 25) + o.ol8h1, kJkg fuel
(5) (6)
wrygas is mole of dry flue gas, mole/kg fuel; M H 2 0 is mole of water in the flue gas, mole/kg fuel; X co2 , X 02 , X N 2 are respectively COz, 0 2 and N2 contents in dry flue gas, % (mole); CcO2is specific heat of C02, 0.004187 x (10.34
(T between 273-1200 OK) (17); CO, is specific heat of 0 2 0.004187 x (8.27
+ 0.00274T-195500/T2) (kJ/mole.k),
+ 0.000258T-187700T2) (kJ/mole.k), (T
between 300-3000°K) (17); C N , is specific heat of N2, 0.004187 x (6.50 + 0.001T) (kJ/mole.k), (T between 3003000'K (17); C H 2 0 is specific heat of HzO, 0.004187 x (8.22+0.00015T+0.00000134T2)(kJ/mole.k) (T between 300-2500 OK) (17);
h is latent heat of water vaporisation, 2.267 kJ/kg; T is combustion gas temperature, OK. In order to satisfy energy balance requirements, the combustion heat output (Qgasprducts) for an ideal case should equal the net calorific value (Qcv)of a biomass fuel. Using this relationship, the combustion temperature (T) can be estimated by approximation as shown in Eqn 7. Predicted combustion temperatures for the biomass fuels tested and derived by this method are shown in Table 4. Error = Qgasprmiucts - Qcv * O (7)
637
Table 4 Prediction of combustion temperature. Fuel Combustion temperature “C
RPB
RPO
RPFR
ENS
ENT
1138
1142
1031
977
1203
COMPARISON OF PREDICTED AND EMPERICAL RESULTS As previously described by using Eqn 3 the flue gas composition for a fuel with known characteristics and excess air ratio can be easily calculated and the heat efficiency of combustion derived. Predicted flue gas compositions and heat efficiencies for the fuels tested are compared in Figs 3 and 4 against experimental measured. The good agreement obtained indicates that this approach is able to provide a simple tool for description of combustion performance of different type of biomass fuels. It is recognised, however, the empirical correlations used are preliminary findings derived from a limited number of data and need to be further verified.
100%
,$ .-
90%
El
80%
50%
F’red RPB
Emp. Red Emp. Red Ernp. Pred RPO RPFR ENS
Emp. Pred Emp. ENT
Fuel type
Fig. 3 Comparison of flue gas composition between the predicted and empirical results for the five biomass fuels. Pred = predicted; Emp = empirical.
e”,
80
8 1
40
p
$b
.
.
. . . . .
60
20 0
Red Emp. F’red Emp. Pred Emp. Pred Ernp. Red Emp. RPB RPO RPFR ENS ENT Fuel type
Fig. 4 Comparison of combustion heat efficiency between the predicted and empirical results for the five biomass fuels. Pred = predicted; Emp = empirical.
638
CONCLUSIONS Preliminary experimental investigation of biomass combustion in a stoker combustor has shown that the useful heat output or heat efficiency of a biomass fuel combustion in lab-scale trials can be adequately represented by using the stoichiometric relationship and measured incomplete combustion correlations. The technique employed provided a simple and useful predicative tool for comparative assessment of the performance characteristics of different biomass fuels.
ACKNOWLEDGMENT The authors would like to acknowledge the financial support of the Public Good Science Fund (PGSF), Foundation for Research, Science and Technology, Wellington, New Zealand.
REFERENCES 1. Hooper R. J., Cox B. Gifford J. & Li J. (2000) Opportunities for renewable energy from woody biomass: A New Zealand perspective. Proceedings (CD Rom) of Chemeca 2000 Conference. 2. Nussbaumer T. (1994) Wood combustion. In: Advances in Thermochemical Biomass Conversion. (Ed. by A. V. Bridgwater), Vol. 1, pp. 575-89. Blackie Academic & Professional. 3. Astrupgaard N. P. (1996) Two stage combustion - the basis for decentralised power production. In: Biomass for Energy and the Environment. (Ed. by P chartier, G.L. Ferrero, U.M. Henius, S. Hultberg, J. Sachau & M. Wiinblad), Vol. 2, pp. 1208-14. Pergamon. 4. Stanzel W. (1994) Computer simulation of a wood chip furnace. In: Advances in Thermochemical Biomass Conversion. (Ed. by A. V. Bridgwater), Vol. 1, pp. 60519. Blackie Academic & Professional. 5. Good J. (1994) Combustion control for automatic wood firings. In: Advances in Thermochemical Biomass Conversion. (Ed. by A. V. Bridgwater), Vol. 1, pp. 590604. Blackie Academic & Professional. 6. Forss M. & Muoniovaara M. (1996) Demonstration of advanced cogeneration of heat and power based on various types of biomass using fluidised bed combustion and fuel drying. In: Biomass for Energy and the Environment. (Ed. by P chartier, G.L. Ferrero, U.M. Henius, S . Hultberg, J. Sachau & M. Wiinblad), Vol 1, pp. 217-21. Pergamon. 7. Martinez J. M., Escalada R. Murillo J. M., Esteban L. S. & Carrasco J. E. (1996) Development of a cogenration plant by combustion of biomass in an AFB combustor and heat conversion into electricity a Stirling Engine V-160. In: Biomass f o r Energy and the Environment. (Ed. By P chartier, G.L. Ferrero, U.M. Henius, S. Hultberg, J. Sachau & M. Wiinblad), Vol2, pp. 1239-44. Pergamon. 8. Peres C.A. & Horta Nogueira L. A. (1996) Measuring performance of fuelwood fuelled small boilers: results and methodological aspects. In: Biomass for Energy and the Environment. (Ed. by P chartier, G.L. Ferrero, U.M. Henius, S. Hultberg, J. Sachau & M. Wiinblad), vol2, pp. 1198-203. Pergamon. 9. Senelwa K., Gifford J., Li J., Hooper R. J., Clemens A. & Gong D. (2000) Combustion performance of New Zealand grown biofuels. In: Progress in Thermochemical Biomass Conversion. (Ed. by A. V. Bridgwater).
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10. Clemens A. H. & Gong D. (1998) Combustion trial of Eucalyptus nitens pretreated by a novel drying process. Prepared for Dr R J Hooper, Forest Research, Rotorua. Coal Research Institute, Ref. 98/1189, New Zealand. 11. Li J., Hooper G. & Nicholas I. (1998) Characteristics of a clean biomass fuel derived from steam pretreatment of Eucalyptus nitens. Proceedings (CD Rom) of Chemeca'98 Conference, Port Douglas, Australia, 28-30 September 1998. 12. Ragland K. W. & Aerts D. J. (1991) Properties of wood for combustion analysis. Bioresource Technology, 37, 161-8. 13. Nussbaumer T. (1997) Overview of biomass combustion. In: Developments in Thermochemical Biomass Conversion. (Ed. by A. V. Bridgwater & D. G. B. Boocock), Vol. 2, pp. 1229-43.Blackie Academic 8z Professional. 14. Tillman D. A. (1981) Review of mechanisms associated with wood combustion. Wood Science 13(4):177-84. 15. Adams T. N. (1980) A simple fuel bed model for predicting particulate emission from a woo-waste boiler. Combustion and Flame 39:225-39. 16. Payne F. A. (1984) Energy and mass flow computation in biomass combustion systems. Transaction of the American Society of Agricultural Engineers. Pp 1432541. 17. Perry R. H. & Green D. (1984) Perry's Chemical Engineers Handbook. 6" Edition, R.R. Donnelley & Sons Company.
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Parametric Modeling Study of Volatile Nitrogen Conversion to NO and N 2 0 during Biomass Combustion G. Loffler, F. Winter, H. Hofbauer Institute of Chemical Engineering, Fuel Technology and Environmental Technology, Vienna University of Technology Getreidemarkt 9/159, A- 1060 Vienna, Austria
ABSTRACT The combustion of biomass is one source of NO, emissions. These pollutants are formed from fuel nitrogen, due to the relatively low operating temperatures. N20 emissions from biomass combustion are low, though the combustion temperatures would result in the expectation of considerable formation. This is normally attributed to the nitrogen functionality in biomass fuels (i.e. amino groups) making NH3 the main nitrogen-containing volatile species. NH3 is mainly oxidized to NO and N2, whereas negligible NzO is formed. However, significant amounts of HCN were found also during the devolatilization stage in the combustion of different biofuels. HCN is the main precursor for N20. Thus a different mechanism must be responsible for the low N20 formation in biomass combustion. Since the volatile content is usually high in biomass fuels, the contribution of this stage to the total emissions is dominating. A recent chemical kinetic mechanism was successfully tested for its capability to describe homogeneous nitrogen chemistry relevant to devolatilization. On this basis an extensive parametric study on the volatile nitrogen conversion to NO, N 2 0 assuming simple plug flow conditions was performed. The parameters, i.e. temperature, air-to-fuel ratio, fuel load, residence time, nitrogen content, HCN/NH3ratio, and volatiles composition, were varied in a wide range. It was shown, that most of these parameters have a strong influence on the emissions. The mechanism in the volatile nitrogen conversion is discussed showing that the composition of the volatiles is an important factor determining the emissions.
INTRODUCTION
In fluidized bed combustion NO and NzO are formed from the fuel nitrogen, which originates from volatiles and char. The relative importance of these two sources depends on the volatile content or rank of the fuel [l-61. Wartha [6] showed the 64 1
contribution of volatiles combustion to be of similar significance as char combustion for a bituminous coal and to be dominating for biomass. Volatile nitrogen is measured as HCN and NH3, where the share between this two species depends on fuel type [7]. These are converted in the gas phase to N2, NO, and N20. In fluidized bed combustion also heterogeneously catalyzed reactions on the surface of the bed material may be significant. These reactions are not further considered within this work. NH3 is mainly oxidized to NO and N2, while significant amounts of N20 are formed from the oxidation of HCN. Selectivity in oxidation of the volatile nitrogenous species depends beside others on the presence of other combustible gases, the oxygen concentration, and the temperature. N20 emissions from biomass combustion are usually low, though the combustion temperatures would be expected to allow its formation. This is normally attributed to the nitrogen functionality in biomass fuels (i.e. amino groups) making NH3 the main nitrogen-containing volatile species [S]. NH3 is mainly oxidized to NO and N2, whereas the negligible N20 is formed. However, Winter et al. [9] found also significant amounts of HCN during the devolatilization stage in the combustion of different biofuels. HCN is main precursor for N20 as shown in different modeling and experimental studies [ 10,111. Thus a different mechanism must be responsible for the low N20 formation in biomass combustion. Since the volatile content is usually high in biomass fuels, the contribution of this stage to the total emissions is dominating [9]. On basis of a recent chemical kinetic mechanism [12,13], which has been tested before to describe HCN and NH3 oxidation in the relevant temperature range [ 141, an extensive parametric modeling study on the volatile nitrogen conversion to NO, N20 assuming simple flow reactor conditions was performed. The parameters, i.e. temperature, air-tofuel ratio, fuel load, residence time, nitrogen content, HCN/NH3 ratio, and volatiles composition were varied in a wide range. The composition of the volatiles and other conditions were assumed to be representative for wood combustion in a fluidized bed, respectively. So the dependency of the NO and N20 emissions on the different parameters and the capability for primary measures reducing them can be studied.
CHEMICAL KINETIC MODELING The calculations were performed assuming simple flow reactor studies applying the program PFRCalc V 2.0 [ 151. The detailed chemical kinetic reaction scheme was taken from Glarborg et al. [12,13]. This mechanism was shown to describe the oxidation of HCN and NH3 in the presence of other combustible gases (i.e. H2,CO, C&, and C2H4) in the relevant temperature range [ 141. However, in that study the concentrations of the combustible have to be kept low, to guarantee isothermal plug flow conditions and the conditions were strongly oxidizing contrary to a typical air-to-fuel ratio in practical combustors. The rates of the reverse reactions were calculated using the thermodynamic data taken from Sandia Thermodynamic Database [ 161 with changes as recommended by Glarborg et al. [12]. For the wood volatiles an average of the compositions measured by Chan [17], who performed pyrolysis experiments with Oregon lodgepole pine pellets, was used. The tar fraction was assumed to crack’and give a secondary gas composition equal to the primary gas composition. The assumed volatiles’ composition for the fuel is summarized in Table I .
642
Table 1 Volatiles composition used in the parametric study (taken from Chan [ 171).
Component H20 CH4 C2H2 C2H4 C2H6
Wood 16.43 8.93 0.16 2.02 1.1 1 44.74 26.6 1
[v-%] [v-%] [v-%] [v-%] [v-Yo] [v-%] [v-%]
co
co2
Based on the volatiles composition shown above calculations were performed at 10 v-YO02,air-to-fuel ratio ilof 1.2, a bed temperature of 85OoC, a residence time of 3 s, and a fuel-nitrogen content of 1 w-%. In accordance to the results of Wartha et al. [ 181 HCN was considered in the volatiles of the biomass fuel. The ratio of HCN to the total volatile N was assumed 0.2. Since the composition of the volatiles and the ratio of HCN and NH3 are uncertain, this parameter was varied in a wide range. To reduce the amount of calculations standard conditions were defined. Coming from these each parameter was varied one by one as indicated in Table 2. Table 2 Parameter variation. Parameters Temp. ["C] Excess air ratio O2 conc. [v-%] Res. time [s] Fuel-N [w-%] HCN/Vol-N
Standard 800 850 900 950 1000 1.2 1.3 1.4 1.5 2 3 10 15 21 5 2.5 3 0.5 1 2 3 0.2 0.3 0.5 0.7 0.8 0 1
700 750 0.5 0.75 1 1.1 1
0.5
1
1.5
0.1
2 0.3
RESULTS AND DISCUSSION
In the following the effect of all parameters are discussed in detail for volatiles thought to be representative for wood.
Temperature Figure 1 shows the evolution of NO, N20, NO2, HCN and NH3 from the fuel-N content in the temperature range 700 "C to lOOO"C, with a residence time of 3 s and 10 v-% oxygen. It is evident that NO is the dominating nitrogen containing emission over the whole range of conditions investigated. The conversion of volatile nitrogen to the other species (i.e. N20, NH3, and HCN) is always below 1.2%. The main part cannot be found as fixed nitrogen (i.e. NH3, HCN, NO, NO2, N20), but is converted to N2. The emissions of HCN and NH3 decrease with increasing temperature, due to the increasing reaction rates and increasing concentration of free radicals (i.e. 0, H, and OH). At around 750°C no NH3 is found, while still some traces of HCN can be found up to 900°C. The increased 0 and OH concentrations at higher temperatures, explains also 643
the monotonically increasing conversion to NO. The NO formation paths (i.e. Rl-R7) are preferred compared to the destruction reactions (R8-Rl4). 2.OE-2
0.6 0.5
c
2-
1.5E-2 X I
z
0.4
s 2 1.OE-2
y
E
?
4
0.2
5.OE-3
2 \
d
0.1
0 700
O.OE+O 750
,850 900 Temperature ["C]
800
950
1000
Figure I Volatile-N conversion versus temperature. Air-to-fuel ratio 1.2, O2 concentration 10 v-%, Volatile-N content 1 w-%, residence time 3 s, HCN / Vol-N ratio 0.2. Volatiles composition according to Table 1.
NHZ +O HNO+H HNO + OH
HNo+o
NH+O NH + 0 2 NCO + 0 NCO + NO NCO + NO NH + NO NH + NO NO + NH2 NO + NH2 NNH + 0 2
HNO+H NO + H2 NO + H20 NO + OH NO+H NO + OH NO + CO N20 f CO N2 + C02 NzO + H N2 + OH N2 + H20 NNH + OH N2 + HO2
t) f)
f) f) f)
f)
t) f)
f)
t) t) f) f)
t)
R1 R2 R3 R4 R5 R6 R7 R8 R9 R 10 R 11 R 12 R 13 R 14
22% of the volatile-N is reduced from NO to N2 at 8OO"C, while this amount is only 13% at 1000°C. As known from literature [lo] N 2 0 emissions show a maximum around 800°C. At lower temperatures HCN, which is the main precursor for N20, is only partly converted. For higher temperatures the formation of N20 from the intermediates NCO and NH via R 8 and R 10 is slow compared to the formation of NO
644
(R 5 - R 7). Moreover the destruction of N 2 0 by radicals and thermal decomposition becomes more important. N20 + M N20 + H
N2+O+M N2 + OH
t) t)
R 15 R 16
The formation of NO2 shows a maximum value around 800OC. It is only about 0.4% of the total nitrogen inlet for wood volatile combustion. The formation of NO2 occurs from NO via NO + H01
NO2 + OH
t)
R 17
The presence of hydrocarbons enhance the formation of H 0 2 by HCO + 0 2 CH30 + O2
f)
t)
CO + H02 CH20 + HO2
R 18
R 19
At higher temperatures, the increased presence of H-radicals may produce a progressive decrease of the NO2 to form NO (refer to R 20). Moreover the concentration of the H02-radicals is lower at higher temperatures decreasing the importance of NOz formation by R 17. NO2 + H
NO + OH
f)
R 20
Fuel-nitrogen content To study the dependence of the NO, and N 2 0 emissions on fuel characteristics, the nitrogen content in the volatiles was varied from 0.1 to 3 w-% keeping all other parameters constant at standard conditions given in Table 2. Figure 2 shows the influence of the N-content variation on NO, and N 2 0 formation. The volatile-N was assumed to consist as HCN (20%) and NH3 (80%). The amount of NH3 and HCN inlet is totally consumed in the reaction process independently on the fuel-N amount inlet. The conversion of volatile-N to NO is shown to strongly decrease with increasing N-content in the inlet. HCN and NH3 are not only oxidized to NO but the intermediates act also to reduce NO via NCO + NO NH + N O NO + NH2 NO + NH2 NNH + 0 2
N2 + C02 N2 + OH N2 + H2O NNH + OH N2 + HO2
t)
t) t) t) t)
645
R 21 R 22 R 23 R 24 R 25
2.OE-2
0.6
-
A
0.5
C
2-
-
1.5E-2 X I
z
I
0.4
zz C
I
1.OE-2
2'0.3
az
y
r 2
K\ _ \:
b 2
0.2
.
5.OE-3 ii
P
0.1
D.OE+O
0 0
0.5
1
1.5 2 N-content [w-%I
2.5
3
Figure 2 Volatile-N conversion versus volatile N-content. Air-to-fuel ratio 1.2, 0 2 concentration 10 v-%, temperature 800°C, residence time 3 s, HCN / Vol-N ratio 0.2. Volatiles composition according to Table 1. As NH2 formed from NH3 acts stronger in reducing NO compared to NH and NCO formed fi-om HCN, the effect of increasing volatile-N content on conversion to NO is stronger for the volatiles of wood assumed to contain 80% NH3.NO2 decreases similar to NO, as it is formed from NO by reaction with HOz. Since NCO and NH reacting with NO form not only N2 but also N20 (R 8 and R lo), N20 emissions increases with increasing N-content. In wood volatiles with a N-content above 2 w-% NH2 formed from NH3 reduces NO to very low levels (i.e. conversion = 8.3% for 3 w-% volatile nitrogen). Thus conversion to N20 decreases again, since NO is a precursor in nitrous oxide formation. The same trend has already been found in experimental studies [6, 9, 181. In Figure 3 the conversion of fuel-N to NO and NH3 during devolatilization of single fuel particles (ranging from bituminous coal to different biomass fuels) at 800°C and 10 kPa oxygen concentration is shown. It can be seen that these experiments provide the same qualitative trend, though for the data obtained in the Formation Rate Unit (FRU, refer to [ 6 ] )the sum of NO and NH3 have to be viewed, since the residence time in this labscale unit is very short.
646
= 5 0
1
0
2
3 4 fuel N (wa9 [w-%]
6
Figure 3 Experimentally obtained conversion of fuel-N to NO and NH3 during devolatilization for a wide range of fuels. Experiments performed in the Formation Rate Unit (FRU) at 800°C bed temperature and 10 kPa oxygen concentration (modified from [ 181). Line indicates exponential fit to the experimental data. Wartha et al. [ 181 studied the conversion of fuel-N to NO in a range from 0 to 6 w-% in biomass combustion under fluidized bed and grate combustor conditions using data provided in literature [ 19-22]. They found a similar strong decrease in fuel nitrogen conversion to NO. In Figure 4 the conversion of fuel-N to NO for these experiments is compared to the modeling results obtained in this work. The agreement between the different combustor units, the different fuels and the modeling results assuming simple plug flow conditions is astonishingly good, though the model prediction are somewhat higher between 0.5 and 2 w-% (Waf)nitrogen in the fuel. This implies that temperature and local oxygen concentration as well as combustor design are not as significant as the N-content in the fuel. Moreover it provides evidence that devolatilization is dominating in NO formation for biomass fuels. During char combustion only a small conversion of fuel-N to NO was found (i.e. about 7%), which was almost independent from the N-content [ 181. The results presented here do not agree to the correlation for the NO, emissions in circulating fluidized bed provided from Brereton et al. [23], which was proposed, however, for fuels ranging from pitch to petroleum coke and anthracite. NO,
cc
[Volatile- content].[N - conrent]
(1)
Equation 1 implies that devolatilization is dominating in NO formation and predicts the conversion of fuel-N independent from the N-content in the fuel. This was shown not to apply for the fuels studied in this work.
647
70
P
-E f-
60 0
Grate
m
diff. furnaces (data form Skreiberg et a1.,1996)
FBC
\
\
50
0
0 Stove (data from Hofbauer,l994)
-Model
2 40
results
\
Po c
30
c
0
'f
20 0
-
V
a
4
5
0
0
1
2
3
6
fuel N (waf) [w-%]
Figure 4 The conversion of fuel-N to NO from grate furnace (200 kW*),fluidized bed combustor (50 kWfi) and stove furnaces for different biofuels compared to the conversion of Vol-N to NO calculated under plug flow conditions for wood: Air-to-fuel ratio 1.2, O2 concentration 10 v-%, temperature 800"C, residence time 3 s, HCN / VolN ratio 0.2. Volatiles composition according to Table 1.
Residence Time The residence time in the plug flow reactor was varied between 0.5 and 3 s. As shown in Figure 5 the residence time has only minor effect on the emissions within the considered range. Most volatiles are consumed before 0.5 s, while ignition takes place shortly after this time (refer to Figure 5). The small effect of the residence time justifies also the assumption of 3 s residence time for the standard conditions, though this value is relatively high for some combustion systems. Only N 2 0 decreases slightly for residence times exceeding 0.5 s caused by thermal destruction (i.e. R 26), which is slow at 800°C. N2O + M
NZ+O+M
t)
648
R 26
........... ...........
0.2
1.2E-2 z
.......
zc
1.OE-2 f
f
8.OE-3
Y
9
6.OE-3
5zz
4.OE-3
f ? 4 -.
2.OE-3
-
0
"
0.5
-
,.
n
I
1
1.5 2 Residence Time [s]
T
O.OE+O
*
fi
2.5
4
3
Figure 5 Volatile-N conversion versus residence time. Air-to-fuel ratio 1.2, 0 2 concentration 10 v-YO,temperature 800°C, volatile nitrogen content 1 w-%, HCN / VolN ratio 0.2. Volatiles composition according to Table 1. Air -to-fuel ratio
The effect of the air-to-fuel ratio was studied by varying it between 0.3 and 3 keeping all other parameters constant at standard conditions (refer to Table 2). The effect on the nitrogen containing emissions is shown in Figure 6.
0.5
0.75
1 1.25 1.5 Air-to-Fuel Ratio [-]
1.75
2
Figure 6 Volatile-N conversion versus air-to-fuel ratio. 02 concentration 10 v-%, temperature 800"C, residence time 3 s, volatile-N content 1 w-%, HCN / Vol-N ratio 0.2. Volatiles composition according to Table 1.
649
Below an air-to-fuel ratio of 1 significant amounts of NH3 and HCN can be found in the off-gas. These decrease with increasing air-to-fuel ratio, while NO increases significantly. Only insignificant traces of NO2 and N20 are formed under substoichiometricconditions. Increasing the air-to-fuel ratio above 1 the N20 emissions increases strongly, small amounts of NO2 are formed, while NO increases only slightly exceeding a ratio of 1.2. Fuel load / 0, Concentration
The oxygen inlet concentration was varied between 1 and 21 v-% at a constant air-tofuel ratio by varying also the fuel load. With the lower heating value of the volatiles (composition refer to Table 1) of 426 kJ/mol the variation of the O2 concentration between 1 and 21 v-% is equal to a variation of the fuel load from 45 to 954 kJ/m3. In Figure 7 the emissions of HCN, NH3, NO, NO2, and N 2 0 versus the oxygen concentration are given. 0.4
1.2E-2
N;O 0.3
.\
-
1.OE-2
%
z-
9 -NO
8.OE-3
z-z c
0
L
' 2 0.2
6.OE-3
P
y
z-
E
0.1
HCN NH, +
-
1
0 e 0
" . 5
I
:
1
"
4.OE-3
-
: I
4 5 c
2.OE-3 0
z
1
10 15 0,concentration [v-%]
O.OE+O
1
F .
20
25
Figure 7 Volatile-N conversion versus O2 inlet concentration. Air-to-fuel ratio 1.2, temperature 800"C, residence time 3 s, volatile nitrogen content 1 w-%, HCN / Vol-N ratio 0.2. Volatiles composition according to Table 1.
As with increasing fuel load also the concentrations of the N-containing species increases the observed trends are similar to those for increasing the N-content in the volatiles. NO decreases strongly due to the enhanced reduction by nitrogen containing intermediates NH2, NCO, and NH. N20 increases with increasing fuel load, as formation through NH + NO and NCO + NO becomes more important. Only the trend of NO2 emissions with increasing fuel load is contrary to the results obtained by increasing the N-content in the volatiles: while it is decreasing there with increasing Ncontent, it increases with increasing fuel load. NO2 is mainly formed by NO + HO2
NO2 + OH
f)
650
R 17
with H02 originating e.g. from H+Oz+M
HOz + M
f)
R 27
Thus the higher concentration of Oz enhances the H 0 2 formation and with that the conversion of NO to NOz. Above 10 v-% O2 inlet concentration this increased formation levels off, as it is partly compensated by the lower NO concentration reducing the formation via R 17.
Composition of the Volatiles As discussed in detail in literature [6, 81 the ratio of HCN and NH3 in the volatiles is still uncertain. Thus in this study the distribution of volatile-N was varied from solely HCN to solely NH3. Figure 8 shows that the HCN slip increases slightly with increasing HCN ratio in the volatiles. More pronounced is the effect on N 2 0 emissions that strongly increases with the HCN fraction. HCN is an effective precursor in NzO formation via the intermediates NCO and NH, while NH3 forms only little N20 [lo]. These results are in agreement with some experimental studies [24-271. Also the emissions of NO increase when the volatile-N contains an increasing fraction of HCN indicating that the selectivity towards Nzis higher in NH3 oxidation than for HCN.
0.3
4.OE-2 $
? 9 z
0.25
z-
3.OE-2
0.2
$ 2 0
c
I
z 0 0.15
2.OE-2
0.1
8
P c
2-
1.OE-2 X
B
0.05 0
O.OE+O 0
0.2
0.4 0.6 HCNIN, [-I
0.8
1
Figure 8 Volatile-N conversion versus HCN / vol-N ratio. Air-to-fuel ratio 1.2, O2 concentration 10 v-YO,temperature 800°C, residence time 3 s, volatile-N content 1 wYO.Volatiles composition according to Table 1.
The variation of the HCNNol-N ratio indicates that the low NzO emissions from biomass combustion cannot solely explained by NH3 being the dominating volatile nitrogen containing species. Assuming all volatile-N as HCN still only 4% is converted into NzO. This is attributed to the composition of the volatiles giving a different selectivity in HCN and NH3 oxidation. To investigate this effect the inlet concentration of each species was varied by increasing it 1%. To compare the effect of the different 65 1
species the relative sensitivity of NO and N20 against variation of the inlet concentrations was calculated as summarized in Figure 9. It can be seen that the strongest effect can be found for the concentration of O2 increasing the formation of NO and N20. Very interesting is the effect of CO. Increasing its concentration increases the conversion to NO and decreases N 2 0 . CO was shown to more strongly increase the 0-radical levels than hydrocarbons (Lofleer, 2000). Thus the formation of NO via R 5 and R 7 is favored, compared to competing N20 formation via R 8 and R 10. Moreover the N20 destruction by 0-radicals and CO is enhanced. From these calculations it may be suggested that the high 0-content in biomass fuels leading to high CO contents in the volatiles is an important fact explaining the low N 2 0 emissions in biomass combustion compared to coals. Contrary to CO, hydrocarbons reduce the conversion to NO and N20. They act in different ways: reducing the air-to-fuel ratio (compare to the effect of O2 in Figure 9), enhancing the conversion of NO to NO2 (R 17), enhancing N20 destruction by increasing the H-radical level (R 16).
4,O
CI
3-0
L
.g .- 2,o .-ln 5 v)
U
1,o
al
$-
0-0
2 -l,o -2,o
-3,O I
Figure 9 Sensitivity of volatile nitrogen conversion against variation of inlet concentrations. Air-to-fuel ratio 1.2, O2 concentration I0 v-%, temperature 800°C, residence time 3 s, volatile nitrogen content I w-%. Volatiles composition according to Table 1.
CONCLUSIONS An extensive parametric modeling study on conversion of fuel nitrogen during combustion of volatiles From biomass was performed. The variations of different parameters such as temperature, fuel-N content, air-to-fuel ratio, fuel load, and HCN/fuel-N ratio have shown contrary effects in the emission of NO, and N20 as summarized in Table 3.
652
Table 3 Schematic resume of the influence of increasing parameters. ~
Parameters Temperature Fuel-N content Residence time Air-to-fuel ratio Fuel load HCN/ Vol-N ratio
Conversion to NO
T
4t 3. t
H
Conversion to N20 /I /I I
t t
/
The emissions of NO increase with increasing temperature, air-to-fuel ratio, and HCN/NH3ratio, while they decrease with vol-N content and fuel load. Contrary, N20 emissions show a maximum with temperature and vol-N content, increase with air-tofuel ratio, fuel load, and HCN/vol-N ratio. The residence time has almost no effect on the emissions in the investigated range (0.5 to 3s), since wood volatiles are highly reactive and almost completely converted before that time. It was shown, that the N-content in the fuel is a governing parameter determining the emissions of NO,. Though rather simple assumptions were applied in this work, reasonable agreement to emissions measured for various fuels in different combustors was obtained. This confirms that devolatilization is dominating NO, emissions for high volatile biomass fuels. Volatiles composition is an important factor influencing the conversion of volatile nitrogen to NO and N20. The high CO content in biomass volatiles seems also to be partly responsible for the low levels of N20 emissions in biomass combustion. As the effect of most parameters is contrary a simultaneous reduction of NO and N 2 0 by primary measures is difficult, especially as other emission limits (e.g. for CO) have also to be captured. However, affecting devolatilization conditions to favor NH3 to HCN release and minimizing air-to-fuel ratio seems promising. ACKNOWLEDGEMENT
The authors want to acknowledge the financial support by the European Union research project “Reduction of Nitrogen Oxide Emissions from Wood Chip Grate Furnaces” (JOR3CT960059). Moreover thanks are addressed to C. Lopez and E. Herrero performing part of the calculations. REFERENCES 1. Bramer, E.A., and Valk, M. (1991): Nitrous Oxide Emissions by Fluidized Bed Combustion. 11” Int. Conf. on Fluidized Bed Combustion, ASME, New York, 70 1-707. 2. Wojtowicz, M.A., Oude Lohius, J.A., Tromp, P.J.J., and Moulijn, J.A. (1991): N 2 0 Formation in Fluidised-Bed Combustion of Coal, 1 1” Int. Conf. on Fluidized Bed Combustion, ASME, New York, 1013-1020. 3. Tullin, C.J., Sarofim, A.F., Beer, J.M. (1993): Formation of NO and N 2 0 in Coal Combustion: The Relative Importance of Volatile and Char Nitrogen, J. Inst. Energy, 66,207-2 15.
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4. Winter, F.; Wartha, C.; Ltlffler, G.; Hofbauer, H. (1996): The NO and N 2 0 Formation Mechanism during Devolatilization and Char Combustion under Fluidized-Bed Conditions, 26' Symp. (Int.) on Combustion, The Combustion Institute, Pittsburgh, PA, 3325-3334. 5. Hayhurst, A.N. and Lawrence, A.D. (1996): The Amounts ofNOxand N 2 0Formed in a Fluidized Bed Combustor during the Burning of Coal Volatiles and also of Char, Combust. Flame, 105,341. 6. Wartha, C. (1998): An Experimental Study on Fuel-Nitrogen Conversion to NO and N 2 0 and on Carbon Conversion under Fluidized Bed Conditions, Ph.D. Thesis, Vienna University of Technology, Vienna, Austria. 7. Aho, M.J., Hiimiillinen, J.P., Tummavuori, J.L. (1993): Importance of Solid Fuel Properties to Nitrogen Oxide Formation Through HCN and NH3 in Small Particle Combustion. Combust. Flame, 95,22-30. 8. Leppiilahti, J. and Koljonen, T. (1 995): Nitrogen Evolutionfiom Coal, Peat and Wood during Gasifcation: Literature Review. Fuel Processing Technology, 43, 1 45. 9. Winter, F., Wartha, C., and Hofbauer, H. (1999): NO and N20 Formation during the Combustion of Wood,Straw, Malt Waste and Peat. Bioresource Techno]., 70, 39-49. 10. Kilpinen, P., and Hupa, M. (1 991): HomogeneousN20 Chemistry at Fluidized Bed CombustionConditions: A Kinetic Modeling Study. Combust. Flame, 85,94-104. 11. Wargadalam, V.J., Ltlffler, G., Winter, F., Wartha, C. (2000): Homogeneous Formation of NO and N20 #om the Oxidation of HCN and NHj at 600-100°C. Combust. Flame, 120,427-438. 12. Glarborg, P., Alzueta, M.U., Dam-Johansen, K., Miller, J.A. (1998): Kinetic Modeling of HydrocarbodNitric Oxide Interactions in a Flow Reactor. Combust. Flame, 115, 1-27. 13. Glarborg, P., Ostberg, M., Alzueta, M.U., Dam-Johansen, K., and Miller, J.A. (1998): The Recombination of Hydrogen Atoms with Nitric Oxide at High . Temperatures. 27*h Symp. (Int.) on Combust., The Combustion Institute, Pittsburgh, PA, USA, 219-226. 14. Ltiffler, G. (2000): A Modeling Study on Fuel-nitrogen Conversion to NO and N 2 0 related to Fluidized Bed Combustion. Ph.D. thesis, Vienna University of Technology, Vienna, Austria. 15. Ltlffler, G., Winter, F., Hofbauer, H. (1999): PFRCalc V 2.0: A Program To Simulate Homogeneous Reactions Including Reaction Flow Analysis and Linear SensitivifyAnalysis in a Plug-Flow Reactor. Report No.VTWS-99-FB-02, Institute of Chemical Engineering, Fuel Technology and Environmental Technology, Vienna University of Technology, Vienna, Austria. 16. Kee, R.F., Rupley, F.M., Miller, J.A. (1993): The Chemkin ThermodynamicData Base. Sandia National Laboratories Report SAND87-8215B, Livermore, California, USA. 17. Chan, W.-C. (1989): Ph.D. thesis, Universtiy of Washington, USA. 18. Waitha, C., Reisinger, K., Winter, F., Gogolek, P.E.G., and Hofbauer, H.(1997): The Importance of NO Formation Characterisitcs for the Prediction of NO Emissionsjiom Grate and Fluidized Bed Combustors. 4~ Int. Conf. on Techn. and Combust. For a Clean Environment, M.G.Carvalho et al. (Eds.), Lisbon, Portugal, 15-21.
-
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19. Nussbaumer, T. (1988): Stickoxidebei der Holzverbrennung, Heizung W b n e , No. 12, 51-62. 20. Hofbauer, H. (1 994): Charakterisierung von biogenen BrennstofSen und Verwertung von Holzaschen. 3. Holzenergiesymposium, 2 1 Oct. 1994, Zurich, Switzerland. 21. Winter, F., Wartha, C., and Hofbauer, H. (1996): A NO/N20 - Classification System of Single Fuel Particles. 4" Int. Conference on Developments in Thermochemical Biomass Conversion, 2 1 February 1996, Banff, Canada. 22. Skreiberg, O., Hustad, J.E., Karlsvik, K. (1996): Empirical NO,-Modelling and Experimental Result fiom Wood Stove Combustion. 4" Int. Conference on Developments in Thermochemical Biomass Conversion, 2 1"February 1996, Banff, Canada. 23. Brereton, C., Grace, J.R., and Lim, C.J. (1991): Environmental Aspects, Control and Scale- Up of Circulating Fluidized Bed Combustionfor Application in Western Canada. Final Report to Energy, Mines and Resources, Canada. 24. Hulgaard, T., Glarborg, P., and Dam-Johansen, K. (1991): Homogeneous Formation and Destruction of N 2 0 at Fluidized Bed Combustion Conditions, 1lth Int. Conf. on Fluidized Bed Combustion, ASME, New York, NY, 991-998. 25. Hulgaard, T., and Dam-Johansen, K. (1993): Homogeneous Nitrous oxide Formation and Destruction under Combustion Conditions.AIChE J., 39 (8), 13421354. 26. Jensen, A., Johnsson, J.E., and Dam-Johansen, K. (1993): Formation of Nitric Oxide and Nitrous Oxidefiom Heterogeneous Oxidation of Hydrogen Cyanide at Fluidized Bed Combustion Conditions. 12' Int. Conf. on Fluidized bed Combustion, ASME, New York, 447-454. 27. Wartha, C., Winter, F., and Hofbauer, H. (1999): The Trade-ofgetween N2, NO, and N20 Under Fluidized Bed Combustor Conditions, Proc. 15" Int. Conference on Fluidized Bed Combustion, ASME, New York.
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Profile Measurements and Modelling Studies for Optimisation of Combustion Processes in Wood Firing Systems S. Unterberger, C.K. Gaegauf*,R. Berger, K.R.G. Hein Institute of Process Engineering and Power Plant Technology (IVD), University of Stuttgart, Pfaffenwaldring 23,70550 Stuttgart, Germany *Centre of Appropriate Technology, 4438 Langenbruck, Switzerland
ABSTRACT The major objective of the present work is to investigate in detail the combustion processes in small and medium scale wood-burning systems with a view to achieve a considerable reduction of gaseous emissions like CO and unburned hydrocarbons. For the characterisationof the combustion behaviour by measurements detailed information about the temperatures, gas compositions, gas velocities and mixing conditions within the reaction zones of the furnaces are necessary. Profile measurements in the combustion chambers of different firing systems are carried out using suction probes for gas analysis and temperature measurements as well as applying Laser-Doppler Anemometry (LDA) for the determination of the turbulent flow field behaviour. Additionally, facilitating a basic understanding of the formation and decomposition of different gas components within the reaction zones, numerical simulation studies of the homogeneous gas phase are carried out using the simulation program AIOLOS including submodels treating fluid flow, turbulence, combustion and heat transfer. Of particular interest in this respect, besides the calculation of gas concentration and temperature fields, the turbulent behaviour of the reacting flow and the mixing conditions between combustible gases and burnout air have a strong influence on the combustion quality. The subsequent comparison of experimental and numerical results shows a good correspondence. Then, using the validated numerical model as an engineering tool, parameters such as mixing conditions, combustion air distribution and furnace geometry, and therefore, the combustion process in the wood heater, can be optimised for a complete burn-out and for reduced emissions.
INTRODUCTION One way of reducing the global COz emission problem is to substitute wood for fossil fuels in heat and power generation to a certain extent, an option also appropriate for small and medium scale combustion systems for domestic and decentralised heating purposes. However, the combustion of natural wood logs; for example in small scale heating appliances (< 15 k W d is often involved with higher emissions of unburned components such as carbon monoxide (CO), volatile organic compounds (VOC) and
656
particulate matter (PM). According to Figure 1 in Germany, the share of wood on the primary energy consumption in the domestic heating area is estimated to about 2%, whereas present investigations show that the shares on the main gaseous emissions released by domestic wood stoves are super-proportionally high, e.g. about 40% of the total CO and about 20% of the total PM emissions from firing systems for domestic heating purposes.
60 energy consumption
50
-8
s
40
C
30
k 20 10
0 natural wood
Figure 1
brown coal
hard coal
oil
natural gas
Small scale wood heaters sold in Germany 1994 [ 11.
Experimental basic investigations focusing on reduced emissions of CO, unburned hydrocarbons and NO,, are described by [2, 31. However, the objective of these investigations is an optimisation of the combustion process and therefore a considerable reduction of emissions by measurements mainly carried out in the stack of the respective wood firing system. Recent research activities in the field of wood utilisation in small and medium scale firing systems are strongly related to various regional designs of wood heaters and stoves, i.e. [4]. Furthermore, the utilisation of catalysts for reduction of pollutant components in the flue gas of wood firing systems are investigated intensively as an appropriate alternative to primary measures [ 5 ] . An additional attempt for a further improvement of the emission behaviour is a detailed analysis of the combustion process by in-furnace measurements which is also carried out by [ 6 ] .In order to get experimental data of gas concentrations, temperatures and velocity fields within the reaction zones of different types of heating appliances the project “Development of Newly Designed Wood Burning Systems with Low Emissions and High Efficiency” [7] was carried out under the JOULE I11 program of the European Commission. Facilitating the insight into the complex inter-linked phenomena of chemical reactions and turbulent flow field behaviour as well as for the investigation of parametric effects, the simulation program AIOLOS is used as an effective tool for the investigation of the combustion processes in small and medium scale wood combustion systems.
657
In the following, the measurement techniques and the fundamentals of the numerical modelling studies as well as their application on small scale domestic wood stoves will be described. Furthermore, relevant experimental results and information obtained by the numerical modelling studies as well as a detailed comparison of measured and computed data will be shown by means of an example for a commercially available tile stove heating insert.
EXPERIMENTAL SET-UP AND MEASUREMENT TECHNIQUES The experimental test field designed for investigations on the combustion and emission behaviour of solid fuelled small scale firing systems is shown in Figure 2. In order to get measurement results under practical conditions the test field is equipped with an insulated natural draft chimney with a height of 6m.
Figure 2
Experimental test field for small scale heating appliances at the IVD.
The flue gas analysis system, shown in Figure 3, is capable to determine the concentrations of CO, COz, 02,HzO and total hydrocarbons in the flue gas. Flue gas analysis is done by using infra-red gas analysers for CO, C02 and HzO. The oxyken content is determined by a paramagnetism gas analyser and the total content of hydrocarbons is determined by using a FID. Additionally, the flue gas temperature, chimney draught, combustion air volume flow as well as the temperature of the preheated combustion or secondary air is measured continuously. The experimental results of these measurements are used to characterise the general combustion and emission behaviour as well as for obtaining information about the comparability of different burn cycles. For a detailed investigation of the formation and decomposition processes of different gas species, gas concentration and temperature profile measurements using
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suction probes are carried out at different positions within the reaction zones of the selected test stoves. The profile gas analysis system is designed to measure the gas components CO, COz. O2 and total hydrocarbons simultaneously to the flue gas measurements. The application of the suction probes is quite sensitive due to the small furnace geometry and the unsteady combustion behaviour of batch-wise fed wood heaters. Therefore, combined suctiodthermocouple (NdCrNi) probes are used facilitating the measurements of gas concentrations and temperatures at the same time. Additionally, for completing the set of profile measurement data wall temperatures at flue gas measurements
natural draft chimney
Figure 3
Flue gas analysis and temperature measurement system.
different positions in the reaction zones are measured. Besides information about the gas concentrations and temperatures inside the reaction zones of the investigated stoves data about the turbulent flow field behaviour and especially about the mixing conditions between combustible gases and burnout air are of special interest. For that, profile measurements of velocities and turbulence intensities are carried out using a Laser-Doppler Anemometer. The LDA is a nonintrusive, optical measurement technique for the determination of fluid velocity components with a high temporal and spatial resolution. The measurement principle is based on the Doppler effect describing that coherent light, scattered from particles suspended in the fluid shows a frequency shift which is proportional to the particle velocity. By detailed frequency analysis of the received signal the mean velocity as well as information about the turbulence conditions can be obtained. For further and detailed information about Laser-Doppler Anemometry see [8,9]. The used Laser-Doppler Anemometer, shown in Figure 4, consists of a 2W argonion laser used with a DANTEChnvent 2D FiberFlow system facilitating the simultaneous measurement of two perpendicular velocity components. The use of a 3D-traversing system (ProMotiod isel Co.) enables to carry out automatically velocity measurements on predefined grids, adapted to different measurement tasks, which is an important feature with regard to the instationary combustion process. Magnesia particles, generated by a rotating brush seeder with a mean diameter in the range of 1 pm, are used as tracer particles.
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Figure 4
Ser-up of LDA system
TILE STOVE HEATING INSERT A tile stove heating insert, illustrated by Figure 5 , with a thermal capacity of 10 k W g representing the state of the art of wood log combustion systems was selected as a test stove. The firing principle of the stove can be described as "backside downdraft". It consists of two spatially separated, well insulated reaction zones for gasification of the wood fuel and burnout of the combustible gases, respectively.
gasificatilon zone I)
' (reaction zone
measuremen
-primary air &start-up
air
(slot)
i
secondary air (3 nozzles) Figure 5
=Onday
air
Investigated tile stove heating insert, 10 kW,h.
Secondary air added to the reacting flow is preheated while flowing through the ash box. A throat inside the burnout combustion chamber leads to an acceleration of the flow and in combination with the following expansion the mixing of combustible gases and burnout air is realised. The injection of the secondary air is realised in two ways: by means of three nozzles before the throat and by a slot in the middle of the throat.
660
zL
loom
X
secondary air I1 (slot)
suction probe/ thermocouple
-90mm
secondary air 1 (nozzles)
Figure 6 Measuring positions of gas concentrations, temperatures and velocities.
As optical access for the LDV measurements a window covering the whole burnout zone is installed in the side wall of the stove. In order to facilitate a fast change of the measuring window during a burn cycle the opening is closed by a changing frame with a glass ceramic pane. The openings for gas concentration and temperature measurements are cut in the opposite side wall of the stove. The arrangement of the different measuring access ports facilitated a simultaneous application of all mentioned measurements. The exact positions of the measuring probes for gas concentrations and temperatures as well as the different velocity measuring positions within the burnout zone of the test stove are shown in Figure 6.
EXPERIMENTAL INVESTIGATIONS In the following, experimental results of gas concentration measurements at the positions 6, 9 and 11 will be shown and discussed. All measurements are done at nominal heat output using suction probes covering the whole depth of the burnout zone. During one test burn cycle the suction probe remains at one position in order to get information about the course of gas composition and temperature at the particular measuring position. Therefore, it is important to take into account that a comparison of measured gas concentrations at three different measuring positions results in a comparison of data from three different burn cycles. Due to small differences among single test burn cycles the discussion is focussed on a more qualitatively evaluation of the obtained results. The characteristic courses of CO concentrations during a whole test burn cycle for the above mentioned profile positions 6, 9 and 11 are shown in Figure 7. While analysing the CO concentration of the position 1 1 the single combustion phases startup, main burning and char burning can be seen very clearly. The start-up phase, up to 8 minutes, is characterised by very high concentrations of CO due to an intensive gasification of the wood fuel. After establishing a stable combustion process the CO concentrations at the inlet of the secondary reaction zone remains during the main
66 1
7 6
5 4 3
2 1
0 0:OO
0:lO
0:20 0:30 0:40 0:50
1:00
Test duration [h:mrn] Figure 7
Measured CO profile concentrations.
burning phase nearly constant at values of about 3.5 Vol.%. The change from homogenous gas phase reactions to a heterogeneous char coal combustion process is defined by a strong decrease of the CO concentration at 42 minutes. The comparison of the different concentration levels for the respective measuring position shows a considerable reduction of the CO concentration between the positions 11 and 9 which is possibly due to the effect of the secondary air injection by the three nozzles, see Figure 5 . The marginal differences of CO concentration for the positions 6 and 9 can be put down to differences during the two test burn cycles. For the positions 6 and 9 the subdivision into the different burn phases is not as obvious than for the position 11.
-
3.5
0.35
0.305
0
0.252 (D 5 m
0.20%
5
b a
0.15
uj 0
5
1.0
0.10 8
I*
0.5
0.05
0.0
0.00 0:20 0:30 0:40 0:50 1:00
0
0:OO
c a
5
0:lO
P
Test duration [h:mm] Figure 8
Measured total hydrocarbon ( C h ) profile concentration.
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The course of the total hydrocarbon concentration (THC) in Figure 8, measured as methane (CI-L,) concentration, shows a similar characteristic. Only the conversion rate of the hydrocarbons between position 11 and position 9 is about 10 times higher as the conversion rate of CO. According to the above observations about the end of the main burning phase the course of the methane concentration shows an identical decrease indicating the completed gasification of the wood logs in the primary combustion chamber.
c
0 .-c
E
+ 5
85 8
6
20 18 16 14 12 10 8 6 4 2 0
I
I
I
I
I
I
I
I
1
I
I
I
0:OO
I
0:lO 0:20 0:30 0:40 0:50
1:00
Test duration [h:mm] Figure 9
Measured O2 profile concentration.
An other parameter having a strong influence on the combustion process is the oxygen concentration of the reacting flow which is shown in the following Figure 9. The course of the measured O2 profile concentration at position 11 shows a high oxygen content of about 6 Vol.% during the main burning phase in the reacting flow at the inlet of the burnout zone. Therefore, taking into account the arrangement of the wood logs, close to the inlet of the secondary reaction zone a very inhomogeneous flow field with stable streaks of combustible gases and oxygen must be assumed. For a further optimisation of the combustion process the improvement of the mixing conditions should be considered besides the other parameters such as residence time and temperatures. Another task is to investigate the turbulent flow field behaviour within the secondary reaction zone and the residence time during the different burning phases. Figure 10 shows five velocity profiles of a single burn cycle characterising the flow field during start-up, main burning and char burning. The time information below the single profiles indicate the duration for scanning the particular velocity profile.
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The main characteristics of the flow field within the secondary reaction zone are: (1) The flow entering the secondary reaction zone is very uniform. Due to the glow bed at the bottom, the entrance velocity is quite high. (2) The throat causes a remarkable acceleration of the flow. (3) The main flow is deflected to the right side due to the injection of secondary air by the nozzles and the slot. (4) At the left hand side a flue gas recirculation zone is formed over the whole height of the burnout zone. (5) The differences on the velocity level due to the course of the burn cycle can be seen very clearly. ( 6 ) The main characteristics of the flow field remain constant during the whole burn cycle.
Figure 10
Flow field behaviour during a complete burn cycle.
On the basis of the measured velocity data the residence time of the combustible gases and the flue gas in the secondary combustion zone of the tile stove insert is estimated for obtaining information whether the absolute value is high enough ensuring a complete burnout of the gases as well as the residence time distribution in the investigated area is homogeneous. The results in Figure 11 shows that the residence time of the gases is about 150 to 400 ms during the main combustion phase. Firstly it seems that the measured residence time is to short. But when the temperature is high enough and the gases and the combustion air are intensively mixed the residence time is sufficient to ensure low gas emissions of unburned species because the combustion process in this zone is controlled by fast gas phase reactions. Besides the residence time the velocity distribution is given by a grey scale contour plot in the background.
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Figure 1I Measured residence time distribution during the main combustion phase.
NUMERICAL MODELLING STUDIES The simulation program AIOLOS is developed for the numerical calculation of three-dimensional, stationary, turbulent and reacting flows in pulverised coal-fired utility boilers. AIOLOS contains submodels treating fluid flow, turbulence, combustion and heat transfer. In these submodels equations for calculating the conservation of mass, momentum and energy are solved, presupposing high Reynolds-numbers and steady-state flow conditions. It is assumed that the flow field is weakly compressible which means that the density depends only on temperature and fluid composition but not on pressure. A body-fitted grid with approximately 200.000 nodes, Figure 12, is used to represent the secondary reaction zone of the investigated test stove. The k,e-model is based on a first order turbulence model closure according to Boussinesq. In analogy to laminar flows, the Reynolds stresses are assumed to be proportional to the gradients of the mean velocities. Transport equations for the turbulent kinetic energy and the turbulent dissipation are developed from the NavierStokes equations assuming an isotropic turbulence. The implementation of this model and the parameters used can be found in [ 101.
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Figure 12 Computational, body fitted grid.
The combustible volatiles are converted by the following reactions { C1 } - {C3}: n
{CZ}:
co + -210 2
(C3): H,
-)
co,
1 + -0, + H2O 2
The distribution of water and hydrogen as products of the hydrocarbon combustion described by reaction { C1) is calculated from the water-gas-shift equilibrium. All reactions are treated as irreversible reactions and the kinetic rates of the reactions (C1 ) - (C3} are takenfrom [ll]. The Eddy Dissipation Concept (EDC) is used for treating the interaction between turbulence and chemistry in flames [ 121. The method is based on a detailed description of the dissipation of turbulent eddies. In the EDC the total space is subdivided into a reaction space, called the 'fine structures' and the surrounding fluid. In the presented reaction scheme the reactions {Cl }, (C2}, and (C3) are treated as taking place only in these fine structures, i.e. only on the smallest turbulent length scales. The discrete ordinates method in a S4-approximation is used to solve the radiation transport equation. Since the intensity of radiation depends on absorption, emission and scattering characteristics of the medium passed through, a detailed representation of the radiative properties of a gas mixture would be very complex and currently beyond the scope of a 3D-code for the simulation of industrial combustion systems. Thus, contributing to the numerical efficiency, some simplifications are introduced, even at the loss of some accuracy. The absorption coefficient of the gas phase is assumed to have a constant value of 0.2/m. The wall emissivity was set to 0.65 for the ceramic walls and to a value of 0.15 for the glass pane inserted in one side wall for optical access. A more detailed description of the used models for combustion and turbulent flow is given in [ 131. Parallel to the determination of the profile measurement data within the burnout chamber boundary conditions at the different inlets are determined. Velocity measurements are carried out on several profiles covering the total depth of the channel between the gasification/ combustion zone and the burnout zone as well as in the outlet area of the slot and the three nozzles for obtaining the combustion air volume flows added to the reacting main flow. In order to get reliable data, results of several burn cycles are used to calculate mean values of velocities and turbulence intensities. In addition to the mentioned investigations by LDV, gas concentration measurements of C02, CO, O2 and hydrocarbons ( C h ) are done at the inlet of the burnout zone. Temperature measurements are carried out for the entering combustible gases, the added burnout air and at two wall sections in order to get complete data sets characterising the combustion conditions in the investigated test stove during the different burn phases of wood log combustion. On the basis of the before mentioned experimentally determined boundary conditions the homogeneous gas phase combustion within the secondary reaction zone of the investigated test stove is simulated. By means of an example the results of the numerical simulation studies during the main burning phase, about 20 - 30 min after
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start of the burn cycle, will be opposed to the respective experimental measurement data. In Figure 13 comparisons of calculated and measured velocity components in x- and z-axis direction are depicted for different heights (z = -90 mm and z = 100 mm, see Figure 6). 4 A v
.
Measurement 1 Measurement2
Measurement 1 A Measurement2
A
v
-;o'
O
6k
"
x in mm 4n A
.
v
'
'
I
40
I
60
x in mm
2
'i
20
0
I
Measurement1 Measurement2
-Alaos
1
.-C
3
A
.-C 3
v
v
1
0-
-40
-20 0
20
40
60
80
100
-'
&o
T
/
-40
-20
0
20
40
60
80
100
x in mm
x in mm
Figure 13
Comparison of calculated and measured velocities within the burnout chamber. u, w-velocity component, height z = -90 mm top: bottom: u, w-velocity component, height z = 100 mm
The comparison of the velocity values shows that the characteristics as well as the computed and measured absolute values are in good accordance. The flow field characteristics in areas with high velocity gradients, for example the secondary air injection by the nozzles (z = -90 mm), are qualitatively well predicted. However, deviations between calculated and measured absolute velocity values can be found, especially for the prediction of the recirculation zone in the upper left part of the burnout zone (z = 100 mm). More detailed information about the mixing of combustible gases and burnout air and therefore about the combustion process can be obtained by an analysis of the calculated mixture fraction and the gas concentration and temperature fields. Figure 14 shows the mixture fraction as well as the gas concentration and temperature field calculated in the centre plane of the burnout zone. For the mixture fraction the gases entering the burnout zone are defined as component "0", the
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Figure 14
Computed mixture fraction, CO concentration and temperature within the burnout zone of a wood stove.
secondary air injected by the nozzles and the slot are defined as "1". During the modelling studies the mixing ratio between component "0" and "1" are calculated and the distribution within the centre plane is shown in Figure 14. The penetration depth of the burnout air injected by the nozzles is approximately half of the combustion chamber width combined with locally high mixing intensity. The lower output momentum of the burnout air injected by the slot ,due to the relatively large outlet area, results in a lower penetration depth and a lower mixing intensity. The total injected air is deflected without complete mixing to the left hand side of the burnout zone. Therefore, a streak is formed on the right hand side retaining approximately the initial gas concentrations.
1
c 2
e3 .8 4
,= 6
.E
i 7
E; 10 500 600 700 800 000 100011001200130014001500
temperature in "C
CO concentrationin Vol.%
Figure 15 Comparison of measured and predicted CO-concentrationsand temperatures.
The calculated gas concentration and temperature fields support the before mentioned characteristics of the investigated combustion process. The CO concentration field shows a rapid reduction of the initially high concentration values
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due to the injection of secondary air by the slot and the nozzles and the high temperatures within this part of the burnout zone. Talung into account the probable existence of stable streaks of combustible gases and air at the inlet of the burnout zone, already mentioned at the discussion of the experimental results, it could be assumed that the main effect of the injected air by the nozzles is the improvement of the mixing conditions between gases and air. Related to the asymmetric arrangement of the air distribution and therefore an insufficient mixing quality for a complete combustion the CO-streak at the right hand side can not be reduced while passing the throat in the burnout zone. In the upper part of the burnout zone lower temperatures and a reduced mixing intensity prevent a further reduction of remaining CO. A further comparison of calculated and measured CO-concentrations and temperatures, given in Figure 15, show also a good correspondence. Differences between measured and calculated data, especially for the temperatures can be put down to the already mentioned fact of an incorrect predicted size of the recirculation zone in the upper left part of the burnout zone by the numerical model. However, the experimental determination of gas concentrations as well as of gas and wall temperatures is often in involved with uncertainties and measuring errors, especially under the prevailing instationary conditions within a domestic wood heater.
CONCLUSION The present work showed the application of different measuring techniques and numerical simulation studies on a selected commercially available tile stove heating insert. For a basic understanding of the combustion process and therefore, for a further improvement of the emission behaviour, gas analysis by means of suction probes, temperature measurements by suction pyrometry as well as velocity measurements by Laser-Doppler Anemometry are carried out within the reaction zones of the stove. Numerical simulation studies using the 3D-CFD code AIOLOS are done facilitating on the one hand the insight into complex inter-linked phenomena of chemical reaction and turbulent flow field behaviour as well as on the other hand the investigations of parametric effects. In order to improve the combustion process and therefore to lower the gaseous emission of incomplete combustion the main task will be to improve the mixing conditions of combustible gases and burnout air and therefore to avoid stable streaks. This should be achieved in combination with an increased residence time of the gases in areas of high temperatures. Therefore, the developed numerical model, based on and supported by detailed experimental investigations, can be used as an effective engineering tool for the optimisation of, e.g. the air distribution arrangement, furnace geometry or injected air volume flows. The developed optimisation procedure, consisting of detailed experimental and numerical investigations on the gas concentration, temperature and flow fields, will be also suitable for different small and medium scale wood firing systems with various combustion techniques.
ACKNOWLEDGEMENTS The project is funded by the European Commission DG XI1 within the non-nuclear JOULE-I11programme under the contract number JOR3 - CT95 - 0056.
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REFERENCES 1.
2.
3.
4.
5. 6.
7.
8.
9. 10. 11.
12. 13.
Pfeiffer, F., Struschka, M., Baumbach, G. (2000): Ermittlung der mittleren Emissionsfaktoren zur Darstellung der Emissionsentwicklung aus Feurungsanlagen im Bereich Haushalte und Kleinverbraucher. Forschungsbericht 295 46 364, Umweltbundesamt,Berlin, Germany. Struschka, M (1984).: Holzverbrennung in Feuerungsanlagen. Grundlagen, Emissionen, Entwicklung schadstoffarmer Kachelofen, VDI-Fortschritt-Berichte VDI Reihe 15, Nr. 108, VDI Verlag, Dusseldorf, Germany Nussbaumer, Th. (1989): Schadstoffbildung bei der Verbrennung von Holz, Dissertation, ETH Ziirich, Switzerland Karlsvik, E. (2000): Development of Low Emission Stoves and Fireplaces in Norway, Proc. of Nordic Seminar on Small Scale Wood Combustion, Turku, Finland Berg, M., Berge, N. (1999): Development of domestic wood fved boilers with catalytic abatement of emissions, Proc. of 2d Olle Lindstrom Symposium on Renewable Energy - Bioenergy, Stockholm, Sweden Kilpinen, P., Hupa, M. (2000): An Overview of Tulisija - The Finnish Research Programme on Wood Firing Technology 1997 - 1999, Abo Akademi University, Turku, Finland Maier, H., Unterberger, S., Struschka, M., Baumbach, G., Hein, K.R.G. (1996) Joint European Project: Development of Newly Designed Wood Burning Systems with Low Emissions and High Efficiency. In: Proceedings of the 1" European Conference on Small Burner Technology and Heating Equipment, Vol. I, pp. 145-154,Ziirich, Switzerland. Ruck, B. (1987) Laser-Doppler-Anemometrie.AT Fachverlag GmbH, Stuttgart, Germany. Durst, F., Melling, A., Whitelaw, J. H. (1987) Theorie und Praxis der LuserDoppler-Anemometrie. G. Braun, Karlsruhe, Germany. Launder, B.E., Spalding, D.B. (1974): The Numerical Computation of Turbulent Flows.,Comp. Meth. Appl. Mech. Engrg., Vol. 3, S. 269-289 Howard, J.B., Williams, G.C., Fine, D.M (1972).: Kinetics of Carbon Monoxide Oxidation in Postflame Gases. 14th Symposium (Int.), The Combustion Institute, S. 975-986 Magnussen, B.F. (1989): "The Eddy Dissipation Concept". XI Task Leaders Meeting -Energy Conversion in Combustion, IEA. H. Knaus, Richter, S., Unterberger, S., Schnell, U., Maier, H., Hein, K.R.G. (2000) On the Application of Different Turbulence Models for the Computation of Fluid Flow and Combustion Processes in Small Scale Wood Heaters. In: Experimental Thermal And Fluid Science, 21, pp. 99-108. Elsevier Science Inc., The Netherlands.
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Visualization and Analysis of SEM-EDS Data of Quartz-Bed Agglomerates M. E. Virtanen, M. S. Tiainen, M. Pudas, and R. S. Laitinen Department of Chemistry, P. 0. Box 3000, FLTN-90014 University of Oulu, Finland
ABSTRACT: Samples from solid fuel combustion are commonly analysed with SEMEDS that usually produces a vast amount of data, comprising analyses for ten elements from several hundreds of chosen domains in one sample. A current emphasis on the analysis is the visualization of the results in a meaningful way. Visual methods have proven to be usehl, and the use of the three-dimensional ternary diagrams has gained popularity among the SEM-EDS users. A logical extension to the conventional ternary diagrams is the quasiternary diagrams, where more than one element can be combined in one comer. In this work we describe and demonstrate a locally written software, which is designed especially for the production of two- and three- dimensional quasiternary diagrams, as well as for calculating statistical parameters from the SEMEDS data. The mean values of the elemental composition and their relative contents within a desired region of the quasiternary diagram can be calculated.
INTRODUCTION Scanning electron microscope combined with energy dispersive X-ray spectrometer (SEM-EDS) is a common tool in the research of solid fuels and other combustion related samples. SEM-EDS is quite versatile and it has been used to analyse coal and its mineralogy, ashed biomasses, fly ash, corrosion samples, bed material, as well as, bed material agglomerates and coatings.'-7The combination of SEM-EDS with image processing facilitates the automation of the analysis procedure, and it is possible to restrict the analysis to the areas of interest, like the coatings on the bed particles or the adhesive material of the agglomerates.' Automation of the analytical procedure enables the analysis to be carried out for over thousand domains from a single sample. Usually, the contents of ten elements (Na, Mg, Al, Si, P, S, K, Ca, Ti, and Fe) are determined with EDS in each domain. The extraction of chemically meaningful information from a vast amount of data can be very time-consuming. Graphical representations can give a fast overview, and the use of ternary &agrams has gained popularity among SEM-EDS users.8 The quasiternary diagrams are a logical extension to ternary diagrams and provide a useful
67 I
visualization of analytical results, when the compositional distribution of more than three elements needs to be taken into account.' Quasiternary diagrams, though powerful, are often laborious to produce, since several different diagrams are needed for a given set of analytical data in order to draw chemical conclusions on the phases present in ash or agglomerates. We have therefore constructed a special software, that contains several tools for a detailed examination of the data.
EXPERIMENTAL The agglomerate that was examined has been produced in a laboratory-scale fluidisedbed reactor in ETC Pitei, Sweden.' Bed material in these tests was pure quartz, and the fuel was wheat straw. Straw is known to be a difficult fuel because of its h g h alkali metal content (K20content ca. 25 %), since high alkali metal content in fuel has been found to cause agglomeration of quartz- and silica sand bed materials.'0"' The tests at ETC were carried out to effect the agglomeration of the bed. The initial agglomeration temperature was determined by principal component analysis of small variations in the measured bed temperatures and differential pressures that preceded the defluidisation. The agglomeration temperature with the wheat straw fuel was quite low, 739°c.12 The adhesive material of the agglomerated sample was characterised using a Jeol JSM-6400 scanning electron microscope combined with a Link ISIS energy dispersive X-ray analyser. The acceleration voltage of 12 kV and a beam current of 120*10-* A were used for the SEM-EDS-analysis. The sample distance was 15 mm. The image analysis was performed with IMQuant software incorporated in Link ISIS that was used to divide arbitrarily the seemingly homogeneous cross sections of the adhesive material into 685 domains to establish the compositional distribution of the material. All domains were independently analysed for sodium, magnesium, aluminium, silicon, s u l k , phosphorus, potassium, calcium, titanium, and iron. The SEM-EDS-results were visualised and examined in detail by the use of locally constructed software package.
RESULTS AND DISCUSSION The initial visual SEM examination of the bed sample clearly indicated agglomerated material. In this case the automated EDS-analysis could be restricted to determine the ten elements in the adhesive material of the agglomerate and the bed particles could easily be excluded fiom the analysis. Since it is known that high chlorine content in straw causes fouling and corrosion during normal operation of the boiler,' a few point analyses were carried out in the initial examination to explore the chlorine content in the agglomerate. It turned out that there is no or very little chlorine in the adhesive material. The software package used for the visualization of the analytical results contains several modules. It is possible to construct three-dimensional quasiternary diagrams. Each comer can be defined in terms of combined content of several elements. The column heights represent the relative content of the material with the elemental composition as controlled by the definitions in the three comers. By varying the
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definitions in the comers, it is possible to gain insight on the phase composition of the material. The information extracted ffom SEM-EDS data by use of three-dimensional quasitemary diagrams can be complemented by utilizing two-dimensional diagrams and involving the custom colouring scheme. It is also possible to select a region in the two-dimensional diagram and make inferences on the details of the compositional distribution within this region by exploring the original analytical data. A series of three-dimensional diagrams have been constructed ftom SEM-EDS data for the adhesive material of the wheat straw agglomerate (see Fig. 1). The elements chosen for the comers and other relevant options needed for the construction of the diagram can be defined in the initial dialog box of the program (see Fig. 2). The mean values of the contents of each element in all domains are also shown in this box.
a)
b)
Al+
Al+ Na+K+ Ca+Mg+ Fe+Ti%
Na+K+ Ca+Mg%
P+S%
Fe+TI%
C)
4
Al+
S Na+K%
Ca+Mg%
K%
Ci%
Fig. 1 A series of three-dimensional quasitemary diagrams of the same analytical data with various comer definitions. (a) All the analysed elements are defined in the comers of the diagram, (b) P and S are replace by Fe and Ti, (c) Fe and Ti are replaced by Ca and Mg, (d) conventional temary diagram with Si, Ca, and K defining a comer each. The column heights represent the relative content of the material with the elemental composition as controlled by the definitions in the comers.
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Fig. 2 Initial dialog box of the s o h a r e , where the comers and other options are defined for a diagram. In the quasitemary diagrams, the height of the columns is relative to the area of the domains that represent composition controlled by the comer definitions. The first diagram [see Fig. l(a)] includes all ten analysed elements in three comers. The number of included elements is gradually reduced [Fig. l(b-d)] so that in Fig. l(d) there are only three elements shown. The criterium to accept the analysed domains in the diagram is based on the 80% rule.’ As the number of elements defining the comers is reduced, fewer domains are included in the diagram. In the diagram of Fig. l(a) A1 and Si were defined in one comer, Na, K, Mg, Fe, and Ti in the second comer, and P and S in the third comer. Phosphorus and sulfur were defined in the same comer to establish whether the adhesive material of the agglomerate contains significant amounts of these elements. This has indeed been found to be the case. The material distribution is seen as a ridge that diverges from the vertex between the Al+Si% and Na+K+Ca+Mg+Fe+Ti%comers. In Fig. l(b) the same analytical data were considered, but the comers have been redefined by replacing P and S by Fe and Ti. It should be noted that 99.4% of the adhesive material is still included in the diagram of Fig. l(b). It can also be seen that only a few domains contain iron and titanium. The redefinition of the comers in Fig. l(c) separate alkaline earth metals from alkali metals. It can be seen that with this definition the material appears more disperse. It can also be seen that alkali metals are more predominant than alkaline earth metals. There is also small distribution maximum (maximum 1) near to the vertex between the Al+Si% and Ca+Mg% comers. When Al, Na, and Mg are omitted from the defintion, this maximum shifts towards the vertex between Si% and Ca% comers [see Fig. l(d)]. The direction of the shift indicates that the material contains sodium in this region. Fig. l(d) represents a conventional ternary diagram with only one element defining one comer. It is notable that 95.8% of the analysed material is still included in the diagram. While the three first figures [Fig. l(a-c)] indicate the presence of either silicate or aluminosilicate phases, the omission of aluminium in the comer defintion has only a small impact on the compositional distribution. Only one minor local distribution maximum (maximum 2) near to the vertex between Al+Si% and Na+K% comers that is observed in Fig. l(c) is missing in Fig. l(d). It can therefore be concluded that the adhesive material contains mainly silicates. Conversely, the vanished maximum can therefore be inferred to be composed of either aluminates or aluminosilicates. It comprises 16 domains i.e. 3.3% of the analysed material. The 674
original analytical data from these domains (see Table 1) indicate that the material is an aluminosilicatephase containing mainly potassium, sodium, and iron. The information extracted fiom SEM-EDS data by use of three-dimensional quasiternary diagrams can be complemented by utilising two-dimensional diagrams and involving the custom colouring scheme. This is exemplified in Fig. 3 where two two-dimensional quasiternary diagrams with different comer definitions are shown. In Fig. 3(a) the elements P and S in the lower left comer are assigned individual colours black and white, respectively, and these colours are mixed in the ratio of the elemental composition providing areas of different shades of grey. Thus, solid black colour in Fig. 3(a) indicates that the material contains phosphorus but not sulfur. The ridge that was detected in the diagram shown in Fig. l(a) appears black indicating that of the two elements, phosphorus and sulfur, phosphorus is much more abundant (see Table 1). Sulfur, on the other hand, is more abundant than phosphorus in the material appearing along the rightmost vertex [see Fig 3(a)]. The total content of these two elements, however, is very low in this region. In Fig. 3(b) black and white stand for K and Na, respectively. It was earlier deduced from Figs. l(c) and l(d) that sodium is more abundant than potassium in the region indicated by a triangle (maximum 1) in Fig. 3(b). The appearance of white colour clearly verifies this conclusion. The region comprises 18 analysed domains that correspond to 3.2% of the area of the analysed adhesive material. The exploration of the original analytical data shows that these domains contain mostly calcium and not magnesium (see Table 1). AI+Si%
Ai+Si%
/
Ca+Mg%
P4
Na+K%
Fig.3 Two-dimensional quasiternary diagrams with custom colouring. In the diagram a) white stands for sulfur and black phosphorus, in the diagram b) white stands for sodium and black potassium. Sodium containing region is selected and marked as maximum 1 in figure b). The principal maximum appearing in the quasiternary diagrams of Fig. 1 was selected as a region in the diagram of Fig. 3(b). The selection contains 540 domains (81.0% of the material) and mainly consists of potassium silicates with calcium and smaller amounts of other elements. The average elemental composition of this maximum based on the full analytical data of these domains is shown in Table 1.
675
Table I Average elemental compositions of selected regions in the diagrams. wt-%
Ridge see Maximum 1 Maximum 2 Principal maximum Fig. l(a) see Fig. l(c) see Fig l(c) see Figs. l(a-d)
Na Mg A1 Si P S K Ca Ti Fe
233 298 0,4 30,2 93 190 33,6 19,l 093 098
593 293 032 36,9 099 094 797 45,3 0,2 03
1,8 093 7,5 59,3 034 093 27,6 097 094 196
274 193 037 50,2 192 0,4 36,8 691 023 0,6
CONCLUSIONS
SEM-EDS combined with image processing enables automation of the analysis and provides several hundred domains for analysis. This paper describes a locally constructed program that enables the visualization of SEM-EDS data. The production of two- and three -dimensional quasitemary diagrams is easy and the use of multiple diagrams with various comer definitions has proven to be crucial in establishing the composition of ash-related materials. Custom colouring of the two-dimensional diagrams shows details in the data that are difficult to visualize in any other way. The software also enables the investigation of the analytical data from the selected regions either by itself or in relation to the whole data set.
REFERENCES 1 . Huggins F. E., Kosmack D. A., Huffman G. P. & Lee, R. J. (1980) Coal Mineralogies by SEM Automatic Image Analysis. SEM, 1,531-40. 2. Baxter L. L. (1990) The Evolution of Mineral Particle Size Distributions During Early Stages of Coal Combustion. Prog. Energy Combust. Sci., 16,261-6. 3. Laursen IS.,Frandsen F. & Larsen 0. H. (1995) Slagging and Fouling Propensity: Full-scale Test at Two Power Stations in Western Denmark. In. Application of Advanced Technology to Ash Related Problems in Boilers (Ed. by L. L. Baxter & R. DeSollar), Waterville Valley, NH, 03215. 4. Alekhnovich A. N., Gladkov V. E. & Bogomolov V. V. (1995) Slagging Prediction When Using the Chemical Composition of Fly Ash Individual Particles and the Slagging Probability Model. In: Application of Advanced Technology to Ash Related Problems in Boilers (Ed. by L. L. Baxter & R. DeSollar), Waterville Valley, NH, 03215.
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5 . Gupta R. P., Wall T. F., Kajigaya I., Miyamae S . & Tsumita, Y. (1998) Computer-
Controlled Scanning Electron Microscopy of Minerals in Coal -Implications for Ash Deposition. Prog. Energy Combust. Sci., 24,523-43. 6 . Heiklunen R., Laitinen R. S., Patnkainen T., Tiainen M. & Virtanen M. (1998) Slagging Tendency of Peat Ash. Fuel Processing Technology, 56,69-80. 7. Virtanen M. E., Heikkinen R. E. A., Patrikainen H. T., Laitinen R. S . , Skrifvars B.J. & Hupa M. (1999) A Novel Approach to Use CCSEM when Studying Agglomeration in Fluidised Bed Combustion. In: The Impact of Mineral Matter in Solid Fuel Combustion, (Ed. By R. P. Gupta, T. F. Wall & L. L. Baxter), pp. 14754, Kluwer AcadernicPlenum Publishers, New York.. 8. Shah N., Huffman G. P., Huggins F. E. & Shah A. (1991) Graphical Representation of CCSEM Data for Coal Minerals and Ash Particles. In: Inorganic Transformations and Ash Deposition During Combustion, (Ed. By S . A. Benson), pp. 179-90, Palm Coast, Florida. 9. Nordin A., Ohman M., Sknfvars B.-J. & Hupa M. (1995) Agglomeration and Defluidisation in FBC of Biomass Fuels -Mechanisms and Measures for Prevention. In: Application of Advanced Technology to Ash Related Problems in Boilers (Ed. by L. L. Baxter & R. DeSollar), Waterville Valley, NH,03215. 10. Ergiidenler A. & Ghaly A. E. (1992) Quality of gas produce from wheat straw in dual-distributor type flidised bed gasifer. Biomas and Bioenergy, 3 , 4 19-430. 11. Nuutinen L. H., Ollila H. J., Tiainen M. S., Virtanen M. E. & Laitinen R. S. (2000) Role of quartz sand in the agglomeration during the FB-combustion using fuel of hlgh sodium content. Ash deposition: Problems, Management & Solution, 8.-11.5. 2000 Park City, Utah, United States. 12. Ohman M., Nordin A., Skrifvars B.-J., Backman R. & Hupa M. (2000) Bed agglomeration characteristics during fluidized bed combustion of biomass hels. Energy &Fuels, 14, 169-178.
677
Model and Simulation of Heat Exchangers and Drying Silo in a New Type of a Boiler Plant J .Yrjola Satakunta Polytechnic, Technology, Pori, Tekniikantie 2, 28600 PORI, FXVLAND
ABSTRACT
A new type of a boiler plant is presented, in which it is possible to use both dry and wet biofuel also in small scale. To achieve that benefit, the new plant contains three additional components not present in a traditional plant designed for only dry or wet biofuel. The additional components are a flue gas heat exchanger, a drying silo and a heat recovery heat exchanger for the exhaust air of the drying silo. A general view of simulation models for these components is given and some calculation results are discussed in the context of the developing work. Verifying measurements of the model are also shown. INTRODUCTION Nakkila Works Ltd. and Satakunta Polytechnic have co-operated in developing a new type of a boiler plant (Fig.I). m.,w
U D CDIIUITIrn
LlS r W
m W
",US,
r
Fig. I The scheme of the new boiler plant.
678
**I)
IiLLUiilPl
I
In this system the fuel (chips, bark, sawdust) is dried before it is led into the combustion chamber of the boiler. The drying air is heated in a recuperative heat exchanger (Fig. 4 ) by the heat of the combustion gases. Hot air is then blown through the bed of fuel in the drying silo (Fig. 12), while the fuel dries and the air cools and humidifies. Heat of the exhaust air of the silo is recovered for the drying air and combustion air by a recuperative heat exchanger (Fig. 9). The boiler is an ordinary warm water boiler for dry biofuels. The most essential benefit of the system is the possibility to use both dry and wet biofuel without lower limitation in nominal heat output. In the new type of a boiler plant three additional components not present in a traditional plant designed for only dry or wet biohel are needed. These components are a flue gas heat exchanger, a drying silo and a heat recovery heat exchanger for the exhaust air of the drying silo. Modelling of those components is necessary in the dimensioning of the plant. Because the plant operates only a part of the year at peak load, it is also important to know the behaviour of the system in changing environments. In order to select an economical and efficient combination of components, knowledge of the effects of changes in size or in geometry of the main components is needed. The first step in simulation is the heat balance calculation. The procedure of the calculation is shown in Fig. 2. Dimensioning calculation procedure
Simulating calculation procedure
Fig. 2 The mass flows and their state in different points of the system are results of the mass end energy balance calculations. They are input data of the dimensioning of the components. After dimensioning it is possible to simulate the behaviour of the system in different operational conditions.
679
The mathematical models of the components are carried out in $e Excel-spreadsheet computation with Visual Basic macros. The theory of the models of is not new, but the combination of these models is not presented before. The calculation results are verified with measurements in a 40 kW,h test plant and in a 500 kWh demonstration plant Calculation of combustion processes, adiabatic combustion temperature and heat and mass balances are not discussed in this paper. The moisture content of the fuel has an essential significance in developing the new system. As an example of the calculation possibilities the efficiency ratio of a boiler as function of the temperature of flue gas with the moisture content of fuel as a parameter is shown (Fig. 3). The results are similar to [ 11. I4
I3
Fig. 3 The efficiency decreases, when the combustion air ratio or the moisture content of the fuel increases. However, if the outlet temperature of the combustion gases is below the dew point of the water vapour, the efficiency calculated according to LHV (lower heating value) begins to increase, when the moisture content of the fuel increases. Temperature of combustion air is 0 "C, RH 60%, excess air ratio 1.4 and composition of dry fuel is CH1.466N~.~~90~.633.
MODELLING AND MEASUREMENTS OF THE COMBUSTION GAS HEAT EXCHANGER The heat exchanger is of shell-and-tube type, where the combustion gases flow in tubes and the heated air in the shell. This construction was selected because of the easy removal of deposits. The principle scheme is shown in Fig 4. The heat exchanger is divided into six computational cells, in which the mass and heat balances are solve. The output data of one cell is the input data of the next cell [2]. The heat transfer coefficient inside the pipe is solved as a forced convection by means of the correlation formulas of pipe flow presented in the literature [3].The effect of the condensing water vapour is taken into account [4]. The heat transfer coefficient outside the pipe is calculated according to the procedure of pipe bundle heat exchangers in [3]. The properties of substances are calculated in the middle of the cell. Since the heat exchanger is of crossflow-counterflow type, the state of the gas leaving the heat exchanger is solved iteratively. Also the pressure drops of the flows are calculated according to [3].
680
Fig. 4 Construction of combustion gas heat exchanger.
Input data are the geometry of the heat exchanger (the pipe dimensions and locations, the information of baffles and surface deposits) and the properties of fluids (elementary analyses, massflows, temperatures entering the heat exchanger). Output data are temperatures of the fluids leaving the heat exchanger, heat output, condensing mass flow, pressure drop, surface temperatures and the temperature distribution etc. Some calculation results are shown in thefigures 5 and 6. 4W %
140
350 %
120
3W% 1W
E
250 %
-Power
80 0
2W %
60 150
k 2
x
of heal exchanger Pressure drop. nw gas Pressure drop. an +Temperalure of air ~t Temperature of flue gas
1
1
40
1W %
20
50 %
0
0% 0
0.2
0.4
0.6
0.8
1
Proportional rate of nuaa n
1.2
1.4
1.6
1.8
2
a of n w r s and air
Fig. 5 Behaviour of Flue Gas Heat Exchanger. Relative heat output (-), pressure and drops (flue gas - - - - and air - - - ), outlet temperatures of gases (flue gas -Aair-) as fbnction of the rate of mass flow of air and combustion gas (changed in the same proportion). The change in heat output is almost linear. When the rates of mass flow are reduced by SO%, the heat output is reduced by 44%. Also the drying air outlet temperature increases and the flue gas outlet temperature decreases from the initial values. Pressure drops for primary and secondary flows decrease about 15%. The heat output does not decrease steeply even if the flow becomes laminar. ~
68 1
---Heat exchange surfacs area . . Pressure dmp. air -.-.Pressure dmp. Rue gas
.. .
Number Velocity Heat exchanger of pipes comb. eff. air gas
70
d s
1 2 3 4
5 6 1
2
3
4
5
50 35 25 45 30 20
d s
1.80 7.38 10.55 4.14 14.80 10.73 1.32 11.63 17.43 3.95 26.16 10.78
6
Fig. 6 Heat exchangers of different size yielding the same heat output with the same mass flows and inlet temperatures. In the heat exchangers 1,2 and 3 the tube diameters (outer/inner) are dJdi=60.3/54.5 mm and in 4, 5 and 6 51145.8 mm. Number and location of tubes vary yielding different velocities of combustion gas and effective velocities of air (see table next to diagram).
The same heat output is achieved with a heat exchance area of 85 m2 as with 29 m2, though the air-side pressure drop increases from 100 to 6500 Pa and the flue gas-side pressure drop from 100 to 1600 Pa. The greater increase of air-side pressure drop derives from the fact that in addition of changing the number of pipes also the distance between pipes was changed. Infigures 7 and 8 the comparison of calculated and measured values are shown.
Fig. 7 Comparison of measured temperatures and temperatures calculated with the model. The measurements are done at the demonstration plant of Pori Forest Institute and the test plant at Nakkila Works, Terasmaki. The uncertainty of the measuring device is f 0.5 "C.
682
im IW
M
I
I
I
I
I
I
0 005
0
01
015
02
025
A t hImYsl
[-Calculated
values ..m.. Measured values
I
Fig. 8 Comparison of measured and calculated pressure drops. The measurements were carried out at the demonstration plant of Pori Forest Institute and the pilot plant at Nakkila Works, Terasmaki. The uncertainty of the micromanometer is f 1 Pa and of the volume rate measuring device f 0.5%. MODELLING AND MEASUREMENTS OF THE HEAT RECOVERY HEAT EXCHANGER The heat recovery heat exchanger is of flat-plate type (Fig. 9). This construction was selected because of the low investment costs. Primary flow is moist, nearly saturated exhaust air from the drying silo. The secondary flow is outdoor air, which is used as combustion air and drying air.
,
Exhaust air from silo
'
Supply air to flue gas heat exhanger Exhaust air out
Condensate out
Fig. 9 Heat recovery heat exchanger of exhaust air from drying silo. The heat exchanger is divided into 3 computational cells in parallel-flow and counterflow types, and into 9 cells in crossflow type. The mass and heat balances of the cells are calculated by means of heat balance of a single plate [3,4]. The procedure of the calculation is similar to the case of the flue gas heat exchanger. The surface temperatures and air temperatures leaving the different types of heat exchangers are shown in Fig. 10. The pressure drop of the heat exchanger is calculated according to
PI. 683
Fig. I0 Outlet temperature and surface temperature in different types of heat exchangers. The outlet supply air temperatures are shown on the left hand side and the lowest surface temperatures (in the middle of the computational cell) on the right hand side depending on the entering outdoor air temperature and the type of the heat exchanger. The rates of mass flow, heat transfer area and the properties of the entering moist exhaust air of the drying silo are constant. According to Fig. 10, the outlet air temperatures are rather close to each other with different type of heat exchangers. The temperature difference between the counterflow and parallel-flow is 1 "C, when the outdoor air temperature is 0 "C and 2 "C when 4 0 "C, thus the heat outputs are also almost equal. The crossflow heat exchanger is between the types above. Regarding freezing the differences are more important. The critical outdoor temperature, when the lowest surface temperature falls below 0 "C, is -26 "C with a crossflow, -28 "C with a counterflow and below -40 "C with a parallel flow heat exchanger with the critical rates of mass flows. The comparison between calculated and measured temperatures of the fluids leaving the heat exchanger are shown in Fig. I I. The heat exchanger is of crossflow type.
Fig. I I Comparison of measured and calculated outlet temperatures of heat recovery heat exchanger. The measurements were carried out at the demonstration plant of Pori Forest Institute. The uncertainty of the measuring device is f 0.5 "C. 684
MODELLING AND MEASUREMENTS OF THE DRYING SILO The fuel is led into the drying silo Fig. 12 by the conveyor and spread evenly by the screw. The fuel then flows down gravitationally and is led into the boiler by the screw feeder. The cross-section of the silo is rectangular and the longer sides are constructed from a hole plate with 3 mrn holes. The hot drying air coming from the combustion gas heat exchanger is blown through the holes and the fuel bed. Water vapour evaporating from the fuel particles cools and humidifies the drying air so that it is almost saturated when leaving the drying silo. The construction of the silo is based on the experimental study concerning the 40 kW pilot-plant, where several alternatives were investigated 161.
1
Moist biofuel
Dried biofuel
Fig. 12 The scheme of the drying silo in the demonstration plant Pori Forest Institute. Modelling is divided into two separate parts: calculation of the dimensions of the silo and the simulation of the thermal behaviour of the silo.
DIMENSIONING The aim of the dimensioning is to give enough information to define the main dimensions rapidly, and in easy-to-use way. The thermal operation of the silo is simplified so that the stages of heating and melting of ice in particles, heating of the particles and evaporating water from the particles are assumed to appear sequentially. In reality these stages are partly simultaneous because of the different sizes of particles and the shape of the silo. Small particles may be almost dry, while the the large ones include plenty of water while being almost dry near the surface, even though they are still icy in the middle. In the front of the silo the particles are rather dry, while in the back they still are icy or very moist. In the stages of heating and melting the silo is treated as a heat exchanger so that the temperature of air and fuel leaving the silo are input data. The structure of the dimensioning model is shown in Fig. 13.
685
moist biofual
PI
b/ ,/
to
Xl
I
xs
Fig. I3 The principle of dimensioning calculations of the drying silo. The mass flows qnl...qa of air needed for heating and melting the ice, heating the wood and remaining water to the final temperature (b)are achieved by energy equilibrium calculations with mass flow, initial and final moisture content (aIand a2respectively) and temperature of fuel as input data. The mass flow of air (qa4)needed to evaporate the water (Aq,) is calculated by assuming the outlet air of the evaporation stage saturated (cps) and the state of the drying air changing along the curve of constant enthalpy. The total mass flow of air needed is the sum of the mass flows of different stages added with the mass flow of evaporated water. The outlet temperature of air ( k x 2 ) is the mixing temperature of the outlet temperatures of different stages.
The energy demand of the heating and evaporating stages depending on the temperature of outdoor air is shown in Fig. f4.The mass flow and the heating effect of the drying air depending on the temperature of the drying air is shown in Fig. f5. .so
.so
1
Fig. 14 The components of heat power demand of drying air in a 1 MW boiler plant at different outdoor air temperatures. No heat recovery is used, the initial and final moisture contents of fuel are 60 % and 35 % and the temperature of drying air is 90 "C. The bars on the left represent the power needed to heat up the air and the bars on the right the power demand of the different stages strictly.
686
The demand of heat for the different stages changes a little depending on the temperature of ambient air. The only greater change is at temperature 0 "C, because of the melting heat of ice. However, the outdoor air must be heated to rather high temperature to be used for heating and evaporating purposes. That is why the heat demand for warming up the outdoor air depends strongly on the ambient temperature. Because of that fact heat recovery is necessary in the system. The enthalpy of the exhaust air from the drying silo is high and the heat exchanger gives the more heat output the lower the outdoor temperature is (Fig. 10) to cover the increasing heat demand. 250 x
200 %
150 %
drying air
100 %
50%
0% 0
20
w
40
loo
80
120
140
160
Tempenturn d drying air [C)
Fig. 15 The rate of mass flow and heat power demand for heating of drying air as a function of the temperature of drying air. The needed rate of mass flow of drying air increases when its temperature decreases, but the needed heating power increases less. According to Fig. 15, the needed massflow of drying air increases by 140%, when its temperature decreases from 90 to 40 "C. That is because of the ability of air to contain water vapour decreases. At the same time the needed heating effect of air decreases 8%. The increasing rate of mass flow of air increases the power input of fans.
SIMULATION AND MEASUREMENTS The silo is divided into cells containing fuel particles, which are handled as spheres of equal size, and their effective diameter is calculated according to [7]. The evaporative mass flow from the wet surface of the particle is calculated [4]: - w Ld" m =h-Mvln-
CPm
P -P",, P-Pv
687
The energy balances are solved numerically for a computational cell:
h,,A(T'
- T , )= mpcpdTP + riz"z(u)Al,+ Ij2"Acpvd(Ta - T,) dt
The term on the left describes the heat of the gas to the particle. Radiative heat transfer is excluded. The first term on the right describes the required heat for increasing temperature of the particle. The second term shows the heat needed to cover the latent heat of evaporating water. The effect of bound water is excluded, because it is small compared to heat of vaporization [8]. Drying the fuel particle is described with an experimentally defined function z(u). The last term describes the heat needed to rise the temperature of water vapour from evaporation to leaving air temperature.
Fig. 16 The moisture content distribution of the fuel particles and the temperature distribution of the drying air. In this calculation the silo (height 2,O m,width 1,0 m) is divided into 1000 calculaton cells. The moisture content of entering fuel is 60 % (wet basis). However, on the right side of the silo there is an area, where the moisture content is greater than 60% (wet basis) because of condensation. Drying of wood bed in hot air was measured in laboratory with different kinds and particle sizes of wood chips.
688
Wood shlps (fir). p.rtisle
dze 30 mm
80
I
0
0
20
40
60
80
1W
120
140
Urn. Imin]
Fig. 17 Comparison of measured and calculated values of temperatures and the moisture content of wood in experiment of fixed bed. The function z(u), which is experimentally defined to describe the mass transfer inside the particle, is chosen so that the most important parameter, the speed of evaporation of water fits well with the measurements of the bed. That can be seen in Fig. 17, in changes of the calculated and measured moisture content of wood in the bed. OPERATIONAL ECONOMICS The models of components are joined together so that it is possible to simulate the whole Ecoflame-system (Fig. 2). In this paper the model of the boiler is not described. However, some computational results concerning operational economics are shown. It is also possible to simulate the behaviour of the traditional systems, where the fuel is led into boiler without drying and no heat is recovered from the combustion gases after the boiler (Fig. 18).
Fig. 18 The behaviour of two base-load plant in different outdoor temperatures. The efficiency (useful heat outputheat input of the fuel) of the traditional boiler system and the Ecoflame-system are near the same, on the left. These base-load plants are dimensioned to produce 50% of the peak-load in the outdoor temperature conditions of
689
the Western Finland, and so the relative thermal output begins to decrease at higher outdoor temperatures than 0 "C. On the right is seen that the flue gas temperature is kept constant at 110 "C in the Ecolame system, but it changes from 100 to 160 "C in the traditional system. The moisture content of the flue gas is smaller in the Ecoflamesystem than in the traditional system, because a part of the moisture of the he1 leaves with the exhaust drying air. The annual fuel consumption of the Ecoflame system is slightly smaller than that of the traditional system, but the need of the electrical energy of the combustion air fan is bigger, because this fan also supplies the drying air. The efficiency of the system depends among other things on the temperature of the flue gas. In this comparison the temperature of the traditional system is nearly as low as possible without any condensation in the chimney at partial load. In practice this temperature is often in the range 220 ... 280 "C at the dimensioning point. In these cases the operational costs of the Ecoflame system are smaller than of the traditional system. However, the most essential benefit of the Ecoflame-system is the possibility to burn both dry and wet biofuel.
690
CONCLUSIONS A new type of a boiler plant is presented, in which it is possible to use both dry and wet biofuel also in small scale. The fuel is dried before it is led into the boiler. Drying is carried out in a silo, where hot air is blown through the bed of fuel. Drying air and combustion air are preheated in a heat recovery heat exchanger by the exhaust gas of the silo and heated in the flue gas heat exchanger. In developing the new system a dimensioning and simulation model was developed. A general view of the models is given for three main components: the flue gas heat exchanger, the drying silo and the heat recovery heat exchanger for the exhaust air of the drying silo. The flue gas heat exchanger is of shell-and-tube type, where the flue gases flow in the tubes and the combustion and drying air in the shell. The heat output of the heat exchanger changes almost linearly with the changes in the rates of mass flow. No sudden decrease in the heat output is detected even if the flow becomes laminar. Many combinations in the heat transfer area and in the geometry of the heat exchanger give the same heat output and the same outlet temperatures at the same inlet conditions. From the changes in the pressure loss of gas flows follows, that there is an optimum heat transfer area regarding to the life-cycle cost of the plant. The heat recovery heat exchanger is of flat plate type. Heat recovery is necessary because of the large heating demand of the drying air. The heat output is nearly the same for parallel flow, crossflow and counterflow types of equal size in the temperature conditions of heating season. That is due to very moist primary flow. Almost the same amount of water vapour condensates in all types of flow arrangements. The heat output differs more at colder outdoor temperatures. The heat output between parallel-, cross- and counterflow heat exchangers differs more with increasing heat transfer area. On the other hand the differences in risk for freezing are greater. Crossflow type is most risky, counterflow is the next and the parallel-flow type heat exchanger has practically no risk. In the new system the parallel-flow heat exchanger is chosen despite the small loss in heat output. The very small risk in freezing in Finnish climate was valued more. The drying silo is constructed of sheet metal. The silo has two screws, one above to spread the fuel evenly and one below to empty the silo. The hot drying air is blown through the fuel bed in the silo. The drying air flow rate needed increases when the temperature of it decreases, because of the smaller ability to contain water vapour. Evaporation of water consumes the most of the heat demand of the silo. The heating and melting of ice and heating of wood need less heating. The simulation calculations were verified by field measurements in the test plant at Nakkila Works and in the demonstration plant at Pori Forest Institute. The calculation results seemed to be accurate enough for practical dimensioning and designing purposes. The most essential benefit of the system is the possibility to use both dry and wet biofuel without lower limitation in nominal heat output.
ACKNOWLEDGEMENTS Support from the Technology Developing Centre of Finland (TEKES), contract 70130/99, and from the Nakkila Works are acknowledged.
69 1
NOTATION
A CP, CP
h 1" Le M m m" n P T t U
z(u) Subscripts a con P S
V
total particle surface area in the cell, m2 molar heat capacity, J/mol K spesific heat capacity, J k g K heat transfer coefficient, W/m2K heat of *iaporization,J k g Lewis number molar weight, g/mol mass mass flow rate/area, kg/ m2s exponent pressure, Pa temperature, "C time, s moisture experimental function air convection particle surface vapour
REFERENCES 1 . Neuenschwander P., Good J., Nussbaumer Th. (1998) Grundlagen der Abgaskondensation bei Holzfeuerungen, Bundesamt f3r Energie, Zurich 2. Hewitt, Geoffrey F. et a1 (1998)., Heat Exchanger Design Handbook, Begell House Inc., New York 3. VDI-GVC (199 l), VDI-Warmeatlas: Berechnungsblatterfr den Warmeubergang, 6. erweiterte Auflage, VDI-Verlag, Diisseldorf 4. Lampinen M. ( 1999), Aineensiirto-oppi, Helsinki University of Technology, Otaniemi 5. Mills, A.F. (1998) Heat Transfer, Prentice Hall, New Jersey 6. Honkasalo M., Yrjola J. (1997) Kattilalaitosteknila kuiville ja kosteille biopolttoaineille. In: Bioenergian tutkimusohjelma vuosikirja 1997 osa III, pp. 169-179. Jyvaskyll 7. Saastamoinen J., Impola R. (1995), Drying of Solid Fuel Particles in hot Gases, Drying Technology, 13, pp, 1305-1315 8. Groenli, M. (1996), A Theoretical and Experimental Study of the Thermal Degradation of Biomass, PhD. thesis, The Norwegian University of Science and Technology
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Testing & Modeling The Wood-Gas Turbo Stove T. B. Reeda9b, E. Anselmo” and K. Kirchef a The Community Power Corporation, 8420 S. Continental Divide Rd., Su 100, Littleton, CO 80127; The Biomass Energy Foundation, 1810 Smith Rd., Golden, CO 80401; The Colorado School of Mines, Golden, CO 80401
ABSTRACT Through the millennia wood stoves for coolung have been notoriously inefficient, unhealthy and slow. A new “wood-gas” cook stove has been developed that has >30% thermal efficiency, can be started, operated and stopped with very low emissions and can use a wide variety of biomass fuels. This “Turbo Stove” operates with 3 W of blower power or other air supply to produce 1-3 kWtheml for coolung. It is simple and inexpensive to build. Data is presented for this stove on a wide variety of fuels. The stove will bring a liter of water to boil in 4-10 minutes and can be turned down to the simmer level for longer cooking and increased efficiency. The stove operates in several different gasification and combustion modes. In the “volatile burning” mode, the stove makes 18-25% charcoal fiombiomass fuels. In the “charcoal burning mode” the charcoal is gasified to produce a CO flame. If longer cooking is required, additional fuel can be fed fiom above, but other modes require more operator skill. For a new stove to be accepted it must fit the fuel supply, cooking practices , construction methods and commercial infrastructure of each country. Therefore, it must be possible to make a variety of stoves and requires understanding of the basic mechanisms of gasification and combustion of “wood-gas”. A model of the wood-gas “Turbo Stove” is described based on the measured parameters in this paper.
INTRODUCTION - WOOD COOKING VS WOOD-GAS COOKING Since the beginning of civilization wood and biomass have been used for coolung. Still, today over 2 billion people cook badly on slow, inefficient wood stoves that waste wood, cause health problems and destroy our forests. Electricity, gas or kerosene are preferred for cooking - when they can be obtained However, they are costly, contribute to global warming, and depend on having a suitable infrastructure often not available in developing countries
693
In the last few decades, many improved wood stoves have been developed (the Chula, the Hiko, the Maendeleo, the Kuni Mbili, the Wendelbro, etc.’ These new wood stoves are often more difficult to manufacture and they do not offer good control of cooking rate. They are often not accepted by the cooks for whom they are developed. Since 1850 the preferred means of cooking has been first gas, then electricity. Gas is still preferred by many cooks. Electric cooking can be 60% electric-efficient, but power generation and distribution is typically 30% efficient, yielding an overall efficiency of 18% for electric cooking. We have developed several simple, inexpensive wood-gas stoves which can bring the “joy of coolung with gas” to everyone while using a wide variety of renewable biomass fuels or ~ o a l . ~ ‘ ~
PRINCIPLES OF DOWNDRAFT GASIFICATION FOR COOKING BIOMASS GASIFICATION
When biomass is burned with insufficient air in a gasifier, it makes a “producer gas” containing primarily CO, H2, C02, H 2 0 and CH4.Over a million gasifiers powered the civilian cars and trucks of Europe and Asia during WW 11. Downdraft gasifiers are “tar-burning, char-making” and are most suitable for biomass which contains 80% volatile material. Updraft, “char-burning, tar-making”, gasifiers are often used for coal which can be 80% char. In conventional downdraft gasifiers, air passes down through the fuel mass, then in the flaming pyrolysis zone burns the volatiles and tars while making charcoal and pyrolysis gas. The charcoal then fiu-ther reduces the C02 and H 2 0 combustion products back to CO and H2 fuel. THE “INVERTEDDOWNDRAFT GASIFIER”
In inverted (top burning) downdraft gasification air passes up through the fuel and meets the flaming pyrolysis zone where the reaction generates charcoal and fuel gas as shown in Fig. 1.2,3 NATURAL VS FORCED CONVECTION
Natural convection provides poor mixing of air with fuel gases and can result in incomplete combustion, soot and emissions in open wood stoves. A chimney can supply 1 rnm water pressure per meter of height. Addition of a chimney for cooking can greatly improve wood combustion in closed models, but also adds complication and requires wasting heat to operate. Forced convection provides good mixing and combustion for gas cooking and is widely used in homes and camping stoves. The 3 W blower used in the Turbo Stove provides 7.5 mm water pressure and makes clean cooking possible.
694
Wood
I
I
Fig. 1 Natural convection gasifier stove Fig. 2 Forced convection Turbo Stove made with 15 cm riser sleeve293 with 3 kW flame fiom a 3 W blower4 THE “TURBO STOVE”
The Community Power Corporation and the Biomass Energy Foundation have developed a new “Turbo wood-gas stove” using forced draft fiom a 3 Watt blower. One design is shown in Fig. 2. It consists of an inverted gasifier close coupled to a burner section to mix air and gas and bum cleanly. A 3 Watt blower generates 7 mm water column pressure, equivalent to the draft of a 7 meter chimney. We have made it from an outer 1 gal paint can, an inner burner can and a fuel magazine or with many other construction method^.^ Several burners can be assembled to make a cooking “range”. An oven can be placed on one of the burners for oven heat. The stove can be started and operated indoors with no exhaust fans and no odor of burning wood. We have taken the stove to India and the Philippines and cooked with the Turbo Stove in small villages and on conference room desks with no odor. While the Turbo Stove currently uses a 12 Volt 3 Watt blower, the power could come from stored compressed air, bellows, wind-up generators, photovoltaic, thermophotovoltaic, windup motors, thermoelectric or other sources.
-
CONSTUCTION AND OPERATING THE TURBO STOVE THE RESEARCH TURBO STOVE
The research Turbo Stove shown in Fig. 3 consists of An inverted downdraft gasifier and fuel magazine 0 A combustion section which bums the gas 0 Supports for a pot 0 Regulated air supply for gasification and combustion
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as shown in Fig. 3. T h s permits independent adjustment of the air to the gasification section and the combustion section for optimizing cooking conditions at both hlgh and low levels. The rate of heating and boiling was used to measure the heat transfer for cooking. Draft meters were used to measure the pressure drop for gasification and for combustion (typically 0.25-0.75 mmw for gasification and 2.5 mmw water pressure for combustion). Water Calorimeter
Turbo Stove Flames 4ir Jets
Flowmeten
:harcoal
Fig. 3 - Research Turbo Stove, showing separate supplies for gasification and combustion air. THE PROTOTYPE STOVE Compressed air and flowmeters do not a practical stove make. We have also built the prototype stove shown in Fig. 2 that is easier to operate and less expensive. It permits adjusting the power level by adjusting the gasification air. Some of the data below were taken on the research and some on the prototype stove.
STARTING AND OPERA TING THE STOVES In a typical run,the stove is filled with weighed pellets of the dry fuel of choice. A layer of starting chips, (chips, charcoal, or other porous materials soaked in alcohol, fat or kerosene) is placed on top. The blower is turned on and the starter chips are lit with a match. For the first few minutes the starter chips ignite the fuel below and make a bed of charcoal that the gas must pass through. In 1-5 minutes, depending on the fuel, the main fuel mass is ignited and burns downward regularly in flaming pyrolysis mode until the reaction zone reaches the grate, making charcoal as it goes. The test variables are shown in Fig. 4 for the research stove and Fig. 5 for the prototype stove.
696
600
500
100
0
-
Time rnin
Fig. 4 Typical operating data on the research Turbo Stove showing weight of fie1 remaining vs time at hgh, medium and low power levels
Table 1 Air Fuel Ratios for gasification and combustion, power levels, turndown and superficial velocity for research stove Fuel rate POWER Maximum Medium Low
g/m 11.3 7.1 4.6
Gasif Combustion air air g/m 18.0 6.0 3.5
g/m 56.6 34.0 21.7
Gasif Combustion Power Turnair/ down air1 Fuel Fuel kW PPmax 1.59 5.01 2.83 1.00 0.84 4.79 1.78 0.63 0.75 4.73 1.15 0.41
697
Sup Vel m/s 0.062 0.028 0.017
1000 ,-,.. ...,........., .... ....
I
................ 1. ....... ......... ...... ...../ .......
~ . . ....... . ...... .., T'
"
~
....... .....I
-
900 BOO
700
'
600
u)
5
500
400
300 200 100
0
10
20
30
-
40
50
60
70
nme mln
Fig 5 Prototype Turbo Stove data on peanut shell pellets, showing water boiled, water temperature, and fuel remaining at high, medium and low gasification and combustion of volatiles and charcoal cooking after 3 1 minutes
TEST PROCEDURE In operation the Turbo Stove is operated on a balance and the loss of weight of fuel and water are recorded as test cooking progresses as shown in Figs. 4 and 5. The stages in cooking are shown in Table 2. In some cases the run is stopped after the flaming pyrolysis zone reaches the grate and the volatiles have been burned. However, at this point the charcoal begins to be gasified by the incoming air and cookmg can continue until all the fuel is gone. DATA O N THE TURBO STOVES
BEHAVIOR OF VARIOUS FUELS A typical set of data for the research and prototype stove are shown in Figs 4 and 5 and data collected on these runs are shown in Tables 1 and 2. The Turbo Stove has been satisfactorily operated on dozens of fuels. The behavior of six fuels tested more extensively is shown in Table 3. In addition to many biomass forms, coal was also found to be very a satisfactory fuel for the Turbo Stove. Some of the tests done for thls paper were made on '"h inch" (6.2 mm) stove sawdust pellets, a readily available, high density, reproducible fuel except where otherwise noted. Table 1 and Fig 4 show data taken on stove pellet fuel made from sawdust. Pellets are widely marketed for pellet heating stoves in the U.S. for $2.50 for a 20 kg bag. This would be sufficient to cook 40 typical meals. Figure 5 and Table
698
2 were made using peanut shell pellets, a potentially good fuel wherever peanuts are produced.
FUEL CONSUMPTION RATE AND COOKING POWER The fuel consumption rate is a direct measure of cooking power, provided all the gas is subsequently burned in the burner section. The heating value of most biomass with 5-10% moisture (Denver dry) is -18 kJ/g. The stove typically produces 20-25% charcoal after the volatiles have been burned. The charcoal typically has a higher heating value of -24 H/g. In the tables the power level is calculated from these values. A gasification rate of 10 g/m gives 2.5 kW, comparable to the large burner on modem gas or electric stoves.
TURNDOWN RA TI0 An important criteria for successful stove cooking is the “turndown ratio” of the stove.
Initially the stove should develop high power to bring water or oils to coolung temperature. After cooking temperature is reached it is desirable to turn the power level down to just maintain that temperature. The turndown ratios for wood and peanut shell pellets are shown in Tables 1 and 2.
Table 2 Stages in test of peanut shell pellets (Fig. 4), showing flame intensity, water evaporated, turndown and efficiency at various stages of heating water Time Condition GasifFuel Water Heat Heat Effic- Flame Turnication remain lost used topot iency intensity down air mins Turns g g k J kJ kW 2 530 0 0.0 Startup 3300 909 28% 4.1 100% Pot on 2 520 1.o flame 2 420 Rocking 7.2 Boil 1/2 310 230 13.5 Turn Down 1050 253 24% 3.2 78% 1 Medium 1/4 240 340 19.0 Turn down 1650 506 31% 2.3 56% 1 Low 1 130 560 31.0 OdY charcoal lef:
4, 64.0
2160
Off
summary
1 NA
40 490
930 930
851 39%
8160 2518
31%
1.1 28% 2.1
31%
It is desirable to be able to operate at lower power levels than maximum and a major advantage of the Turbo Stove is that it can be operated at lower powers reducing the air for gasification and burning less gas. For most biomass the energy content is 18 kJ/g for fuel with about 10% moisture. In the volatile burning mode the
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volatile energy content is about 15 H/g, whle the 25% charcoal contains 24 W/g. Calculations included here have been made on this basis.
BOILING EFFICIENCY There has been entirely too much emphasis placed on “coolag efficiency”, but it is certainly one importuned factor in evaluating stoves. Equally important is low emissions, h g h intensity and good turndown ratio. The data shown in Table 3 for boiling efficiency was calculated from the ratio of energy used for boiling water (in an 18 cm diameter pot filled with lliter of water) to the fuel consumed after the water reached the boiling point until the end of volatile burning. It typically ranges from -40%, depending on fuel, length of boiling time, pot s u e and other factors.
Table 3 - Operating and derived data for tuns on selected fuels Test fuel
’
Peanut Wood Coconut Palmnut shell pellets shell shell 5/25 519 4/22 4/27
Wood Coal chips 4/26 4/28
Test Date FUEL DATA 6.4 6.5 6.2 7.8 3.1 Moisture Content 6 Fuel Wt. -g 500 305 150 180 260 500 0.64 0.48 0.26 Fuel Density-g/cm3 0.58 0.265 1.69 RUN DATA 15 37 41 Volatile burn time19 13 30 410 215 130 Volatiles burned g 490 150 150 6.0 10.0 Time to Boil min 7.2 7.0 13.0 8 Charcoal yield - g 130 30.0 130.0 90.0 90.0 20 850.0 220.0 100 Water boiled - g 930 145.0 850.0 DERIVED DATA Charcoal Yield-%4 17 18 29 13 50 26 Boiling Efficiency 31.8 37.5 33 31 20 24 2.5 2.4 Average Intensity 2.1 2.5 2.8 2.5 ~. Notes: (1) The peanut shell pellets were 3/8” diameter from Birdsong Peanuts, Georgia; the wood pellets are standard % inch wood heat pellets from Ace Hardware; the coconut shells were obtained from the Philippines hammermilled to 1 corn on an edge; the palm nut shells were obtained in Indonesia; the wood chips were mixed tree chips from Denver 2 cm on an edge; the coal was bituminous from Denver crushed to 2 cm on an edge. ( 2 ) Moisture in wt %, wet basis; (3) Apparent fuel density calculated from magazine dimensions and weight; (4) Charcoal yield calculated from char remaining at end of run and initial fuel weight; (5) Boiling efficiency calculated from fuel consumed during steady boil, based on 15 H/g for biomass volatiles, 35 H/g for coal volatiles; (6) based on 15 kJ/g for volatiles and bum time of volatiles.
’
-
EFFECT OF FUEL MOISTURE CONTENT The fuel moisture content is recommended to be <20% (wet basis) for the operation of engines. However, we have found that fuels with up to 30% moisture can be used in the Turbo Stove quite satisfactorily. We believe that this is because it is necessary for
700
each layer of fuel to ignite the next lower layer. When the layer is dry, the fire propagates easily, but with wet fuel more charcoal is consumed to dry the layer before the reaction can proceed. T h s is born out by the fact that with bone dry fuel charcoal often exceeds 25%, but with 30% moisture fuel only 4% charcoal remained after the pyrolysis was complete. AIR-FUEL RATIOS FOR GASIFICATION The airhe1 ratio is a very important criterion for solid, liquid and gas fuels since there is only one theoretical value that produces maximum flame temperature and minimum emissions. Table 1 shows the air/fuel ratio (based on the sum of gasification and combustion air) for three conditions, the maximum, medium and low output as 4.5-5.7. The theoretical value for “typical dry biomass is 6.3 but depends on fuel composition and moisture content. The aidfuel ratio is an important parameter in the clean gasification and combustion of all fuels including biomass, charcoal and coal. The Aidfuel ratio was measured for wood pellets in the research stove and is shown in Table 1. GAS HEATING, VALUE The gas heating value of raw producer gas containing significant condensable volatiles (tars) is difficult to measure, since measurements are usually made at room temperature after the tars have been removed. The gas higher heating value varies with the aidfuel ratio used for gasification and the superficial velocity. We are in the process of measuring it, but we expect is to vary in the range 5-7 MJ/Nm3. DIFFERENT MODES OF TURBO STOVE OPERATION
Coolung is typically a batch process and successful operation of the Turbo Stove requires the cook to estimate how long a particular task will require. If further cooking or water heating is required, there are several ways to extend the heating time. On startup, the gasifier converts the biomass fuel volatiles to gas, whch is burned and leaves behind up to 30% charcoal whch can be saved or burned for additional cooking. In the stove described in Table 3 times of 19-37 min were recorded for the various fuels. After the volatiles have all been burned a dramatic change in the flame occurs and with the air then gasifies the charcoal to CO, giving a different flame and much hotter temperatures at the grate. If even longer cooking is required, more fuel can be added judiciously in an “updraft” mode in which charcoal combustion supplies heat to pyrolyse the new fuel.. These other methods may require more operator slall and different design and construction. SAFETY Producer gas was the only gas fuel widely available until 1940 when natural gas pipelines became common. Since producer gas contains 10-30% CO, it is a real health hazard if the flame is extinguished or incomplete combustion occurs. (Smoky
70 1
open fires and insufficient cooking fuel are also major health hazards in the world today.) Therefore it is necessary to mandate good practice in using the Turbo Stove. In the volatile combustion mode CO is a minor hazard because if the flame should go out, the copious smoke warns the operator to re-ignite the fire or move the stove outside. However, in the charcoal combustion mode the CO is odorless and could pose a health hazard. It is recommended that all stoves including the Turbo Stove should be operated under a hood carrying the cooking odors and possible stove emissions to the outside by natural or forced convection. That is the practice in most kitchens in developing countries today and should be followed as the rest of the world develops.
MODELING THE TURBO STOVE TAYLORING THE STOVE FOR VARIOUSAPPLICATIONS For a new stove to be accepted it must fit the fuel supply, cooking practices, construction methods, size of servings and commercial infrastructure of each country. Therefore, it must be possible to make a variety of stoves and requires understanding of the basic mechanisms of gasification and combustion of “wood-gas”. For this reason it is desirable to have a complete model of the wood-gas stove from which new designs can be constructed with a minimum of testing. Optimization of stove behavior ultimately depends on reaching a maximum heat transfer to the cooking pot while minimizing emissions and soot. This is a multivariable problem and must be broken down into its component parts for solution. SUPERFICUL VELOCITY The superficial.velocity (SV) is defined as “gas production ratetcross section” and is measured in d s , btu/W-hr, etc. It is an important figure of merit of gasifiers and combustors. Each device will have a maximum SV that it can operate at satisfactorily. In downdraft gasification the SV determines the intensity of the flaming pyrolysis reaction and so controls gas, charcoal and tar production. The World War I1 gasifiers typically operated at 0.1-1.0 d s to produce low tar gas and consume most of the charcoal.’ The Turbo Stove operates in the range 0 to 0.06 gasification superficial velocity because at higher gas rates the charcoal is blown out of the top of the gasifier. The SV for three conditions are shown in Table 1 for the research gasifier operating on wood pellets. ESTABLISHING STQ VE CRITERIA To model the stove it is necessary to define the application in t e r n of maximum power required, minimum burn time at fill power and whether charcoal is desired as a by-product.
702
THE GASIFICATION SECTION OF THE TURBO STOVE The gasification section is relatively simple to model as shown in Table 4. From the burn time at full power one first calculates the fuel requirement, using 18 kJ/g for volatiles, 24kJ/g for the charcoal or 21 kJ/g for the fuel (adjusted to the moisture and ash content). The maximum power required determines the rate of production of gas, the aidfuel ratio determines the gasification air that must be supplied and the superficial velocity determines the diameter of the gasifier chamber. In Table 4 the fuel magazine diameter and height for a hypothetical 12 kW gasifier required to burn for 2 hours for community coolung are derived. We look forward to building it.
THE COMBUSTION SECTION The combustion section of the stove is less easy to model, since it depends on the mixing of the combustion air with the rising gases. Many combustion devices are rated in terms of “combustion Intensity”, which can range from lo3 to lO9Hh-m3for devices ranging from ovens to special burners6 The combustion zone in the Turbo Stove measures 10 cm diameter X 6 cm tall with a multitude of small holes for air injection. At a power level of 2.5 kW, the combustion intensity is -3 kJ/h-m3 , moderately high. We find that at power levels above 2.5 kW the flame rises above the burner and may blacken the pot due to incomplete combustion. We believe that the combustion chamber for other Turbo Stoves should have the same combustion intensity, but not necessarily the same diameter as the gasification section.
SUMMARY We have measured many important gasification and combustion properties of biomass gas made in the Turbo Stove and believe that this stove could solve many problems in world cooking. We present here a simple model for sizing other stoves.
ACKNOWLEDGMENTS We wish to acknowledge the generous support of the Biomass Energy Foundation and the Community Power Corporation for the research reported here. We are also indebted to the National Renewable Energy Laboratory and their personnel for the long term support of biomass research that has made h s work possible.
703
Table 4 - Hypothetical model of gasification section of a 12 kW* Turbo Stove designed to burn 2 hours at 12 kW* INPUT REQUIREMENTS: Maximum Power - kW* Cooking time @ Pmax - hr Charcoal (Yes/no) FUEL PROPERTIES Fuel HHV (dry) W/g Fuel Moisture Content % Ash Content - % Adjusted fuel HHV - W/g Density - kg/l GASIFIER REQUIREMENTS Fuel Rate - g / s Fuel consumed in time-g AirEuel Ratio Gas Produced - g/s Molecular Wt - M(g) Gas Produced - N d / s Maximum SV - m/s Gasifier area - m2 Gasifier diameter - cm Fuel volume - cm3 Fuel magazine height-cm
12 2 no
Source Assumed Assumed Assumed
Coconut Shells Assumed 21.0 Typical biomass, ash free, dry 6.0% Denver dry 0.7% Measured 19.6 Calculated 0.48 Measured 0.61 4408 1.60 1.59 25 0.00143 0.06 0.0238 17.4 9183 38.6
W/s+kJ/g Run time X fuel rate ( 1+A/F)XFuel rate Assumed (22.4 Vmole)XM(g)XlO”/M Measured m3/s-m/s = m2 D = (4*area/pi)”’ Weighddensity Volume/area
REFERENCES 1. Kammen, D. M., Cookstoves for the Developing World, Scientific American, 2.
3.
4.
5. 6.
July, 1995, p.72 La Fontaine, H. and Reed, T. B., An Inverted Downdraft Wood-Gas Stove and Charcoal Producer, in Energy from Biomass and Wastes XV, D. Klass, Ed., Washington, D. C., 1993. Reed, T. B. and Larson, R., A wood-Gas Stove for Developing Countries, in Developments in Thermochemical Biomass Conversion, Ed. A. V. Bridgwater, Blackie Academic Press, 1996. Reed, T. B. and R. Walt, The “Turbo Wood-Gas Stove, in Biomass: Proceedings of the 4* biomass Conference of the Americas in Oakland, Ca, Ed R. P. Overend and E. Chornet, Pergamon Press, 1999. Reed, T.B., Walt, R. Ellis, S., Das, a. and Deutch, S, Superficial Velocity - the Key to Downdraft Gasification, ibid. Combustion Handbook, Vol 1, North American Mfg. Co., Cleveland, OH, Third Edition, 1986, p. 13. ”
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Effect of GR GRANULE used as Bed Material to Reduce Agglomeration in BFB Combustion of Biomass with High Alkali Metal Content J. H. Daavitsainen, R. S . Laitinen, L. H. Nuutinen, H. J. Ollila, M. S. Tiainen, and M. E. Virtanen. Department of Chemistry P. 0. Box 3000 FIN-90014 University of Oulu, Finland
ABSTRACT The combustion of fuel containing large amounts of alkali metals has proven to be very problematic due to the bed agglomeration. The h g h alkali metal content in the fuel lowers the melting point of silicates and aluminosilicates that are normally formed in a silica sand bed. The best conventional way to utilize fuel with a high alkali metal content is to co-combust it with other fuel types. It is also possible to avoid the agglomeration and defluidization by using additives like kaolin in order to bind alkali metal in the ash. In this paper we describe a new bed material that does not react with alkali metals in a harmful manner. Numerous boiler tests have been carried out in different boiler plants co-combusting plywood-waste with bark or sawdust without any additives. We also describe a test in which only plywood-waste was used as fuel. The main source of alkali metals in these tests was the adhesive in plywood. No agglomeration problems related to the reactions of ash with the bed material were observed. In fact, these boilers and the fuels have been fully utilized ever since. During the tests, fuel and bed material samples were collected. These samples were mainly investigated with SEM-EDS. Bulk elemental analyses were also carried out. The unproblematic behaviour of the bed material was confirmed and no agglomerate formation was observed except in the close proximity of occasional quartz centres that were unintentionally present in bed. It was observed that the majority of the bed particles were coated with several thin superimposed layers. The composition of each layer was different. The outermost coating layer contains a significant amount of magnesium. It may be this outermost layer that hmders the agglomeration of the particles in tlus new bed.
705
INTRODUCTION
The utilisation of the fluidised bed technique has become more common in combustion because it allows the combustion of low heat capacity fuels. The environmental benefits associated with biomass fuel (e.g. forest residue, bark, sawdust and wood chip) and waste material (e.g. RDF, industrial based waste) has increased their use in energy production during the last decade. These fuels are often problematic and cause problems like fouling and agglomeration during combustion. The tendency towards agglomeration also depends on bed material, additives, and on the boiler conditions. The combustion of fuels with high alkali metal content are known to show high propensity towards agglomeration when using silica sand and quartz bed, which may lead to the defluidisation of the bed.' Problems are often avoided by co-combusting high alkali metal content fuel with other fuel types, or by using additives (kaolin and dolomite) during combustion.'^ lo In this paper we discuss the results from numerous heating plant tests utilising BFB-combustion of biomass with high alkali metal content (plywood waste). The plywood ash contains 33% ofNa20."
'-*
EXPEFUMENTAL
The heating plant tests were camed out in four different BFB-boilers utilising quartz-free bed material GR GRANULE. The fuel mixtures and the capacity of boilers are described in Table 1. Tests 1 and 2 were full-scale test m, whereas the samples taken from Tests 3 and 4 were collected during the normal operation of the heating plants.
Table 1 The fuel mixtures used in heating plant tests and the capacities of the boilers. Boiler test Test 1 Test 2 Test 3 Test 4
The fuel mixture
Plywood-waste and sawdust (spruce and pine) in different ratios Plywood-waste Plywood-waste, dust and bark (spruce, pine and birch) Plywood-waste and bark (spruce and birch)
The capacity of the boiler
6Mw 5Mw 25 MW 15 MW
We explored unused bed material and plywood, as well as samples from the bed that were collected from the boiler during the tests. The bed samples were obtained from bottom ash chute after the water-cooled screw. The ashing of plywood was carried out according to ASTM standard (D 317489).12 The sample preparation for the ICP- and DCP-AES determinations were carried out according to ASTM standard (D 3682-97)13 by mixing the bed samples or
706
fly ash with Li2B407.The determinations were performed with a Philips PU- 7000 ICP-AES or a Spectraspan IIIB DCP-AES spectrometers. Bed samples were characterised by using a Jeol JSM-6400 scanning electron microscope combined with a Link ISIS energy dispersive X-ray analyser. The acceleration voltage of 15 kV and a beam current of 120*10-'A were used for the SEM-EDS-analysis. The beam current during the recording of the X-ray maps was 670*10-' A. The sample distance was 15 mm. The magnification used for bed samples was x50. All SEM-samples were mounted with resin, cross-sectioned, polished, and coated with a thm carbon layer. About 1000 domains were analysed for sodium, magnesium, aluminium, silicon, sulfur, phosphorus, potassium, calcium, titanium, and iron contents by utilising ca. seven different image fields.I4 The image analysis was performed with IMQuant software incorporated in Llnk ISIS. The SEMEDS-results were visualised by use of quasiternary diagrams using a Konpad software package that is locally designed especially for this purpose.I5
RESULTS AND DISCUSSION The SEM-EDS-results of the used bed samples indicate only little agglomeration in the bed. The small agglomerates were formed solely around quartz particles that were present in the bed as an impurity from fuel, or as a residue of the previous bed. The quartz particles had probably reacted with sodium from the fuel forming sodium aluminium silicates of a low melting point. The coating thus formed may eventually lead to agglomeration.I6 Although most bed particles did not agglomerate, they were coated with several thin coating layers. During the initial stages of the tests, the layers were very thin,but they gradually grew as the test proceeded. The innermost coating layer on majority of the particles were first coated mostly with calcium and phosphorus with the outermost layer being mamly magnesium-rich. It is probable that the formation of these layers protects the bed particles from agglomeration.". These coating layers are discussed in detail elsewhere.I6 The bed material in the heating plant during its normal operation (Test 3) was sieved once a week in order to remove small agglomerates formed during the combustion. The bed material containing agglomerates was disposed, and the sieved bed material was returned to the boiler. No new bed material was added during the 9month of continuous operation. The size of the bed particles increased during the observation period. The compositional distribution of the unused bed material is shown in Fig. 1 (a). The quasiternary diagram of the bed sample from Test 3 containing agglomerates is presented in Fig. 1 (b). The compositional distribution of the disposed bed material is shifted towards the Al+Si/Na+K+P+S-edge. The adhesive material of the agglomerates can be seen in the diagram as a lighter grey area. When comparing the unused bed material [Fig. 1 (a)] with that after the operation of two months [Fig. 1 (c)], it can be seen that the compositional distribution is shifted towards the Mg+Ca+Ti+Fe-corner of the diagram.
707
a+Ti+Fe YO
a+Ti+Fe %
Na+P+S+K YO
Na+P+S+K %
+Ca+Ti+Fe YO
Na+P+S+K %
Fig.1 Series of quasiternary diagrams of bed samples from Test 3. (a) unused bed material, (b) bed material containing small agglomerates (disposed), and (c) sieved bed material (returned to boiler) A quasiternary diagram of coating layers from Test 4 is shown in Fig. 2 (a). The sharp maximum in the Mg+Ca+Ti+Fe-corner is due to a magnesium-rich layer and the broader maximum is due to calcium- and phosphorus-rich layers. When the quasiternary diagrams [Fig. 2 (b-d)] is compared to that of the unused bed material [Fig. 1 (a)], it can be seen that a second maxima is formed towards the Mg+Ca+Ti+Fe-corner from the that of the unused bed material indicating the growth of the calcium and phosphorus-rich coating layers. It can be seen that the compositional maximum due to unreacted bed material gradually disappears.
708
Na+P+S+K %
Na+P+S+K %
a+Ti+Fe YO
a+Ti+Fe %
Na+P+S+K YO
Na+P+S+K %
Fig. 2 Quasiternary diagrams of point analysis of coatings (a) and the bed samples taken on test days (b) 10, (c) 15, and (d) 21 from test 4. To investigate the effect of dolomite, few samples of silica sand bed with dolomite addition were studied. Interestingly, the similar coating layers could be found in the silica sand bed to those found in a quartz-free bed material. In X-ray maps (Fig. 3) these coating layers can easily be seen. All bed particles have an inner coating layer containing mainly calcium and phosphorus. A discontinuous outer layer of magnesium can also be seen in Fig. 3.
709
Fig. 3 (a) The BE-image (x50) of coated silica sand bed particles, (b) X-ray maps indicating major elemental composition of the particles.
The surface area based mean values of the SEM-EDS-results of the bed material and the ICP-AES determinations of the bed samples from Test 4 are shown in Fig. 4. It should be noted that the bulk information given by ICP-AES determinations agree well with the average values calculated from the compositional distribution based on SEM-EDS determinations.
CONCLUSIONS The bed material presented in this work was found suitable for combustion of plywood waste that is known to be problematic because of its high sodium content. Only little agglomeration was observed during the tests. These agglomerates were solely formed around quartz particles that were present as impurities in the bed. The coating of the bed particles exhibited several thin layers. The majority of the particles were first coated with calcium and phosphorus with the outermost layer being magnesium-rich. It is probable that the formation of these layers protects bed particles from agglomeration. Further evidence for this conclusion is found from a test where dolomite was added to silica sand bed to prevent agglomeration. A similar layered coating was found in the silica sand bed particles. The beneficial properties of the quartz-free bed that was initially observed in two full-scale tests"' l7 were verified during the nine months of normal operation a heating plant.
710
60
50
Day 10 SEM Day10 ICP
40
Day 16 SEM Day 16 ICP
30
D Day 21 SEM
20
10
1
0 Na
Mg
Al
Si
P
S
K
Ca
Ti
Fe
Fig. 4 The results from SEM-EDS and ICP-AES analyses of bed samples. The results are normalized to 100 % by the analysed elements.
ACKNOWLEGDEMENTS Financial support from Technology Development Centre Finland, Putkimaa Oy and Vapo Oy, is gratefully acknowledged.
REFERENCES 1 Natarajan, E., Ohman, M., Gabra, M., Nordin, A. Liliedahl, T., and Rao, A.N., “Experimental determination of bed agglomeration tendencies of some common agricultural residues in fluidised bed combustion and gasification”, Biomass and Bioenergy 15 (1998) 163-169. 2 Salour, P., Jenkins, B.M., Varaei, M., and Kayhanian, M., “Control of bed agglomeration by fuel blending in a pilot scale straw and wood fuel AFBC”, Biomass andBioenergy 4 (1993) 117-133. 3 Skrifvars, B-J., Backman, R., and Hupa., M., “Characterization of the sintering tendency of ten biomass ashes in FBC conditions by a laboratory test and by phase equilibrium calculations”, Fuel Processing Technology 56( 1998)55-67. 4 Mann, M.D., Swanson, M.L., and Yagla, S.L., “Characterization of alkali and sulfur sorbents for pressurized fluidised-bed combustion”, Proceedings of the I 3Ih of International conference on fluidised bed combusition, Ed. Heinschel, K.J. (1995) pp.333-340. 5 Lind, T., Kauppinen, E.I., Jokimemi, J.K., Maenhaut, W., and Pakkanen, T.A., “Alkali metal behaviour in atmospheric circulation fluidised bed coal combustion”, The Impact of Ash Deposits on Coal Fired Plants, Ed. Williamson, J. & Wigley, F., (1993) Sollhull, Birmingham, UK, pp. 77-88. 71 1
6 Mansaray, K.G., and Ghaly, A.E., “Agglomeration characteristics of alumina sandrice husk ash mixtures at elevated temperatures”, Energy Sources 19 (1997) 10051025. 7 Ergiidenler A. and Ghaly A. E., “Quality of gas produced from wheat straw in dual-distributor type fluidised bed gasifier”, Biomass and Bioenergy, 3 (1992) 419430. 8 Ergudenler A. and Ghaly A. E., “Agglomeration of silica sand in a fluidised bed gasifier operating on wheat straw”, Biomass and Bioenergy, 4( 1993) 135-147. 9. Steenari B-M. and Lindqvist 0. (1998) High-temperature reactions of straw ash and the anti-sintering additives kaolin and dolomite, Biomass and Bioenergy, 14, 67-76. 10. Kallner P. & Zintl F. (1997) Orsaker till pislag p i varmeoverforingsytor vid tradbransleeldning och additiv fir att minska problemen. Varmeforsk. 11. Nuutinen L. H., Ollila H. J., Tiainen M. S., Virtanen M. E., and Laitinen R. S., An Improved bed material for the BFB-boilers. Case 1: Co-combustion of sawdust and plywood waste. 5” International Conference on industrial furnaces and boilers, April 11-14.4.2000 Porto, Portugal. 12. Standard Test Method for Ash in the analysis sample of coal and coke from coal, D 3174-89. 1989 Annual book of ASTM standards, Vol. 05.05, Gaseous Fuels; Coal and Coke, pp. 302-304. 13. Standard Test Method for Major and Minor Elements in Coal and Coke Ash by Atomic Absorption, D 3682-97. 1998 Annual book of ASTM standards, Vol. 05.05, Gaseous Fuels; Coal and Coke, pp. 345-350. 14. Virtanen M. E., Skrifiars B. J., HeikkinenR. E. A., HupaM., PatrikainenT., and Laitinen R. S. (1999) A novel approach to use CCSEM when studing agglomeration in fluidized bed combustion. In: The impact of mineral matter in solid fuel combustion, (Ed by R. P. Gupta, T. F. Wall & L. L. Baxter), pp. 147-154. Plenum Press, New York. 15. Virtanen M. E., Tiainen M. S., and Laitinen R. S., SEM-EDS image analysis in the characterisation of coatings and adhesive material in quartz bed. 5“ International Conference on industrial furnaces and boilers, April 11-14.4.2000 Porto, Portugal. 16. Nuutinen L. H., Ollila H. J., Tiainen M. S., Virtanen M. E., and Laitinen R. S., Role of quartz sand in the agglomeration during the FB-combustion using fuel of high sodium content. Ash deposition: Problems, Management & Solution, May 8.11.5.2000Park City, Utah, United States. 17. Laitinen R. S., Nuutinen L. H., Tiainen M. S., and Virtanen M. E., An Improved bed material for the BFB-boilers. Case 2: Combustion of fuel with high sodium content. 5” International Conference on industrial furnaces and boilers, April 1114.4.2000Porto, Portugal.
712
Elemental Gas-Particle Partitioning in Fluidized Bed Combustion and Gasification of a Biomass Fuel D. Pfaff', B.M. Jenkins' and S.Q. Turn' 'Department of Biological and Agricultural Engineering, University of California, One Shields Avenue, Davis, CA 95616-5294 2Hawaii Natural Energy Institute, University of Hawaii, Honolulu, HI, 96822
ABSTRACT The fates of alkali and other elements during combustion and gasification are important to the process equipment used in these biomass conversion applications. Partitioning of alkali metals and other elements to the various product phases impacts the formation of fouling deposits on heat exchangers, slagging, corrosion of equipment, and deactivation of catalysts. The partitioning is also important to the control or removal of alkali species in gas turbines and other advanced generation technologies. Alkali metal (K, Na) partitioning during combustion and gasification of a blended wood-almond shell fbel was investigated in a bench-scale atmospheric fluidized bed reactor. The same fuel was used in both combustion and gasification regimes. Product streams were characterized in terms of gas and solid phase compositions and flow rates. Deposit samples from the disengagement zone of the reactor and from the downstream convection pass were collected and analyzed. Particle and gas phase samples were also collected via extractive probes directly from the disengagement zone of the reactor and downstream following cyclone separation. Element balances for all major components were computed. Combustion deposits on in-bed equipment were composed principally of potassium chlorides and sulfates. Gas phase potassium concentrations ranged from 0.6 to 0.9 ppmw during combustion (60% excess air) and up to 2.8 ppmw during gasification (0.17 air factor). Gas phase chlorine concentrations ranged up to half stoichiometric yield. Gasification resulted in a reduction of total potassium volatilization relative to combustion. 1 INTRODUCTION
Many biomass fuels contain high concentrations of alkali and alkaline earth metals (K, Na, Ca, Mg)'. These elements, especially in combination with chlorine, sulfbr, and silica contribute to ash slagging, the formation of fouling deposits on heat exchangers
713
used in boilers and other combustion equipment, corrosion, and deactivation of catalysts in selective catalytic reduction (SCR)systems used for NO, emission control on power ~lants.2.~Alkali metal concentrations must be carefully controlled in combustion gases entering gas turbines, a critical issue in the application of gasifiers in biomass integrated gasifier combined cycles (IGCC) and other gas turbine based power generating concepts. Recommended vapor phase alkali concentrations are frequently below 0.1 ppm in the gas ahead of the turbine. Reciprocating engines are more tolerant of contaminants in the fuel gas, but also require gas cleaning, and system efficiencies are generally lower than turbines for combined cycle applications. In fluidized bed reactors, alkali metals and other elements contribute to bed media agglomeration, leading to enhanced channeling and bed defluidization:" The detrimental effects of inorganic constituents on power systems cause increased costs for maintenance, decreased plant efficiency, reduced plant availability, and reductions in energy revenues leading to overall poorer plant economic perf~rmance.~Fuel selection, i.e., selecting higher quality fuels with lower tendencies to foul and slag, has been a common technique in the biomass power industry to mitigate these impacts. This technique excludes the use of many poorer quality fuels that can be available at lower cost. Gasifiers have been proposed as a means of reducing alkali impacts on advanced power systems by reacting the solid fuel at lower temperatures with the intent of reducing alkali volatilization. A number of studies have been conducted examining alkali release during gasification and combustion of biomass, under highly controlled laboratory conditions as well as in pilot- and full-scale Exiting the reactor, alkali compounds occur in the vapor phase and in or on particulate matter. The principal alkali compounds released from the fuel are KCI and NaCl if chlorine is present in the fuel.'-" The temperature for condensing vapor phase alkali is typically between 400-550°C.'s Condensation of alkali metals on particles, with subsequent removal of particles is one means of reducing alkali concentrations to acceptable levels for gas turbines. The research described here compares gas-particle phase alkali and other element partitioning at different locations in an atmospheric bench scale fluidized-bed reactor operated under combustion and gasification regimes with the same fuel.
2 METHODS AND MATERIALS 2.I FUEL
All tests were performed using the same biomass fuel composed of 80 wt.% wood blended with 20 wt.% almond shells. Wood fuel was obtained from the fuel feed conveyors of an operating biomass power plant and was largely a mixture of Douglas fir and Ponderosa pine. The blended fuel material was milled and sieved through a 16 mesh (1 mm) screen. Before each test, the moisture content was determined by drying the fuel for 24 hours in an air oven at 103OC. The moisture content ranged from 9.1 to 9.4 wt.% wet basis. Properties of the fuel are listed in Table 1.
714
2.2 REACTOR SYSTEM
Each test was carried out in an atmospheric bench-scale fluidized bed reactor shown schematically in Figure 1. The main reactor consists of a 321 stainless steel tube with an inside diameter of 73 mm and a length of 1 m. The tube is encased in an electric furnace used to preheat the reactor. At the top, the reactor expands into a 127 mm square section with a cross-sectional area 4 times larger than the main reactor column. This section is used to disengage larger bed and partially reacted fuel particles from Table I Fuel properties. Higher heating value (constant volume) (MJkg, dry basis) Moisture content (wt.% wet basis) Proximate analysis (wt.% dry fuel) Volatile Matter Fixed Carbon Ash Total
Ultimate analysis (wt.% dry fuel) C H 0 (by difference)
19.65 9.1 - 9.4 71.4 24.3 4.3 100.00
49.1 1 5.76 39.93 0.87 0.03 0.03 4.27 100.00
N S
c1 Ash Total
Elemental composition of ash (wt.% of ash) 32.69 11.77 0.53 6.52 15.16 3.83 1.36 13.56 2.64 0.99 0.03 9.39 1.53 100.00
Undetermined Total
the flow for internal recirculation to the bed. Exiting the side of the disengagement section, the flow passes through a horizontal deposit test section and cyclone.
715
Removable U-shaped deposit probes are located in the disengagement sectiori and the horizontal pass. The deposit probes are air-cooled with surface temperatures measured via type K thermocouples welded into the tube wall. Probe cooling air is exhausted externally and is entirely separate from the reactor gas flow. Some ash and elutriated bed media particles are removed by settling in the horizontal pass, and the cyclone located downstream of the horizontal pass provides additional particle separation. Gas and small particles not separated by the cyclone are discharged into a vertical ceramic-lined exhaust duct drawing laboratory air for dilution and cooling. When the reactor is functioning as a gasifier, the product gas is flared in the duct at the cyclone stack. Primary fluidizing air is distributed at the bottom of the reactor through a multi-hole nozzle. A series of electric heaters upstream of the distributor nozzles preheat the air. For combustion, secondary air is added for more complete combustion of the fuel. Secondary air can be injected anywhere along the reactor, but for all combustion tests reported here secondary air was injected only at the center of the main reactor column. Secondary air was not preheated. Fuel was loaded onto a flat belt feeder metering onto a high speed auger injecting into the lower bed. Fuel feed rate was controlled by varying the feeder belt speed and the height of the fuel bed on the belt. Average fuel feed rate was determined from the total fuel consumption and duration of each test. Instantaneous feed rate was estimated from the belt speed. Thermocouples and pressure taps were situated at various heights along the reactor column, and at locations downstream. Thermocouples were also located in the deposit probe walls, and at the cooling air inlet and outlet of each deposit probe. Temperatures and pressures were recorded by electronic datalogger every 10 s. Prior to each test, the reactor was preheated to a temperature of 800°C and primary air was preheated to 350 400°C. In combustion, secondary air injection was started immediately after initiating fuel feed to the reactor. Temperature equilibrium and uniform temperature distribution were achieved typically within 15 min from start. Sampling was conducted only after reaching steady operating conditions.
-
Each test was terminated by stopping fuel feed and air flow. After allowing the reactor to cool, remaining bed material was removed, ash was collected, and heat exchanger deposits were recovered. All samples were weighed and sealed pending further analysis.
716
4 Exhaust
Horizontal Pass IPA npnncit -".. Air Coo.,, Probes
Access PortslO Bed M*eup" 2i Air
f
127 x 127 mm disengagement section
Air Cooled Deposit Probes Optical Access Ash Dropout
\
I
E
Electric Furnace
Ash Dropout
Distributor Nozzle
d ThermocouplePressure Taps Bed Drain
Fig. 1 Schematic of the bench-scale fluidized-bed reactor system.
717
2.3 BEDMEDIA For all tests a refractory alumino-silicate grain was used for the bed media. Due to higher air flow rates and fluidizing velocities, a larger particle size (Investocast 35, North American Refractory Company, Ione, CA) was used during combustion than was used during gasification (Investocast 60). Average grain size of the Investocast 35 is 0.42 mm, and 0.21 mm for the Investocast 60. Fresh bed media was weighed into the reactor at the start of each test. The bed material was recovered after each test, weighed and sampled for analysis. Samples of fresh and spent bed material from combustion and gasification tests were analyzed for Si, Al, Ti, Fe, Ca, Mg, Na, K, P, S and C1 by X-Ray fluorescence (Hazen Research, Inc., Golden, Colorado). 2.4 ALKALI SAMPLING Gas and particle samples were withdrawn from either two or three locations depending on the mode of operation. During combustion and gasification tests, samples were collected from the disengagement section of the reactor and from the cyclone stack (Fig. 1) prior to dilution. During gasification, samples were also collected post-flare in the exhaust duct (Fig. 1). Samples were extracted non-isokinetically through a 7.8 mm id. stainless steel tube. The sampled stream passed immediately through a heated 2 pm sintered stainless steel filter with outside diameter of 12.7 m and an effective length of 152 mm. The filter was housed inside an electric heater, with the filter held at or somewhat above the temperature of the gas at the extraction point. Particle separation was thereby accomplished without cooling of the gas. For sampling in the disengagment zone during gasification, a size selective inlet was located inside the reactor to separate coarse particles and increase sampling time for the sintered filter. The inlet consisted of a stainless steel cyclone with both top and bottom outlets ducted outside the reactor wall. The coarse particle fraction was collected from the bottom outlet, while fine particles were collected on the heated sintered filter connected to the top outlet. The fine particle filter was again maintained at a temperature equal to or above the gas temperature at the extraction point. Coarse and fine particles were thus separated without cooling. After passing the fine filter, the sample stream was cooled in a vertical, 0.6 m long, 10.9 mm i.d. stainless steel water jacketed condensor draining directly into the first of four ice bath cooled impingers. Prior to each test, the first two impingers were loaded with 200 mL of distilled water, the third impinger with 100 mL distilled water, and the fourth impinger left dry. Following the impingers, the gas stream passed through a weighed silica gel desiccant, sample pump, dry test meter for total gas volume, and rotameter indicating instantaneous sample gas flow rate. Following each test, the sample line ahead of the impinger train was rinsed with distilled water and the rinsate collected. Impinger liquids were recovered, liquid volumes determined, and the impingers were rinsed with distilled water. Sample line and impinger rinsate were added to the first impinger liquid, the other impinger liquids were collected separately to check for breakthrough. All lines and impingers were acetone washed prior to each subsequent test. All liquids were submitted for analysis of C, Al, Ca, C1, Fe, Mg, P, K, Si, Na, S and N (Hazen Research, Inc., Golden,
718
Colorado). Desiccant weight gain was combined with impinger liquid volume increase to determine water content of the sample gas. Filter cake was recovered dry from the fine particle filter. The filter including housing and sample line was rinsed with distilled water to recover any residual sample. After soaking 24 hours to dissolve alkali, the rinse was filtered through an ash-free filter (Whatman 42) on a Buchner funnel to remove insoluble solids. The filter dry weight before and after filtration was recorded to determine soluble and insoluble material mass fractions. The filtrate was analyzed for C, Al, Ca, C1, Fe, Mg, P, K, Si, Na, S and N (Hazen Research, Inc., Golden, Colorado). The collected filter cake was weighed and submitted for ultimate analysis (gasification mode only due to the high residual organic fraction) and X-ray fluorescence analysis giving inorganic composition (Hazen Research, Inc., Golden, Colorado). Coarse particles were similarly analyzed. 2.5 GAS ANALYSIS
During combustion, and post-flare during gasification, CO, C02 and NO, (NO, NO2) were monitored via continuous gas analyzer (Horiba Instruments, Irvine, CA). During combustion, the gas sample was extracted from the cyclone stack prior to dilution in the exhaust duct. In gasification mode, continuous monitoring was conducted only from the post-flare sampling location in the exhaust duct (Fig. 1). The gas sample stream was cooled through two icefingers in series to condense liquids and then passed through a fiberglass plug and a molecular sieve (Linde 3A, Union Carbide, South Plainfield, NJ) to separate coarse particles and desiccate the sample. The sample was further cooled through a refrigerated coil and any remaining particles removed through sintered stainless and glass fiber final filters. NO, was analyzed via chemiluminescence, CO and C02 were analyzed via NDIR. Gas concentrations were recorded by electronic datalogger every 10 s. Filters and desiccants were replaced and all analyzers were calibrated using standard gases prior to each test.
Gas grab samples were also collected in glass sample flasks during combustion and gasification tests, and analyzed via gas chromatography for COZ,CO, N2, 02,CH4 and H2. Grab samples were collected from each sampling location during steady reactor operation.
2.6 DEPOSIT PROBES Two sets of removable stainless steel U-tube heat exchangers were utilized as deposit probes. At the conclusion of each test, deposits were collected dry by brushing and weighed. As brushing did not remove all deposited material, the surface was also rinsed with distilled water. The rinse was allowed to stand for 24 hours to dissolve solids, then filtered through ash-free filter paper (Whatman 42) to separate insoluble materials following the same procedure used for the fine particle filter of the alkali sampling system. Dry filter weight gain was recorded. Filtrate was analyzed for C, Al, Ca, C1, Fe, Mg, P, K, Si, Na, S and N (Hazen Research, Inc., Golden, Colorado). In gasification mode, the dry fraction of the deposit was analyzed for ultimate elemental composition (C, H, 0, N, S, and Cl) due to the large residual fraction of organic materials, and for elemental ash composition (Si, Al, Ti, Fe, Ca, Mg, Na, K, P, S, Cl). Probes were rinsed and rinsate processed and analyzed as described for the combustion tests.
719
3 RESULTS 3.1 OPERATING CONDITIONS
In combustion mode, duplicate tests were carried out with a dry fuel feed rate of 0.33 g s-' and an air factor of 1.6 (ratio of actual air-fuel ratio to the stoichiometric air-fuel ratio). Two tests were required to obtain alkali samples from both the disengagement section of the reactor and the cyclone stack. Four gasification tests were conducted, three to obtain alkali samples from the disengagement section, cyclone stack, and the post-flare flow, and a fourth test again sampling from the disengagement zone but using a 5 pm sintered filter to extend the sampling time at this location. However, both filters loaded rapidly in gasification mode at the two sampling locations ahead of the flare. Dry fuel feed rate in gasification mode averaged 1.43 g s-' with an air factor of 0.17. Average reactor temperature in combustion mode was 870°C and in gasificaton mode 750°C. Test conditions for both combustion and gasification tests are summarized in Table 2. 3.2 DEPOSITS Deposits were collected from the disengagment section and from the horizontal pass. A summary of properties for deposits removed from both locations appears in Table 3. Surface temperatures were controlled to about 500°C for probes in the disengagement zone. Specific deposition rates in the disengagment section are higher for the combustion tests than the gasification tests due to the higher fuel feed rate for gasification with the same total deposit mass. Compositions of the deposits differed as a function of the testing regime (combustion or gasification). Deposits from gasification have large carbon concentrations (30 to 50 wt.% as shown in Table 4) due to the large residual organic fraction under reducing conditions. Carbon in the combustion deposits is largely present as carbonates, at much lower concentrations. The combustion deposits are composed principally of potassium sulfates and chlorides. The gasification deposits contain more silica and alumina due to attachment of bed media particles that occurs as a result of the higher organic fraction. Heat exchangers in commercial gasifiers would not normally be positioned at the locations of the probes here. The deposit probes were principally intended for combustion experiments, but were left in place during gasification to compare with the combustion results.
720
*
(“(3
I
re 2 Summary 0 perating condii ins and results. Combusl n Mode ication Mode Alkali sampling location: Disengagement Cyclone Disengagement Disengagement Stack Zone Zone’ Zone’ Total wet fuel burned ( g ) 2,652 2,652 4,627 12,261 3,25 1 4,903 Fuel moisture content (wt. % wet basis) 9.4 9.1 9.4 9.4 9.2 9.2 Wet fuel feed rate (g s-I) 0.37 1.58 0.37 1.58 1.58 1.58 Reactor preheat temperature (“C) 800 800 800 800 800 800 Alkali filter temperature 750 350 250 730 730 350 Primary air preheat temperature (“C) 400 400 350 350 350 350 Fresh bed material (g) 433 433 433 433 433 433 Removed bed material (g) 474 414 389 419 449 549 Primary air flow rate (L mid’) 60 140 140 60 60 60 --Secondary air flow rate (L mid’) 15 15 Air factor 1.6 0.17 0.17 0.17 1.6 0.17 Alkali filter cake (g) 0.74 0.16 1.34 1.54 1.63 0.01 Total collected ash (g) 116 403 247 83 432 1,335 Condensed liquid in impingers ( g ) 20.1 29.6 42.2 23.0 24.4 6.8 Desiccant weight gain (g) 4.9 4.2 12.1 8.1 4.3 4.2 Total sampling time (min) 29.5 60 90 20.0 19.6 123 Total sampled dry gas volume (L)4 161 284 448 100 99 654 Gas flow rate through alkali filter (L mi^^-')^ 5.1 5.0 5.5 5.3 4.7 5.0 ’2 pm fine particle filter. ’5 pm fine particle filter. ’average mperature inside er during sampling. 4volumeat 20°C and atmospheric pressure.
Table 3 Summary of deposit probe conditions and results. Combustion Gasification Dis. Horiz. Dis. Horiz. zone' pass' zone' pass' 496 426 507 361 Probe surface temperature ("C) 0.4 0.1 0.4 0.7 Total deposit mass (g) 8.2 1.5 20.5 42.7 Specific deposition rate (g m-*h') 156 28 90 189 Specific deposition rate (mg deposivkg fuel fed) 3.8 0.7 0.9 2.0 Specific deposition rate (mg deposivg ash fed in fuel) 'Disengagement section probes. 'Horizontal pass probes.
Table 4 Deposit compositions (wt. YO). Combustion Gasification Dis. Horiz. Dis. Horiz. zone' pass' zone' pass' C
0.6
nd nd
nd nd
17.0 11.7 0.9 0.9
13.3 12.7 2.2 2.2
nd <0.1 5.3 0.3 1.5 61.5 0.1 100.0
nd
34.5 0.5 0.5 0.5 0.6 20.3 2.9 1.2 6.2 15.2 11.1
2.2 4.3
54.3 0.8
0.1 0.7 0.2 14.4 1.4 1.o 7.8 11.3 6.4 1.2 0.4
Gas compositions as determined by GC are listed in Table 5 for the two combustion sampling locations and the three gasification sampling locations. Carbon dioxide was also measured on-line by continuous gas analyzer (NDIR method) during the combustion tests and averaged 10.8 vol.% at the same conditions. Average NO, concentration (as NO') in combustion was 165 ppm (volume) and average CO concentration was 1,252 ppmv. Post-flare gas composition from the gasification tests reflects the air dilution in the exhaust stack at the point of sampling. The air dilution ratio varied from about 10 to 30. Average CO and NO, concentrations measured by continuous gas analysis were 349 ppmv and 33 ppmv, respectively. The volumetric stoichiometric air-fuel ratio for the producer gas is about 1.15.
722
Nz 0'
co2 co
Combustion Cyclone Dis. Stack Zone 78.5 80.8 9.7 7.6 10.8 11.6
0.1
__
Cyclone Stack 48.9
--
<0.05
Gasification Dis. Zone 46.3
Post flare 79.2 19.1
--
16.1
16.2
1.7
20.0 10.5 4.5
22.1 10.2 5.2
<0.05
-_ H2 -CH4 Heating Value' --_ 5.12 5.6' (MJm-') 'higher heating value, 300K. 2calculated,excludes C2and higher hydrocahons.
__
I
__
-_
__
3.4 ALKALI FILTER SOLIDS COMPOSITIONS
For combustion, the filter cake on the alkali filter was composed principally of ash, while for gasification the filter cake contained ash, char, and condensed tars. Particle concentrations in gasification were substantially higher than in combustion. Compositions of the dry solids from the filter are listed in Table 6. The mass of dry solids recovered (10 mg) from the filter at the post-flare sampling location was too small to be analyzed directly and thus no composition appears in the table.
3.5 ELEMENT BALANCES 3.5.1 Combustion
Element mass balances were determined using both 1) the direct air flow (rotameters) and fuel feed rate measurements, and 2) by closing the carbon balance. The extent to which the element balances are closed by the two methods appears in Table 7 (closure on the carbon balance for the second technique is necessarily 100%). Excess air was 57% by the direct method and 86% via carbon balance using CO and C02 concentrations measured by the continuous gas analyzer. Small amounts of ash were found in the bed material and most ash was deposited in the ash dropouts below the horizontal pass and the cyclone. A small fraction of carbon was found in ash and spent bed material. Values in Table 7, with the exception of the closure values in the last column, are percent of element or compound in each material fraction relative to input in fuel and air. In the case of bed material, the values are computed as the ratio of the difference in constituent mass between spent bed and fresh bed to the input in fuel and air. Negative values associated with the bed material arise from carryover of bed particles from the reactor with a corresponding decrease in total spent bed mass relative to fresh bed mass. Most of the carryover bed is found in the ash collected at the horizontal pass and cyclone, although some fine particles originating from bed media pass the cyclone and are associated with the particle phase of the stack flow. The closure values in the last column of the table are overall values based on the ratio of total component mass in products to total component mass in inputs, including fresh bed media, fuel, and air.
723
Combustion Mode Dis. Cyclone Zone' Stack C H 0 (diff.) N S
c1
Ash Total Ash composition (wt. % ash
Dis. Zone2 natter)
Gasification Mode Dis. Dis. zone3 zone4
Cyclone Stack
63.21 0.65 2.94
63.56 0.38 2.14
62.80 0.40 0.54
68.08 1.70 2.80
0.53 0.03 0.10 32.54 100.00
0.5 1 0.15 0.07 33.19 100.00
0.78 0.02 0.04 35.42 100.00
0.98 0.08 0.24 26.12 100.00
22.99 24.71 22.47 24.59 25.02 1.09 1.74 2.98 2.20 3.39 1.66 1.57 1.66 1.79 1.04 12.38 14.20 12.26 9.80 10.31 30.75 29.26 29.77 33.64 25.25 4.83 4.86 4.52 3.47 5.70 0.91 0.89 0.91 0.40 1.11 1.11 20.12 17.09 17.60 24.39 10.66 15.35 3.40 3.35 3.34 2.77 3.48 3.59 0.3 1 0.11 0.16 0.11 1.08 1.10 0.91 0.31 0.21 0.07 0.60 1.21 4.18 1.99 0.74 2.71 O5 O~ Und. 05 4.94 Total 100 100 100 100 100 100 hction from size selective inlet. 'Disengagemnt zone. '2 pm fill ' 5 Drn filter. 4~oarse SCompositionnormalized to reduce toil to 100%. 26.83 5.28 1.25 11.33 28.52 5.70
The hydrogen and nitrogen balances support the higher air flow derived from the carbon balance, closing at 102 and 100% respectively by the indirect method. Sulfur and chlorine, present in the fuel in concentrations of 0.03 wt.%, were recovered at less than 60%. Both elements were present in fuel at concentrations close to the detection limits and closures were sensitive to small errors in analysis and material flows. A substantial fraction of chlorine remained in the gas phase past the cyclone at a stack exit temperature of 350°C. Total concentrations of silica and alumina in the fuel were 1.4 wt.% and 0.5 wt.% respectively (both values imply substantial soil contamination of the fuel, predominantly the wood fuel). Recoveries are good given the high silica and alumina concentrations of the bed media. Titanium, phosphorus and iron were mostly found in the coarse ash fraction and in the particle fraction of the stack flow. Alkaline-earth and alkali metals in the bed are primarily indicative of the fuel ash remaining in the bed. Substantial fractions of these elements were captured in the coarse ash fractions removed in the horizontal pass and by the cyclone, although about 10% is associated with the fine particle hction passing the cyclone. Yields for all elements on the deposit probes are low due to the low deposit masses.
724
Table 7 Material balances for combustion (%). Deposit Particle Gas Bed Ash Probes Phase Phase C <0.5 <0.5 nd < 0.5 861100 H nd nd nd nd 881102 N nd nd < 0.5 nd 1021100 S -18 39 2 23/27 < 0.5 c1 6 14 2 16119 8/10 SiOz -112 183 < 0.5 516 < 0.5 A1203 -109 124 < 0.5 213 < 0.5 Ti02 -780 633 nd 13115 nd Fe203 7 117 < 0.5 10112 < 0.5 CaO 24 96 < 0.5 12/14 < 0.5 MgO 12 56 < 0.5 10111 < 0.5 Na20 14 54 < 0.5 718 < 0.5 KzO 61 42 < 0.5 9111 < 0.5 p2os 7 75 < 0.5 9/10 < 0.5 nd = not determined. 'via Direct methodlcarbon balance.
C H N S
c1 SiOz A1203 Ti02 Fez03 CaO MgO Na20
K20 PzOs
Bed 1 nd nd 19 22 -5 1 -63 676 12 58 9 9 83 131
Closure (DIC)' 861100 881102 102/100 55158 46/50 97/98 91 94 1161117 130/132 79/80 78/79 1111112 94/95
Table 8 Material balances for gasification (%). Deposit Particle Gas Ash Probes Phase Phase Closure 13 < 0.5 1 56 72 1 < 0.5 < 0.5 106 108 < 0.5 < 0.5 < 0.5 101 101 19 < 0.5 3 < 0.5 46 67 < 0.5 10 49 149 51 < 0.5 4 < 0.5 84 1 < 0.5 < 0.5 73 9 200 1 19 nd 134 127 1 11 < 0.5 132 113 1 12 < 0.5 182 54 1 8 < 0.5 73 39 1 4 < 0.5 56 70 < 0.5 9 < 0.5 156 52 < 0.5 8 < 0.5 165
nd = not determined
3.5.2 Gasification Material balances for gasification are listed in Table 8. Roughly 13 wt.% of carbon from the fuel was retained in ash. The carbon recovery is 72%. The relatively poor closure is possibly the result of error in the primary air flow measurement or in the gas composition, but is also partly due to exclusion of tars and higher molecular weight hydrocarbon gases from the analysis. The hydrogen and nitrogen balances, however, show relatively good closure utilizing direct air flow measurements. Sulfur again shows poor recovery (46%), but in this case chlorine is found in excess (149%). The low concentrations in the fuel again lead to high sensitivity in the balances for these elements. A subtantial fraction of C1 appears in the gas phase past the cyclone at temperatures around 250°C.
725
Spent bed material contains carbon and ash along with silica and alumina from the original bed media (54% SiOz, 44% A1203). About 50 wt.% of silica from the fuel was retained in ash and a smaller amount of alumina was detected in ash, The alumina recovery is 73%, with 84% recovery for silica. The titanium, iron and phosphorus recoveries are high, ranging from 132% to 165%. Recoveries of potassium and calcium are also high (156% and 182% respectively), whereas magnesium and sodium balances fail to close (73% and 57%). In general, element recoveries for the gasification conditions are poorer than for combustion, possibly due to higher concentrations of organic material in the ash and bed fractions.
Turn, et al.,” investigated alkali partitioning in fluidized bed gasification of sugarcane bagasse and banagrass. Fuel element recoveries varied widely for the different fuels, but for most elements are similar to values obtained here. Alkali recoveries in Turn’s investigation ranged from 90% to 120% for potassium and 58% to 151% for sodium using similar detection techniques. Calcium closure ranged from 97% to 143%. Results for the wood-shell blend used here give alkali recoveries in gasification of 156% for potassium, 56% for sodium and 182% for calcium. Overall elemental balances are impacted by the concentration of each element and recoveries for elements present only in small quantities typically deviate substantially from 100%. Mojtahedi, et aI.,l4 reported experimental results of peat gasification in a pressurized fluidized-bed gasifier. A combined alkali (potassium and sodium) recovery of 105% was obtained. Alkali was distributed between the cyclone catch (97%), post-cyclone particulate matter (6.7%) and vapor phase in flue gas (1.2%). Unlike the current work, alkalis were not accounted for in spent bed material but the results are otherwise similar. For combustion of the wood-shell blend, potassium was distributed among the cyclone and bed ash (103%), post-cyclone particulate matter in the stack flow (9-11?40), and the vapor phase of the stack gas (0.1%). Sodium and calcium in combustion were distributed among ash (68% and 119%) and post-cyclone particulate matter (8% and 14%), with small amounts in the vapor phase of the stack gas (
3.6 GAS-PHASEI N 0 RGANICS 3.6. I Cornbustion In combustion, vapor phase concentration of potassium inside the reactor at the disengagement section was 0.9 ppmw and decreased to 0.6 ppmw at the cyclone exit. The measurements of the cyclone exit flow also indicated vapor phase silica at a concentration of 0.3 ppmw, but some of this may be in the form of fine particles passing the alkali filter. Thus, some amount of the alkali attributed to the vapor phase may also have been present as or on fine particles. Gas phase chlorine concentration inside the reactor at the disengagement section was 16.4 ppmw, and 2.4 ppmw at the cyclone exit (indicating an 85% removal rate via condensation or reaction on particles). A small amount of iron (0.03 ppmw) was present in the sampled gas from inside the reactor, again potentially as a component of fine particles passing the filter and originating in part from the reactor steel itself.
3.6.2 Gasification In. gasification, gas phase chlorine at the disengagement section was detected at concentrations of 38.3 and 40.8 ppmw for the 2 and 5 pm filters, respectively. The increasing concentration for the larger filter is suggestive of greater fine particle mass passing the filter and landing in the impingers, although the difference may also be due to
726
small differences in fuel and operating conditions for the two tests. Gas phase chlorine concentration at the cyclone exit was 73.5 ppmw, higher than in the disengagement section, although here again the difference may be the result of small changes in fuel and operating conditions. Gas phase alkali was not detected in the post-flare flow due to dilution with ambient air in the exhaust duct. Only Si was detected with a concentration of 0.4 ppmw. A vapor phase potassium concentration of 2.8 ppmw was detected in the reactor at the disengagement zone. Gas phase chlorine was detected in all cases except during post-flare sampling. Chlorine concentration ranged from 2.4 ppmw to 16.4 ppmw during combustion, and from 38.3 ppmw to 73.5 ppmw during gasification. The higher concentrations in gasification are partly the result of the higher fuel-to-air ratios relative to combustion. Generally chlorine concentration in the reactor during gasification was higher by a factor 5 or more compared to combustion for sampling at the same location (hel-air ratio was higher by a factor 10). The fraction of C1 in the gas phase at the disengagement section was similar for both gasification and combustion (45 and 58%, respectively), but a greater fraction of C1 remained in the gas phase at the cyclone exit in gasification. Alkali are more readily released from the he1 as alkali chlorides when chlorine is present.ls Residence time and temperature in the combustion zone are sufficiently high to convert alkali chlorides to alkali sulphates in the presence of SO*,and this mechanism is likely contributing to the high sulfate concentrations of the probe deposits. Alkali in the gas phase condenses on solid particles as the process stream cools, trending towards completion in the interval of 550 to 400°C. Cooling occurs rapidly in the disengagement zone due to the increasing surface area and separation of particles. Temperatures at the disengagement section ranged from 640-750°C (combustion) to 590°C (gasification), cooler than the hottest areas in the reactor (870°C to 750°C). As a result of the lower temperatures some of the alkali vapor would have condensed on particles. However, the fraction of K in the gas phase during combustion was always higher than in gasification, varying between 0.5 and 1.6% for combustion, but less than 0.05% for gasification. This implies a lower volatilization of potassium from the fuel during gasification, largely the result of the lower temperatures relative to combustion. 4 CONCLUSIONS
Gas phase K concentrations during combustion of the wood-almond shell fuel blend decreased from 0.9 ppmw at the disengagement section of the reactor (sampling temperature 750°C) to 0.6 ppmw at the cyclone exit (350°C). Higher fuel feed rates for gasification led to higher K concentrations, 2.8 ppmw at the disengagement section (59OoC), but lower temperature at the cyclone exit (250°C) resulted in substantially lower K concentration (<0.05 ppm). The ratio of vapor phase to fuel K at the disengagement section during combustion was twice that for gasification, 9.3 x vs. 5.6 x loa.
Gas phase chlorine concentrations decreased from 16 ppm in the disengagement section to 2 ppm at the cyclone exit during combustion. Gas phase C1 concentration varied from 38 ppm (disengagement section) to 74 ppm (cyclone exit) in gasification. For both combustion and gasification, gas phase CI concentrations at the disengagement section ranged up to about half the theoretical stoichiometric yield from the fuel (29 pp~llwfor combustion, 150 ppmw for gasification). Sodium concentration in the fuel was one-tenth that of potassium, and gas phase concentrations remained below the detection limit (0.05 ppmw).
727
Compositions of combustion deposits from probes in the disengagement zone and horizontal pass show that virtually all of the potassium in the deposits can be accounted for as chloride (30%) and sulfate (70%). Although the probe locations employed are not representative of commercial gasifier designs, the collected deposits show K to be present primarily in other forms, either as bed particles or soil minerals adhering to these organic material rich deposits. Both combustion and gasification regimes show substantial retention of inorganic constituents in the bed. Potassium was partitioned roughly equally between the bed and the coarse ash fraction removed by settling in the horizontal pass and by the cyclone. Of the inorganic elements in the fuel, approximately 5-10% of each appeared in the fine particle phase of the cyclone exit flow, both in combustion and gasification. A greater fraction of K was found in the gas phase during combustion. At the cyclone exit with a temperature of 350"C, 1.6% of the potassium was still in the gas phase, compared with less than 0.05% in the gas phase during gasification at an exit temperature of 250°C. The reduction in both the volatile fraction of fuel K and the fraction of K appearing in the gas phase at the cyclone exit support the use of gasification to mitigate alkali fouling of downstream equipment. 5 REFERENCES
L.L.,Miles, T.R. Jr, Miles, T.R., 1998, Combustion properties of biomass. Fuel Processing Technology 54 (1998) 17-46, Amsterdam, Elsevier. 2 Salmenoja, K., Mgkelii, K., Backman, R., 1996, Superheater corrosion in environments containing potassium and chlorine. Journal of the Institute of Energy, September 1996, 69: 155-162. 3. Miles, T.R., T.R. Miles, Jr., L.L. Baxter, R.W. Bryers, B.M. Jenkins and L.L. Oden. 1996. Boiler deposits from firing biomass fuels. Biomass and Bioenergy 1O(2-3): 125138. 4. Salour, D., Jenkins, B.M., Vafaei, M., Kayhanian, M., 1993, Control of in-bed agglomeration by fuel blending in a pilot scale straw and wood heled AFBC, Biomass and Bioenergy 4(2):117-133. 5 Natarajan, E., Ohman, M. Gabra, M. Nordin, A., Liliedahl, T., Rao, A.N., 1998, Experimental determination of bed agglomeration tendencies of some common agricultural residues in fluidized bed combustion and gasification, Biomass and Bioenergy 15(2):163-169. 6. Natarajan, E., Nordin, A., Rao, A.N., 1998, Overview of combustion and gasification of rice husk in fluidized-bed reactors, Biomass and Bioenergy 14(5/6):533-546. 7. Rizeq, R.G., Shadman, F., 1989, Alkali-induced agglomeration of solid particles in coal combustors and gasifiers, Chem. Eng. Comm. VoI.81,83-96. 8. Dayton, D.C., French, R.J., Milne, T.A., 1995, Direct Observation of Alkali Vapor Release during Biomass Combustion and Gasification. 1. Application of molecular beadmass spectrometry to switchgrass combustion, Energy & Fuels 9:855-865. 9. French, R.J., Milne, T.A., 1994, Vapor Phase Release of Alkali Species in the Combustion of Biomass Pyrolysis Oils, Biomass and Bioenergy 7( 1-6):315-325. 10. Dayton, D.C., B.M. Jenkins, S.Q. Turn, R.R. Bakker, R.B. Williams, D. Belle-Oudry and L.M. Hill. 1999. Release of inorganic constituents from leached biomass during thermal conversion. Energy and Fuels 13(4):860-870. 1. Jenkins, B.M., Baxter,
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1 1 Valmari, T., Lind, T.M., Kauppinen, E.I., 1998, Field study on ash behaviour during circulating fluidized-bed combustion of biomass. 2. Ash deposition and alkali vapor condensation, Energy & Fuels 13: 390-395. 12.Turn, S.Q., Kinoshita, C.M., Ishimura, D.M., Zhou, J., 1998, The fate of inorganic constituents of biomass in fluidized-bed gasification, Fuel 77(3):135-146. 13. Manzoori, A.R., Agarwal, P.K., 1991, The fate of organically bound inorganic elements and sodium chloride during fluidized bed combustion of high sodium, high sulphur low rank coals, Fuel 7 1 : 5 13-522. 14.Mojtahedi, W., Kurkela, E., Nieminen, M., 1990, Release of sodium and potassium in pressurized fluidized-bed gasification of peat, Journal of the Institute of Energy, 63(456):95-100 15. Olsson, J.G., Jaglid, U., Pettersson, J.B.C., 1997, Alkali metal emission during pyrolysis of biomass, Energy & Fuels 1 1 :779-784 16. Salo, K., Mojtahedi, W., 1998, Fate of alkali and trace metals in biomass gasification, Biomass and Bioenergy 15(3):263-267.
729
Evaluation of a Novel Granular Bed Filtration System for High Temperature Applications H. Risnes and 0.K. Sarnju Department of thermal energy and hydropower, The Norwegian University of Science and TechnoIoD, Trondheim, Norway
ABSTRACT: Panel Bed filter is a new alternative for particle removal. The filter is based on the use of sand or other granulated material as a filtration medium, and is operating in surface filtration mode. The concept was developed in the United States and further developed in Norway. Studies were undertaken to test a full-scale single filter element (1150m3h) of commercial design on combustion flue gas from a 5MW bark-boiler. Results from the field filtration experiments, in total 80 hours of operation, show a reduction in the dust concentration from approx. lgMm3 at the filter inlet to an average dust content of 1.7mg/Nm3at the filter outlet (average filtration efficiency equal to 0.9983). The filter unit was operated at approximately 190"C, with a superficial filtration velocity ranging from 4 to 15 c d s . The results obtained from the field tests on a single element has been verified in a full scale pilot plant with a capacity of approx. 30,000m3/hoperating at 200°C. Further, a lab scale filter unit, based on the German filtration standard VDI-3926, type 2, has been built for testing on biomass gasification gases, at higher temperatures.
INTRODUCTION The fast increase in severity concerning environmental legislation has provoked a renewed interest in particle removal from gas streams. Emphasizing on the collecting of micrometer particles and tolerance to fluctuations in the gas stream conditions. Also the potential usage of alternative energy camers has recently attracted considerable attention. Substantial effort has, therefore, been directed toward developing advanced technology for power generation with high efficiency and reduced emissions. Power systems, such as the integrated gasification combined cycle (IGCC) and pressurized fluidized-bed combustion (PFBC, requiring gas turbines), are widely anticipated as future technologies to produce electricity and steam from coal and biomass. Removal of particulate and other emissions is required for these processes, to protect the gas turbines against erosion and corrosion. To maintain high thermal eficiency, the gas cleaning should be carried out at high temperatures (at least 350°C for IGCC and 750°C for PFBC). Thambimuthu' and
730
Clift and Seville2 have published extensive reviews of gas cleanup at high temperatures. For particle cleanup; cyclones, impact separators, fabric and fiber filters, granular beds, and electrostatic percipitators for low-temperature gas cleaning have been widely used in the industry. However, for hot gas filtration with ceramic filters, the long-term durability, alkali corrosion, cleanability, thermal shock, and particulate penetration into filter media has been recognized as major concerns. Traditional bag filters are sensitive to high temperature and hot particles, electrostatic percipitators have a relatively high cost of installation in smaller plants and multi-cyclone cleaners are not sufficiently efficient to meet the new purification demands. The term “granular bed filter” describes devices in which particles are removed from a fluid by passing it through a bed of an unbonded granular medium. Granular bed filters are cleaned in situ, as described by Coury et al.3, or alternatively the medium may be displaced intermittently, as in the “panel bed” configuration described in this paper. The objective of this paper is to describe and evaluate field tests conducted on a new type of cake forming granular bed filter element, of commercial design, and relate these findings to an existing pilot plant of full scale. Sand filters have also numerous advantageous properties compared with electrostatic percipitators and bag filters:
0
0
0
The filter material is easy accessible and cheap to replace, insensitive to hot particles, temperature fluctuations and corrosive elements in the flue gas. Compactness, due to a superficial filtration velocity which is up to 10 times higher than for bag filters Granulated material can be added to the filter to provide chemical processing of the gas. Measured filtration efficiency is at least equal to or better than for bag filters.
The Panel Bed Filter exposes a large number of sand surfaces for collecting deposits of fly ash. Louvered walls, somewhat resembling venetian blind, hold a sand bed in a tall relatively narrow “panel”. Gas flows in the horizontal direction across the sand bed, and the dust in the gas deposits forming a filter cake upon the sand surfaces at the gas-entry “face” of the panel. Cake filtration occurs when the deposit layer becomes thick enough so that the properties of the dust cake, rather than the supporting medium, determine filtration characteristics. Besides particle properties and forces, such as electrostatic forces, the structure of the dust cake can be influenced by fluid conditions such as gas velocity and temperature and the geometry of the filter medium. The dust loading process for cake forming filters generally proceeds through four regimes: i) clean filter filtration, ii) initial depth filtration regime, iii) transition filtration regime, and iv) dust cake filtration regime. Under high particle loading conditions, the process will often pass through the first three regimes in a very short time to reach the dust cake filtration regime. For a low and stable conditioned pressure loss, successful pulse cleaning of the filter is clearly essential. There are two parts to this problem: i) the geometry and reservoir pressure of the cleaning system must be selected to deliver the required conditions for dust cake removal at the filter surface, and ii) the magnitude of the cleaning action which is necessary to detach the cake must be known.
73 1
THEORY The pressure loss across the dust layer during filtration as well as the energy necessary to remove the deposited layer in the cleaning cycle are strongly related to filter cake properties, including cake porosity. The porosity can, in principle, be estimated from macroscopic values of the pressure drop across the dust layer for a given gas flow rate, if the total cake mass is known (Darcy's law). Generally the cake porosity is known to decrease; i) with increasing particle size, ii) towards the cakefilter interface and iii) as cake grows with time, for a given position within the cake4-'. On-line measurements of cake thickness during filter-system operation have not been feasible in most systems. However, when the pressure drop across the filter system, the gas flow rate, and the particle concentration in the incoming gases are measured, these data can be used to calculate the mass of filter cake on each filter. The overall pressure drop can be modeled
Applying K , as the specific resistance of the filter medium itself, including deposited particles, and K2 as the specific cake resistance, Eq. 1 can be rewritten as
-
AP = K , p f + K , p , W A
,where
W A = cvfAt
(2)
Where ,u is the gas viscosity, vf- gas velocity across the exposed filter cake surface and W, equals dust load. It is common to combine the specific resistance and viscosity, and replace K I and K2 (in Eq. 2 ) with
K2 ' is m a d y determined by the filter cake structure; thus closely related to the physical properties of the dust itself.
THE WORKING PRINCIPLE Fig. 1 shows a cross section of a filter element and the principle for the construction of the filter is illustrated. The figure shows that the outer surface of the filter element consists of louvers. There are two layers of sand along each outer surface. A layer of fine sand on the outside and a layer of coarse sand on the inside facing the clean gas duct in the center. The sand fractions are physically separated. All sand has an angle of respose fixed on the basis of the dimension of the sand grains and the distribution of size. By placing the louvers at a common distance adapted to the angle of respose for the fine sand, the sand remains stable between the louvers. In this way the open sand surface is exposed to the dust containing gas.
732
Most granular filters are based on the principle of deep bed filtration (particle penetration deep into the filtration medium) and the purification degree has been relatively poor for particle sizes in the area 0.1-5pm.
N O W OPERATION
CLEANING
Fig. 1 Cross section of a filter element and the principle for the construction of the filter. The Panel Bed Filter differs from other granular filters in that filtration takes place on the surface of the granular medium, through the building up of a filter cake with "roots" in the outside layer of sand. The size of the sand grains and the speed of the gas impinging on the filter surface are important parameters with regard to the formation of the filter cake and the efficiency of the filter. When the filter cake has built up to a thickness whereby resistance over the filter has reached its desired maximum value. The filter is cleaned by stopping circulation and transmitting a short pressure pulse, with duration of only a few milliseconds, into the filter element in the reverse direction to normal flow. The pressure wave fluidizes the sand and the sand moves slightly horizontally between the louvers. As a result, some of the sand together with the filter cake falls from the filter element, thereby cleaning the filter surface. The mixture of dust and sand is precipitated into a hopper for dust separation arid sand recirculation. EXPERIMENTAL SET-UP Field tests were conducted on a slip-stream connected downstream of the multicyclone, on a SMW-bark fired boiler. The instrumentation of the set-up was as indicated in Fig. 2. Connections for isokinetic dust sampling (gravimetrical determination) were mounted on the upstream and downstream side of the filter unit. The following signals were connected to the data acquisition unit for continuous sampling of: Pressure difference, between filter d e t and outlet (PT1); volume flow 733
(torbar pitot-tube installed downstream of the filter (PT2)); temperature at the isokinetic sampling points (TT1-inlet / TT4-outlet) and connections for pressure difference measurements (TTZ-inlet / TT3-outlet). The volume flow measurement is located close to TT4.
n Pressure pulse Clean gas side Dividing sheet Dirty gas side
-c3
Isokinetic sampling point
I@
Filter element
-
Electrically heated filter wessel
Clean gas to stack
Isokinetic sampling point Slip stream from flue gas channel
Filter material & dust
Fig. 2: Flow-sheet for the experimental set-up during the filtration experiments. THE FILTER ELEMENT Fig. 3 shows a cross section of a filter element. The design is quite similar to that of bag-house filters and ceramic / metallic filters. Each filter element is suspended fkom a dividing sheet between pure and contaminated gas at the top of the filter housing. The filtration experiments were conducted on a full-scale filter element: height of approximately 3500mm, width of 600mm and a thickness of 250mm. This size provides a filter area of approximately 3.25 m2.
734
Louvres
Fig. 3: The design of a commercial scale filter element.
EXPERIMENTAL RESULTS FILTER CAKE FORMATION AND FILTER REGENERA TION
The Panel Bed Filter unit was successfully operated for 80 hours on flue gas from a 5 MW bark boiler. Stable operation was maintained, despite significant fluctuations in flue gas dust concentration (at the filter inlet, c = 600-1200mg@4m3) and filter operation (41US114cm/s).The observed instability in dust load makes extraction of exact values for the specific cake resistance less accurate. The influence of filtration velocity on resulting pressure build up and cycle-time is illustrated in Fig. 4. Note that the displayed offset in the base pressure loss for a clean filter (H)has been used to enhance readability Measured pressure build up versus accumulated amount of fly ash is shown in Fig. 5 a). Calculated fly ash accumulation is based on averaged dust load measured gravimetrically during the respective filtration cycle. Observed pressure loss is close to ideal behavior for surface filtration (i.e. incompressible cake formation and constant dust concentration). Calculated specific cake resistance is shown in Fig. 5 b).
735
Fig. 6 shows the variation in K2 ‘ and residual pressure drop during six hours of operation. K2’ is calculated from Eq. 4. Calculated mean specific cake resistance equals 6.5.104 [s-’1, with a standard deviation of approx. 15%. vJ is defined as the velocity across the exposed filter surface (VJ= 2 Us).Observed fluctuations in pressure build-up did not result in any increase in the residual pressure drop. The residual pressure drop could be maintained at a constant level. The average filtration efficiency was 0.9983. Filter regeneration was conducted with off-line pulsing (P,unk= 2barg and total sand spill of 10 kg).
Table 1 Commercial PBF-filter element: operating conditions during the field tests. Operating conditions Filter material
Olivine sand (AFS50) Virgin and circulated up to three times 190°C 500-2000 m3/h 4-14 C ~ Xs0,, = 12.3 pm; X95,3 = 62.2 pm 600- 1200mg/Nm3 approx. 30 ms 5-10 liters 1-2 barg approx. 10 kg
Mean filter temperature Volume flow Superficial velocity (Us) Fly ash from a bark boiler Fly ash concentration Duration of pressure pulse Tank volume Tank pressure Sand spill during cleaning
~
M,av [ c d s ] 4.0
~
M,av [ c d s ] 8.9
~
S
M,av [ c d s ] 10.8
15 10
0
1000
2000
3000
4000
5000
6000
7000
time [s]
Fig. 4 Comparison of the pressure build-up at three different operating conditions (4 (top curve), 8.9 and 10.8 (bottom curve) c d s ) .
736
t
5.0
4
8.0
m
mIn
7.0
4.0
4
1
z
1.o 1.o
: l i l L = 0.00
0.0
2
4
6
8
10 12 14 16
accurmlated m s s of fly ash [ng/cm2]
Fig. 5 (a) Pressure build-up and (b) Calculated specific cake resistance at operating conditions close to design load. (A - vf = 17cm/s, c = 962 mg/Nm3; * - 21.3cm/s, 1226 mg/Nm3; 0-23.lcm/s, 810mg/Nm3; 0 - 18.8cm/s, 1261 mg/Nm3)
0
1
2
3
4
5
6
7 8 91011121314151617 Filtration cycle
Fig. 6 variations in specific cake resistance and residual pressure loss during a filtration cycle (cycle time vaned from 20 to 27 minutes, withm this series).
RESULTS AND DISCUSSION
The calculated specific cake resistance (K2'), for design conditions, is found to be within the range of published values, see Table 2. Note the high value of vf applied in the present work. Due to the significant fluctuations in the flue gas dust concentration calculated values should be considered as being of approximate nature.
737
Fig. 7 a) shows both base and maximum pressure loss (just before and after regeneration respectively) for some sample filtration series. The base pressure loss (M0)shows a linear dependency of Us.Increase in the residual pressure drop due to depth filtration, dust re-entrainment and fractional cleaning is well known to cause serious filter degrading and may lead to total filter blinding. During the 80 hours of field operation such systematic degrading effects has not been observed. Table 2 Comparison to literature values, derived at approximately 200°C
4
Dust
[pm] X50,3 = 12.3; X9,3 = 62.2 0.5-2 3 s & ~ s00~
Present work NaCI, NH,CI, Alz03 (HEPA-filter)" Steinkohle-flughashe (Ceramic element)" Silica (Metal fiber fleece)"
filter
X50,o= 0.53; x9,0=1.6 X5,,,,,=0.7; X90.0=2.74
Quarz (Metal fiber fleece & ceramic fiber fleeze)"
c vs K2'.10-' [g/m3] [ c d s ] [s-'] 1 18 0.65f0.01
2.45-3.0 3
2-20 1.5-1.7
1
2.5
0.028
25
2.5
0.001
3500 T
0.0
2.0
4.0 6.0 8.0 10.0 Superficial wlocity, Us [cmls]
12.0
14.0
Fig.7 Operationalrange for the Panel Bed Filter, pressure loss before ( 0 ) and after ( 0 ) regeneration. Normally bag house filters operate with superficial velocities in the range 0.1-2 c d s and Doapproximately equal to 900Pa. Superficial filtration velocities typical
for the Panel Bed Filter, is approximately ten times higher than for traditional fabric filters. Thus, the Panel Bed Filter can operate at substantial higher filtration velocities at comparable pressure loss and filtration efficiency (above 98%). Filtration efticiency as discussed above, has been verified in a commercial scale filter plant. This pilot plant, Fig. 8, includes a system for discontinuous separation of dust from the used filter bed material. The cleaned bed material is re-circulated by pneumatic transport to sand silos located at the top of the filter vessel. A summary of
738
typical operating conditions, at design load, is given in Table 3 . Compared to the field-testes described above, this system represents a scale up by a factor of 27. A picture of the filter plant is given in Fig. 9. Cleaningsystem
iP_
. i I
n*A
cyclone
Fig. 8 Full-scale pilot plant, flow-sheet. Table 3 Commercial PBF-filter element: operating conditions during the field tests. Operating conditions Bark boiler of 5 MW Flue gas origin: Olivine sand (AFSSO) Filter-matieral approx. 200°C Filter temperature 27 Number of filter elements Design load: Superficial velocity (Us)9 c d s Volume flow approx. 30,000 m3/h Fly ash concentration (after multi-cyclone) 1000mg/Nm3
739
Fig. 9 The commercial scale pilot-plant, capacity of 30,000m3/h. The filter vessel is seen at the right side, with the 5MW boiler house located in the background.
In addition the development of this filter concept for a bark fired bolier application, a research program has been undertaken for evaluation of high temperature applications such as biomass gasification gases. For this purpose, a laboratory filter test unit for application at temperatures higher than >550°C, has been built. The filter closely resembles the system described in VDI-3926,type 2 13. Tests on this unit will be started shortly. CONCLUSION
Field tests, on a single filter element of commercial design, have been successfully conducted at typical operating condition. Measured filtration efficiency is comparable with bag house filters, for filtration velocities up to 14 cmls. The applied filtration velocities are extremely high compared with bag filters which normally operates in the area of 1.0 to 2.5 c d s . The pressure drop over the sand filter shows that even if filtration speed is increased dramatically, it is possible to work within the same field as bag filters, i.e. from 900 to approximately 2200 Pa. Average dust concentration downstream of the filter unit was measured to be well below 5 mg/Nm3.Measured filtration efficiency at 190°C and 9 c d s corresponds to a filter efficiency above 99.83%, verified in a full-scale pilot plant. Calculated average 740
specific cake resistance (K2') equals 6.5 (+ 0.1) .lo4 [s-'1 at operating conditions, which is withm the range found in the literature. A more detailed study of the filter cake behavior will be carried out with the new laboratory unit. The Panel Bed Filter concept has been successfully demonstrated in full scale operation, and more detailed filtration behavior has been studied. Filter cake data have been obtained at very high filtration velocity and reasonable pressure drops. Very low dust concentrations are observed.
NOMENCLATURE dp - pressure loss [Pa] APo- pressure loss across a clean filter element [Pa] &'Firrer - pressure loss across filter material, including deposited particles [Pa] dpcuk- pressure loss across filter cake [Pa] Q - volume flow [m3/h] K - specific resistance [s-'1 W, - accumulated dust mass [g/m2] t - time [s] Us- superficial filtration velocity [ c d s ] vy- average gas velocity thorough filter cake [ d s ] c - dust concentration in flue gas [kg/m3] ,u - gas viscosity [Pa s]
REFERENCES 0. A. M. Squires, Professor at Virginia Polytechmc Institute & State University 1. Thambimuthu, K. V. (1993) Gas cleaning for advanced coal-based power generation. IEACW53, IEA Coal Research, London 2. Clift, R. & Seville, J. P. K. (1993) Gas cleaning at hlgh temperatures. Blackie Academic & Professional, New York 3. Coury, J.R., Thambimuthu, K.V. & Clift, R. (1987) Capture and rebound of dust in granular bed filters. Powder Technology v 50 n 3 p 253-265 4. Aguiar, M. L. & Coury, JR. (1996) Cake formation in fabric filtration of gases. Industrial & engineering chemictry research Vol. 35, No. 10, p 3673-3679 5. Hoflinger, W., Stocklmayer, Ch. & Hackl, A. (1994) Model calculation of the compression beahviour of dust filter cakes. Filtration and Separation v 3 1 n 8 p 807-811 6 . Kono, H. O., Jordan, B., Ohtake, T. & Duane, H. S. (1998) Formation and Measurement of the porosities, tensile strengths, and deformation coefficient of gasification filter cakes at operating temperatures and pressures. Aerosol Science and Technology v 29 p 236-245 7. Rothwell, E.(1985) Analysis of fabric dust filtration. I: Model observation of dust cake formation. Filtration and Separation v 22 n 5 p 3 18-324 8. Schmidt, E. & LoeMer, F. (1991) Analysis of dust cake structures. Particle & Particle Systems Characterization v 8 n 2 p 105-109 9. Stocklmayer, Ch. & Hoflinger, W. (1998) Simulation of the filtration behavior of dust filters. Simulation Practice and Theory v 6 p 281-296
74 I
10. Pilz, T. & Loeffler, F. (1995) Mluence of adhesive and cohesive particle properties in filtration on surface filters. Chemie Ingenieur Technik Vol. 67, Nr. 6, p 745-149 11. Hajek, S. Peukert, W. (1997) Comparison of ceramic and metallic filter units for bigh temperature filtration Chemie Ingenieur Technik Vol. 69, Nr. 3, Mar 1997, side 341-345 12. Novick, V.J., Momon, P.R. & Ellison, P. E. (1992) The effect of solid particle mas lading on the pressure drop in HEPA filters. J. Aerosol. Sci., Vol. 23, No. 6. p 657-665 13. VDI-Richtlinie 3926 Testing of filter media for cleanable filters, Beut Verlag, Berlin, 1994.
742
Combustion processes in a biomass fuel bed Experimental results Marie Ronnback, Monica Axell, Lennart Gustavsson SP Swedish National Testing and Research Institute, BOX 857, SE-501 15 Bords, Sweden Henrik Thunman, Bo Lecher Chalmers University of Technology, Energy Conversion, SE-412 96 Goteborg, Sweden
ABSTRACT: Combustion processes in a biomass bed are investigated experimentally. Special attention is paid to the influence of primary airflow and particle properties on the ignition front, its temperature and on the composition of the gas leaving the front. Two test rigs have been built: a large rig in the same size as a boiler for domestic use and a small laboratory test rig. In both rigs the ignition front moves in opposite direction to the primary airflow. Three combustion regimes are identified: a sub-stoichiometric regime with incomplete consumption of oxygen, a sub-stoichiometric regime with complete consumption of oxygen and an over-stoichiometricregime. The results show that a fuel with higher density and thermal conductivity (but in other respects similar to other fuels) has a wider sub-stoichiometricregime where oxygen is cornpletely consumed. If the particle size is increased (for the same fuel quality) the airflow range of this regime becomes shorter and starts at higher airflow. INTRODUCTION Fixed or moving bed combustion is the most common technology for biofuels. Incineration of wastes and gasification of biomass are also often performed in a bed. The design of the grate (stationary, vibrating, reciprocating or moving) on which the bed rests and the method of fuel supply (from underneath, from the side or from above) depend on fuel characteristics and on the size of the plant. Devices with bed combustion have as common features drying, devolatilization, gasification and combustion occurring in particular zones in the bed. The extension of these zones depends on fuel and air supply and on the initial ignition of the fuel bed. The bed may be operated as a gasifier producing a combustible gas, but it may also be operated with excess air. The purpose of a boiler is to attain complete combustion, although this is achieved downstream of the bed in the form of gas phase combustion controlled by secondary air. For fuels with a high content of volatile matter, the gas combustion downstream of the bed is crucial for emission control. The fuel bed is the first stage in the combustion process and generates the conditions for the latter part. A review of available literature on the experimental simulation of solid fuels confirms that the knowledge of coal combustion is more detailed than that of biomass and municipal solid waste. The knowledge of biomass gasification devices today is extensive; however, see for example La
743
Fontaine and Reed (1) and this knowledge also suits the combustion bed. Also, recently several publications have described experiments on high volatile fuels in onedimensional batch-type reactors. The fuel beds are ignited on the top and the ignition front propagates against the primary airflow. The results from such experiments can be extended to steady-state combustion on a stationary or moving grate. Gort (2) did a thorough investigation on the effects of moisture and volatile content, as well as of particle size. He burned wood cubes, coke and municipal solid wastes in a batch-type laboratory grate furnace. He distinguished three global reaction regimes, depending on the ratio of ignition rate and superficial air velocity. Shin and Choi (3) studied the effects of air supply rate, fuel particle size and calorific value on the combustion of wood cubes in a similar combustor. Depending on the availability of oxygen, distinct reaction zones were identified. Fatehi and Kaviany (4) performed similar experiments in a 7x7-cm2furnace on wood spheres and described two combustion zones. Furthermore, Kuo et al. ( 5 ) pointed out the importance of the arrangement of air supply for combustion efficiency. In their experimental device primary air could be supplied through the grate and through the walls of the fuel chamber. They found that the CO-emissions were related to the oxygen content of the flue gas, depending on the mixing of gas and air in the bed. Horttanainen et al. (6) presented results from experiments on biomass particles of various sizes and moisture content. They focused on the speed of the ignition front as an important factor determining the release of volatiles and that affects the combustion power and the stability of combustion. The aim of the present work is to further analyse the influence of primary airflow and fuel particle parameters on the combustion process. The work ranges from substoichiometric to over-stoichiometric supply of primary air to the bed, to cover the influence of the parameters studied. Experiments were performed with the combustion front moving both upward (as in the studies referred to) and downward (as in a modern boiler with down-draught combustion). The work aims at forming a basis for modelling of wood burning, as well as for design of small boilers. The experiments were conducted so that the results can be used for both batch and continuously fired plants.
EXPERIMENTAL CONDITIONS The experiments were carried out in two test rigs, here called the large and the small rig. The small rig is cylindrical with a diameter of 0.2 m and a height of 0.6 m. The inner wall is made of 3-mm stainless steel to minimise the heat capacity of the wall. It is insulated with a 50-mmglass-wool cover. Primary air is supplied through the steel grate at the bottom of the cylinder and exhaust gases leave through the top, where secondary air is supplied before the gases enter the chimney. The fuel is put on the grate and ignited by a torch from the top. Then, the ignition front moves against the primary airflow towards the grate, see Figure 1. This is a onedimensional representation of a continuously burning fuel bed on a grate, where fuel and air are supplied co-currently and the fuel is ignited on the top of the bed, see e.g. Thunman and Leckner (7).'Two thermocouples (type K) located IS0 mm and 300 mm above the grate measure the temperatures inside and above the fuel bed. Gas analysis is carried out in the gases leaving the bed, before they reach the secondary air supply. The large rig is of the same size as a domestic boiler. The design is chosen to ensure well-defined start- and boundary conditions for the fuel bed. The rig is equipped to measure airflow, weight loss and bed height. Temperatures can be measured upstream, in and downstream of the fuel bed and in the grate by shielded 1-mm thermocouples (type K), mounted both from the side (orthogonal to the movement of the ignition
744
front) and from the air supply. A comparison with shielded 2- and 3-mm thermocouples showed that the mounting of the thermocouples and their sizes gave no significant differences in results. The fuel bed can be observed through sight glasses. The rig works with downdraught combustion, i.e. the primary air is supplied from the top of the fuel bed and the fuel is ignited from the bottom (in this case by electrical spirals in the grate). The ignition front moves against the primary airflow towards the top of the fuel bed, see Figure 2. The fresh fuel moves downwards into the ignition front and influences the ignition front with its weight. Air Combustion gases
I
I
e
I
I Initial ignition
Ignition front
Ignition front
4I
I
Initial ignitlon
Combustion gases
Air
Figure 1 Ignition front moving against the Figure 2 Ignition front moving against the airflow in the small experimental rig. airflow in the large experimental rig.
The fuels were pellets and wood cylinders (pine). The pellets were made of compressed sawdust and had a diameter of 8 mm. The wood cylinders had three diameters, 8, 12 and 34 mm. The proximate analyses and elemental composition of the fuels were almost identical, as seen in Table 1. The density and the thermal conductivity of the pellets are about twice those of the wood. The pellets were burned both in the large and in the small rig, while the wood cylinders were burned only in the small one. Table 1 Fuel characteristics.
Pellet
C (weight-%, dry basis) H (weight-%, dry basis) 0 (weight-%, dry basis) N (weight-%, dry basis) S (weight-%, dry basis) Moisture (weight-%) Hi (MJkg, dry basis) Thermal conductivity (J/mK) (8) Diameter (mm) Fuel density (kg/m3) Bed density (kg/m3) Bed void Air mass flow (kg/m2s)
Wood
50.5 51.4 6.2 6.7 43.1 41.7 0.15 0.047 0 0 8.3 9.1 20.2 19.3 0.32 0.16 8 8 12 34 1259 579 585 581 680 307 305 279 0.46 0.47 0.48 0.52 0.035-0.41 0.07-0.53 0.07-0.53 0.07-0.53 745
Emissions of CO, COz, 02 and THC (Total Hydro Carbon) were measured on-line. The gas was sampled as close to the fuel bed as practically possible. In the large rig, gas was extracted 50 mm downstream of the grate. The distance to the ignition front is then 100-350 mm, increasing as the ignition front moves upward in the fuel bed. In the small rig, gas was extracted in a position 450 mm above the grate. The distance to the ignition front is then up to 450 mm. The gas probe is cooled to quickly stop any combustion reaction, and the temperature of the extracted gas falls to 100°C in less than 200 mm. Because of the high levels of CO downstream of the fuel bed, a dilution devise was used to dilute the sampled gases about 10 times. A conventional instrument could then measure the CO-concentration. The dilution factor was continuously calculated by comparing the COZ-concentration before and after dilution.
RESULTS AND DISCUSSION The ignition rate of the fuel bed is defined as kg ignited fuel per grate area and time (kglmzs). In the small rig this quantity was determined by the times when two thermocouples at a distance of 150 mm reached 500"C, multiplied by the (original) bed density of the fuel. The exact level of temperature chosen (500°C) is not critical, because the temperature change in a fuel layer is quite fast, see Figure 3. The primary air mass flow through the grate is presented as kg air per grate area and time (kg/m2s). Figure 3 shows the temperature at 150 mm and at 300 mm above the grate. The short peak in temperature is the front temperature, most likely influenced by the surface temperature of the nearby particles. In this example the air mass flow is low, and all fuel is not converted but accumulates upstream of the ignition front. The temperature in the partly converted layer is about 100°C lower than in the front. As the bed shrinks the upper thermocouple finds itself above the bed. Because of radiative cooling by the walls the uncorrected temperature is about 200°C lower than the bed temperature.
- -
___
_ *
_ .- - -
Time (min)
Figure 3 Temperatures at 150 mm and 300 mm above the grate and airflow in the small rig. The fuel was 8-mm wood cylinders and the air mass flow 0.11 kg/m2s.
The temperature curves show how the width of the accumulated layer broadens as the front moves downwards. When the ignition front reaches the grate, there is a peak in temperature (at about 43 minutes) caused by combustion of the accumulated char.
746
As can be seen from the gas analysis in Figure 4 there is also a peak in CO- and a dip in COz -emissions followed by a peak in COz while CO goes to zero. The peak in CO coincides with a dip in temperature. When char combustion starts the production of water vapour from drying and combustion falls and the CO-emission increases because of lack of water vapour for CO-oxidation as illustrated by the expression: -d[CO]/dt = k.,[CO][02]'R[H~0] '"exp(-ERT) from Howard (9). The final increase in COz follows from an increase in O2-concentration,enhancing combustion, when the bed is almost finished and very shallow. This dip-and-peak behaviour is not seen at high air mass flows (above 0.17 kg/m2sfor wood cylinders and above 0.40 kg/m2s for pellets).
Time (min)
Figure 4 Emissions of 02,C 0 2 . CO and THC downstream of the grate in the small rig. Data as in Figure 3. The peaks in O2 with corresponding dips in C02, CO and THC originate from back-blowing the analysis system.
In the large rig electrical heating of the grate ignites the fuel. The ignition front moves upward, opposite to the direction of the airflow. The temperature profiles are steep and parallel during the devolatilization. Figure 5 shows temperatures at different levels in the centre of the bed between 10 mm and 400 mm above the grate. Two phases, devolatilization (ca 125-180 min) and char combustion (after 180 min) are indicated by the temperature curves and by the gas analysis downstream of the grate. Figure 6 shows the gas analysis from the same experiment. After bed break-through the bed consists mainly of char and the fuel bums from the top downwards. The temperature in the top layer, where the air meets the fuel, shows peaks to maximum 1100°C. The temperature increases continuously in the remaining bed during char combustion. In this rig, the ignition rate of fresh fuel cannot be determined in the same manner as in the small rig, because the bed shrinks during conversion of the fuel. Instead, the time between the moment when the gas downstream of the bed reaches 500°C and when the ignition front breaks through the top of the fuel bed was measured. The ignition rate in kg/m2sis calculated by dividing the mass of the fuel in the bed by the grate area and the corresponding time. The break-through of the ignition front at the surface of the bed can be determined in several ways: visual observation of flames above the bed, a rise in temperature in the topmost fuel layer and changes in the gas analysis. In the ignition front there is a balance between heat generated by chemical reactions and heat transfer into the fuel particles, to the combustion air and to new layers of fresh 747
y20
140
160
180
200
220
Time (min)
Figure 5 Temperatures in "C at levels above the grate from 10 mm to 400 mm in the large rig. The air mass flow was 0.12 kg/m% and the fuel 8-mm pellets.
Y20
140
160
180
200
220
Time (mln)
Figure 6 Emissions of 02.C 0 2 and CO downstream of the grate in the large rig. Data as in Figure 5 . The peaks in O2with corresponding dips in COz, CO and THC originate from back-blowing the gas analysis system. fuel. The heat transport to new fuel is dominated by radiation, because the direction is opposite to the airflow. The heat transport into the fuel particles depends on the thermal conductivity of the material. The particles used were all thermally thick, i.e. the Biot number = hDpk > 1, and a temperature gradient was present in the particles during devolatilizationand char combustion. As a particle is heated, it dries and devolatilizes. Water vapour and combustible gases are produced and transported to the surface of the particle and out into the gas between the particles, where the gases ignite if the conditions are suitable. Devolatilization of biomass starts already at temperatures about 200°C and spontaneous ignition of
748
wood at 500-600°C.If a pilot flame is present, such as the flame front in a fuel bed, Saastamoinen et al. (10). ignition can take place already at 300-4OO0C, When the volatile matter has left the particles, heterogeneous reactions, such as gasification with COz and H 2 0 and char combustion, begin. Char combustion is a slower process than gas combustion and demands a higher temperature (2 800°C) to be complete. As long as oxygen is present combustion dominates, since gasification is slow.
COMBUSTION REGIMES Figure 7 shows the maximum temperatures and Figure 8 the ignition rates in the small rig with 8-mm wood cylinders and 8-mm pellets as fuel. The lines in Figure 8 represent the theoretical stoichiometric combustion rate that would occur if the fuel was exposed immediately to oxygen, i.e. if the fuel particles were thermally infinitely thin. 15001
I I
1300t----
--
Y
f
c
f 5
0
c
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000
o m
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.
m m
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sub-stoich. b
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7508.01 -
0
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Tmaxwood8mm Tmax pellet 8 mm
00 0.10 0.20 0.30 0.40 0.50 0.60 Air mass flow (kg/m2s)
Figure 7 Maximum front temperature for 8-mm wood cylinders and 8-mm pellets as function of air mass flow rate. The sub-stoichiometricregimes with complete oxygen consumption for the different fuels are marked in the figure.
Based on the results, the combustion process has been divided into three regimes. (1) Sub-stoichiometric combustion with incomplete consumption of oxygen. This regime is found at low primary airflow and is characterised by a clear division in time in a devolatilization phase followed by char combustion. During the devolatilization, partly converted fuel accumulates downstream of the ignition front. Despite the sub-stoichiometriccondition, oxygen is not fully consumed in the bed, either because of slow kinetics or of insufficient mixing between the devolatilized combustible gases and the primary air. As a result, just after the bed where the gas sampling was made, unburned gases as well as oxygen can appear in the sampling gas. The ignition rate and the ignition front temperature are strongly influenced by the primary airflow. Figure 3 and Figure 4 show combustion in this regime from 7 to 11 minutes. (2) Sub-stoichiometric combustion with a complete consumption of oxygen. This regime occurs at higher primary airflow, see Figure 7 and Figure 9,and is characterised by complete consumption of oxygen in the bed. A layer of partly converted
749
fuel forms downstream of the ignition front, but the layer does not grow in thickness as in Regime 1. The devolatilization phase is followed by a short char combustion phase. The influence of primary airflow on ignition rate and ignition front temperature is not so pronounced. The maximum ignition rate and front temperature are found in this regime. Figure 5 and Figure 6 show an example. 0.12:
.
! I
Stachiomtw wood 8 nun
0.08’ .O
......... Stoichiomtrypellet 8 mm I 0.1
0.2
0.3 0.4 0.5 0.6 Alr mass flow (kglm2s)
Figure 8 Ignition rate for 8-mm wood cylinders and 8-mrn pellets as function of air mass flow rate. The lines represent stoichiometriccombustion.
(3) If the primary airflow is increased even more the combustion moves into the overstoichiometric regime and excess oxygen leaves the fuel bed. As the excess air is heated, the bed cools. With increasing airflow the ignition rate slows down and the front temperature decreases. Finally, the combustion is extinguished. The final char combustion phase is negligibly short. INFLUENCE OF DENSITY AND THERMAL CONDUCTIVITY The density and the thermal conductivity of the pellets are about twice those of the wood cylinders, but the fuels used differ very little in elemental composition and moisture content. In general, the differences in front temperatures and ignition rates are small for the fuels, as seen in Figures 7 and 8. The ignition rate of pellets is less influenced by airflow in a wider range, and the temperature in the front is higher for pellets at higher airflows when for wood cylinders. The interesting difference lies in the combustion regimes. Pellets have a wider sub-stoichiometric, full oxygen consumption regime, starting at a lower airflow. At the temperatures and residence times during the change from Regime 1 to Regime 2, the conversion of CO, according to the rate expression of Howard (9). should be complete and no oxygen should be present. The reason for finding CO and 0 2 downstream of the bed is then most likely bad mixing. The pellets are thermally thinner than wood, and their higher thermal conductivity leads to a higher devolatilization rate at the same airflow. This may have contributed to extending the sub-stoichiometric regime with full oxygen consumption to higher airflows in the case of pellets. The higher temperature in this regime (compared to wood) can be explained by a higher burning rate (kg mass losdarea and second) for pellets.
750
INFLUENCE OF PARTICLE SIZE Figure 9 shows the maximum temperature and Figure 10 the ignition rate in the small rig for 8-, 12- and 34-mm wood cylinders. No gas analysis was carried out on the 12and 34-mm wood cylinder experiments and the temperature curves were used to separate the combustion regimes.
h
P
Y
...... . ..... ...
II
. . .... ,,.., . ,..,..... ,... .................,. .. .,. .. ... ,
1300 1100
1
0
v
I
-v v
8
sub-stolch, mp, 0, .I0
.V
Q
sub-stoich, 12 Am, 0,- 0
V
sub-staich, 34 mm, 02=0
V V
"%!O
0.1
0.2
TmaxwoodlZmm Tmaxwood34mrn
0.: 0.4 0.5 0!6 Air mass flow (kglmzs)
Figure 9 Maximum front temperature for 8-, 12- and 34-mm wood cylinders as function of air mass flow rate. The sub-stoichiometriccombustion regimes with complete oxygen consumption for the different sizes are marked in the figure.
Air mass flow (kglm28)
Figure10 Ignition rate for 8-, 12- and 34-mm wood cylinders as function of air mass flow rate. The line represents stoichiometriccombustion. At low airflow the front temperature is lower and the ignition rate appears to be somewhat higher for the largest fuel particles. However, the variations are generally
75 1
small and have to be compared with the accuracy of the experiments, especially for large particles whose size reduces the precision in the calculation of the ignition rate. The sub-stoichiometric regime with complete oxygen consumption starts at a higher airflow for the larger particles. This can be explained by the larger particle being thermally thicker. The resistance towards heat flow into the particles increases with the diameter. More of the heat produced is transported to new layers of fresh fuels and not towards the centre of the fuel particle. This is reflected in a higher ignition rate and a lower front temperature at low airflow. Less combustible gases reach the surface of the large particles, the power produced is lower and the burning rate also becomes lower. When the particle size increases the transition from Regime 2 to Regime 3 is extended to higher airflow. The reason for this has to be further investigated. Parameters of interest are devolatilization rate compared to ignition rate, heat transfer between particles and primary airflow.
COMPARISON BETWEEN THE TWO RIGS A comparison between ignition rates from 8-mm pellets burned in the small and in the large rig shows a reasonable agreement as seen in Figure 11. The front temperatures agree at low air mass flows, but at higher flows the temperatures in the large rig are lower, see Figurel2. Four differences between the rigs can be noted: (1) The small rig has a smaller grate area and channelling inside the fuel bed is less
probable than in a large rig. In the large rig, horizontal temperatures were measured in the bed and in the gas phase downstream of the bed. At several occasions an uneven start phase in the large rig resulted in an uneven ignition profile. An uneven ignition front could lead to a faster ignition of the bed, but the measured ignition rates are consistent and similar in the rigs (Figure 11). The method of measuring the ignition rate over the batch in the large rig is possibly removing any unevenness in the ignition profile. (2) In the small rig there is a risk for air passing between the outer edge of the fuel bed and the walls without taking part in the combustion process. In the large rig there is a seal between the grate and the walls, and any air and gas “channelling” between the fuel bed and the walls is forced to move over the surface of the grate and leave through holes in the grate. This difference in design may lead to higher 0 2 levels downstream of the bed in the small rig. (3) The large rig has a larger heat loss during batch firing, caused by the heavier design. The ignition front, however, is well insulated by fresh fuel on one side and by partly converted fuel or char on the other, and there is only a heat loss at the walls. (4) The most noticeable difference between the results from the two rigs is in the temperature of the reaction front (Figurel2) at airflows higher than 0.3 kg/m2s when the temperature is 200°C higher in the small rig. Also the burning rate in the small rig is higher at these airflows. There are two explanations for the higher temperatures. The first is that in the small rig secondary air is supplied a few decimetres downstream of the bed, and the combustible gases leaving the bed burn with a gas flame that radiates upstream and heats the bed. At lower airflow the ignition front is covered by partly converted material and is not influenced by the heat from the flame. At higher airflow the layer of partly converted material is very thin and the front temperature increases by the heat from the flame. The second explanation is that in the combustion front of the small rig, particles with reduced size may start
752
to fluidise at high airflow, as shown in experiments of Fatehi and Kaviany (4),and that would increase the conversion rate and the temperature in the front. 0.12-
3 0.10 E
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COMPARISON WITH OTHER AUTHORS Figure13 shows the maximum temperature in the small rig using 8-mm wood cylinders with a moisture content of 8.3 % compared to 10-mm wood cubes with a moisture content of 10 % from Got? (2). Figure14 shows the ignition rate of the same cylinders and cubes and of 5-20 mm wood chips with a moisture content of 10.8 % from Horttanainen et al. (6). The results coincide quite well, except for the higher ignition rate of
753
the wood chips. Horttanainen et al. suggest that smaller chips act as pilot flames for larger chips, and this enhances the ignition rate. 1500 - .
1300.-
........... ....................................
.........................
700
........................
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o wood 8 m m m wood 10 mm,Gort
0.3 0.4 0.5 0.6 Air mass flow (kg/mzs) Figure13 Maximum front temperature for 8-wood cylinders and 10-mm wood cubes as function of air mass flow rate.
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0.1
0.2
0.3 0.4 0.5 0.6 Air mass flow (kglmb) Figure14 Ignition rate for 8-mm wood cylinders, 10-mm wood cubes and 5-20-mm wood chips as function of air mass flow rate.
Figure15 and Figure16 show the maximum temperature and the ignition rate in the small rig with 34-mm wood cylinders having a moisture content of 8.3 8 compared to 30-mm wood cubes with a moisture content of 10 8 from Gort (2). The results from the 34-mm wood cylinders are more scattered, probably because the cylinders form a less homogenous bed, which decreases the accuracy of the calculation of the ignition rates. Despite the scatter, the temperatures coincide well. For the larger particles (the cylinders) the maximum ignition rate is found at a higher air mass flow.
754
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wood 34 mrn
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0.2
0.3
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Air mass flow (kglm2s)
FigureZS Maximum front temperature for 34-wood cylinders and 30-mm wood cubes as function of air mass flow rate. 0.12 0.10
0
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(I)
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Other authors have identified combustion regimes similar to this work. Gort (2) divides the combustion process into 1) a partial gasification regime characterised by accumulation of partly converted fuel and full oxygen consumption, 2) a complete gasification regime characterised by a layer of partly converted fuel with constant thickness and full oxygen consumption and 3) a complete combustion regime characterised by a layer of partly converted fuel with constant thickness and incomplete oxygen consumption. These regimes coincide in principle with the ones defined in the present paper with the exception that Gort found no excess oxygen after the bed in the first regime. Shin and Choi (3) identified three combustion modes depending on the air supply rate. When the air supply is low, the reaction rates are controlled by the oxygen supply 755
(oxygen-limited combustion). When the air supply increases, the flame propagation speed increases. However, the flame propagation speed is limited by the reaction rate of the fuel (reaction-limited combustion). When the air supply further increases, excess air cools the bed and puts an end to the flame (extinction by convection). Fatehi and Kaviany (4)identified in a similar way two modes and called them an oxygen-limited and a fuel-limited mode. None of these authors related the modes or zones to a description of bed events. Nevertheless, the oxygen-limited mode should correspond to the sub-stoichiometric regimes, and the reaction- or fuel-limited one to the overstoichiometric regime.
CONCLUSIONS Two test rigs have been built for investigation of the combustion processes in a bed of solid fuels and particularly the influence of primary airflow and of particle properties (size, density and thermal conductivity) on the rate and temperature of the ignition front. Also the gas composition downstream of the bed has been investigated. The following conclusions can be drawn from the results: (1) The two test rigs show the same ignition rates, but the temperatures in the ignition front diverge at higher airflow. This divergence may be caused by a difference in design of the secondary air supply, giving a higher heat transfer to the bed from the gas flame in the small rig. The higher temperatures may also be influenced by fluidisation of particles in the front in the small rig, where the air is supplied from beneath. (2) Three combustion regimes were identified. At low airflow a sub-stoichiometric combustion regime with incomplete consumption of oxygen was found. This regime is characterised by a clear division in a devolatilization phase followed by a phase dominated by char combustion. During the devolatilization, partly converted fuel accumulates downstream of the ignition front. Although the combustion is sub-stoichiometric, oxygen is not fully consumed in the bed. At higher airflow there is a sub-stoichiometric combustion regime with a complete consumption of oxygen in the bed. A layer of partly converted fuel forms downstream of the ignition front, but it is not growing in thickness as in Regime 1. The devolatilization phase is followed by a short char combustion phase. When the primary airflow is increased even more, the combustion moves into the over-stoichiometric regime with excess oxygen leaving the fuel bed. The excess air cools the bed. At higher airflow the ignition rate slows down and the front temperature falls. Finally, the combustion is extinguished. In this regime, the final char combustion phase is negligibly short. (3) The two fuels compared show differences with respect to the combustion regimes. The thermally thinner fuel has a wider sub-stoichiometric, full oxygen consumption regime, starting at a lower airflow. Because of the higher devolatilization rate, the thermally thinner fuel has a sub-stoichiometric regime with full oxygen consumption sustained at higher airflow. (4) Three sizes of the same fuel have been compared. At low airflow the ignition rate is higher and front temperature lower with the larger fuel. The resistance towards heat flow into the particles increases with diameter, and more of the heat produced is transported to new layers of fresh fuels and not towards the centre of the fuel particles. The larger particles are thermally thicker, and the devolatilization rate is lower. The sub-stoichiometric regime with complete oxygen consumption starts at
756
a higher airflow for larger diameter fuel, and the transition from Regime 2 to Regime 3 is extended to higher airflow. The reason for this has to be further investigated. Parameters of interest are devolatilization rate compared to the ignition rate and heat transfer between particles and primary airflow. ( 5 ) In modem combustion equipment, the air supply is divided between primary air to the bed and secondary air to the produced gas. The primary air is kept substoichiometric to enhance reduction of NO. The secondary air is important as a measure for mixing during gas phase combustion. This means that the most important regions for the primary air are the low velocity regimes. Knowledge of the influence of fuel characteristics on the regimes is important for the division between primary and secondary air under various load conditions. However, depending on the design (e.g. batch-wise combustion or steady-state combustion on a transporting grate) the build-up of char or not is also an important issue to consider.
ACKNOWLEDGEMENT This work has been supported by the Swedish National Energy Administration, which is gratefully acknowledged.
REFERENCES 1. La Fontaine, H. & Reed, T. B. (1993) An Inverted Downdraft Wood-GasStove and 2. 3. 4. 5. 6.
7. 8. 9.
10.
Charcoal Producer, Energy from Biomass and Wastes XV, Washington D.C. Gort, R. (1995) On The Propagation of a Reaction Front in a Packed Bed, Ph D Thesis, Universiteit Twente, Enschede. Shin, D. & Choi, S. (2000) The Combustion of Simulated Waste Particles in a Bed, Combustion and Flame Vol. 12 1, pp. 167- 180. Fatehi, M. & Kaviany, M. (1994) Adiabatic Reverse Combustion in a Packed Bed, Combustion and Flame Vol. 99, pp. 1-17. Kuo, J. T., Hsu, W.3. & Yo, T.-C. (1997) Effect of Air Distribution on Solid Fuel Bed Combustion,Journal of Energy Resources Technology Vol. 119, pp. 120-128. Horttanainen, M. V. A., Saastamoinen, J. J. & Sarkomaa, P. J. (1999) Ignition Front Propagation in Packed Beds of Wood Particles, The Swedish-Finnish Flame Day, International Flame Research Foundation, Vaxjo. Thunman, H. & Leckner, B. (2000) Ignition and propagation of a reaction front in cross-current bed combustion of wet biofuels, Fuel (in press). MacLean, J. D. (1941) Thermal Conductivity of Wood, Transaction American Society of Heating and Ventilation Engineers, Vol. 47, pp. 323-354. Howard, J. B., Williams, G. C. & Fine, D. H. (1972) Kinetics of Carbon Monoxide Oxidation in Postflame Gases. Fourteenth Symposium (International) on Combustion, The Combustion Institute, pp. 975-986. Saastamoinen, J. J., Huttunen, M. & Kjaldman, L. (1998) Modelling of Pyrolysis and Combustion of Biomass Particles, ECCOMAS 98, John Wiley & Sons, Ltd. pp. 814-819.
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Combustion performance of New Zealand grown biofuels K. Senelwa', J. Gifford', J. Li', R.J. Hooper2, A. Clemens3 and D. Gong3 New Zealand Forest Research Institute Limited, Rotorua, New Zealand Centrefor Advanced Engineering, Christchurch, New Zealand CRL Energy Limited, Lower Hutt, Wellington,New Zealand
'
ABSTRACT Combustion performance of biofuels such as forest arising, wood process residues (mainly from Pinus rudiutu plantations), and a purpose grown short rotation forest energy crop (Eucalyptus nitens) were analysed and compared to those of a typical New Zealand industrial coal in a 50 kW laboratory scale stoker type combustor. Combustion conditions and products varied significantly among the biofuels and differed from those of coal reflecting the differences in feedstock characteristics. Biofuels had hlgher combustion rates (1.7 - 2.4 g/s, dry basis) than coal (1.6 g / s ) with heat outputs of 22.9 - 44.1 kW (compared to 41.4 kW for coal). The corresponding wet flue gas production rates ranged from 5.4-5.6 m3kg (9.7 m3kg for coal). The flue gas from biofuels had negligible SO2 content. Coal flue gas contained no CO and no CH4 indicating a more complete bum compared to biofuel combustion. The tendency of the bottom ash to slag at the temperatures experienced in the firebox measured by the quantity of ash retained on a 1.676 mm sieve was higher for coal ash than for biomass,. These differences determine efficiencies and viability of biofuel combustion systems, and have important consequences especially when biofuels substitute for, or are cofired with coal in existing coal-fned facilities.
INTRODUCTION The role of biofuels in New Zealand's energy supply is expected to increase from its 1996 contribution of 32 PJ approximately 5% of total primary energy supply (l), to a possible doubling by the year 2005 (2). Drivers for the increase include (i) the potential of biofuels to mitigate energy-related greenhouse gas emissions; (ii) available forestry by-products estimated to be 6.7 million m3 by the year 2020, sufficient to provide 1015% of the total electricity requirements (3, 4); and (iii) the need to diversify national energy supply sources. In the near- to mediwn-term, biofuels can be expected to be utilised in stand-alone combustion plants sized to match site and energy demand, or possibly co-fired with coal in existing coal fired power plants. To ensure increased use of biofuels in New Zealand, more information is required on biohel properties, treatments, process technologies, and the environmental issues associated with their 758
use. Knowledge of important fuel properties that affect ash formation and deposition can lead to the better design of combustion system to minimise fouling and slagging problems. T h s paper outlines ongoing work to evaluate the combustion performance of mainstream woody biomass fuels (forest arisings, wood processing residues, and short rotation forest crops) and fuel mixture of biomass/coal when compared to coal. The trials utilized a fully instrumented 5OkW laboratory facility configured to simulate a typical overfeed stoker combustion unit. The work undertaken involved detailed assessment of the effects of fuel characteristics on thermal output, flue gas composition (including the particulate contents and the levels of toxic trace inorganic constituents), ash deposition and the tendency for slag formation. MATERIALS AND METHODS
FUEL SAMPLES Four mainstream biofuels grown in New Zealand were used: RPB, bark from radiata pine (Pinus rudiutu ) from a sawmill; RPO, radiata pine offcuts with saw dust from a Veneer & Plywood factory; RPA, radiata pine forest arisings (harvesting residues); ENS, stem wood of Eucalyptus nitens (E. nitens, 7 yr), a short rotation forest (SRF) biomass crop. The other fuels used were: 5) RPB/coal, a blend fuel mixed in the ratios of 15:85 bark:coal (thermal input), equivalent to 19:81 bark:coal on a mass basis. The bark was from the same sawmill as above and the coal was a sub-bituminous coal (below); 6 ) Coal, a sub-bituminous coal (Kopako) which is considered to be representative of industrial coals used in and around the Central North Island region of New Zealand.
1) 2) 3) 4)
All biofuel samples were hogged in a 25 hp hogger to provide a particle size range of up to 35 mm. For the set of trials reported here, the fuels were not pre-dried in an attempt to simulate industrial situations. It is recognised that there are opportunities to reduce moisture contents to improve the overall heat output and general combustion performance. Characteristics of the sample fuels and their ash were determined according to the standard ASTM procedures presented in Tables 1 and 2 (5). Furnace gases passed over a convective tube bank containing a fouling probe to sample fouling deposits. Fly ash was removed by a cyclone before the gases were sampled and analysed for their composition using a gas chromatograph, and particulates and trace elements were also assessed (6). In all runs,a gas sample was taken isokinetically and passed through a quartz sampling train to a quartz disc filter to remove any remaining particulates. The sample then entered a series of bubblers containing solutions to trap targeted trace elements. The trace element concentrations in the impinger solutions were measured by ICP-MS. After exiting the sample line, the gas passed through a desiccant and on to a dry gas meter for volume determinations. The moisture content of the flue gas was determined from the increase in weight in the desiccant and impinger solutions and the concentrations of trace elements calculated in
759
micrograms per dry standard cubic meter of flue gas. Particulates were also calculated as milligrams per dry standard cubic meter of gas. Besides the particulates and trace elements, the flue gas was also monitored for oxygen, nitrogen, carbon dioxide, carbon monoxide, methane, and nitrous oxide by withdrawing gas sample from a t h d sampling port near the isokinetic line, passing it through a drying tower and into an on-line gas chromatograph. SOz and NOx were measured using an electrochemical Testoterm 350 analyser. After each run, samples of bottom, cyclone and fouling probe ash were recovered, weighed and analysed for major and trace inorganic constituents. The bottom ash was also sieved to determine the fraction passing through a 1.7 mm opening to serve as an indicator of the extent of slag formation.
Table I Fuel characteristics. RPB RPO Fuel type Proximate analysis, YOwet basis Moisture 38.9 36.7 Ash 3.5 0.3 Volatile 39.7 52.2 Fixed C 17.9 10.8 Calorific value, MJ/kg wet basis 12.98 12.84 Ultimate analysis, YOoven dry basis C 53.86 51.01 H 5.38 6.11 0 34.68 42.12 N 0.25 0.21 S 0.03 0.02
RPA
ENS
PRB/coal
Coal
52.9 3.2 36.8 8.8
51.4 0.7 39.1 8.8
22.7 4.7 37.6 35.0
24.2 5.1 34.3 36.5
8.99
8.37
19.92
21.06
49.41 5.98 37.81 0.1 1 0.01
49.59 5.82 42.96 0.21 0.02
64.28 5.02 23.54 0.89 0.19
67.16 4.78 19.87 1.23 0.26
Table 2 Ash characteristics. RPB RPO RPA Fuel Ash composition*, % in ash Si02 59.50 32.56 69.77 A1203 14.70 9.18 14.21 Fez03 3.30 13.55 3.05 CaO 7.30 16.96 3.3 1 MgO 1.90 4.84 1.09 NazO 3.40 3.96 3.88 K20 6.60 9.26 3.83 TiOa 0.40 1.13 0.35 Mn304 0.20 1.14 0.14 Ash fusion temperature, "C reducing condition Softening 1126 1070 1175 Spherical 1200 1150 1220 Hemispherical 1220 1170 1275 Fluid 1380 1190 1410
ENS
PREYcoal
Coal
7.30 5.00 8.20 36.30 8.50 3.80 19.70 0.35 1.90
24.14 19.70 5.03 28.78 3.46 1.72 1.25 3.01 0.07
21.10 20.50 5.40 30.60 3.70 2.10 0.30 3.60 0.10
1480 1490 1500 1510
1120 1140 1250 1240 1260 1260 1320 1300 *: Ash compositions do not add to 100% as some constituents e.g. SrO, BaO, SO3,B203etc. were
not measured.
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Particulate & Trace elements sampling port
To: Testoterm SOXBox
a
Flue gas cooler
Cyclone
m iL 2 - J Biomass
+- Combustion Chamber
Fig. I Combustion rig utilised for the combustion trials. RESULTS AND DISCUSSION
FUEL AND ASH CHARACTERISTICS Fuel analysis results presented in Table 1 provide a comparison between the biofuels and coal. On an as-fired basis, the four biofuels had significantly lower calorific values (8-13 MJkg) compared to coal (21 MJkg), partly attributable to the higher moisture content of the biofuels. Also, biofuels had lower ash and fixed carbon contents but relatively higher volatile matter contents than both coal and barWcoal blend. On an oven dry basis, the various biohels contained similar C, H and 0 contents, but overall a much lower carbon content than coal, reflecting their lower heating values. Again typically biofuels had lower nitrogen and sulfur contents than that of coal. The biofuel ash characteristics defined by the composition (major and minor components) and fusion temperatures were also different (Table 2). All radiata pine wood residue biofuels (RPB, RPO and RPA) showed higher SiOz contents but lower
76 1
CaO contents when compared to both E. nitens (ENS) and coal. E. nitens ash had the lowest SiOz and highest CaO content.
COMBUSTION CONDITIONS Although it was intended to maintain the same experimental conditions over the six combustion trials, all the materials exhibited different combustion conditions (Table 3). It is felt, however, the results allow for meaningful comparison of the feedstocks. The variation in firing conditions probably resulted from the variation in the nature of the feedstocks (Tables 1 & 2) and the results serve to indicate the likely performance of the fuels when fired in industrial applications. Excess air ratio was measured by the oxygen content of the flue gas. The excess portion of dry air volume is derived from the oxygen content of the flue gas and the stoichiometric oxygen demand. Of interest was the combustion rates (as-fired basis), which determine the actual thermal output. The relatively higher combustion rates of biofuels suggest that biofuels are more reactive than coal. Within the biofuel samples tested, higher combustion rates were observed for wood residues, predominantly of radiata pine origin than those for E. nitens, a hardwood. Thermal output was influenced by the fuel caloric value. When adjusting the thermal power for equivalent combustion rates, coal produced the highest thermal output of 41kW, while woody biomass had a maximum thermal power of only 29 kW for bark. This could be attributed to the higher moisture contents and lower carbon content of biofuels.
Table 3 Test firing conditions. Fuel Barometric pressure, kPa Fuel weight, kg Excess air ratio Total test time, min Temperature*, "C Fan inlet Cooler entry Cooler exit Stack Orifice Stack velocity, d s Combustion rate as fired basis, g/s Combustion rate as oven dry basis, g/s Thermal power, kW
RPO
RPB 100.7 58.30 1S O 249
102.8 65.41 1.45 28 1
RPA 101.6 71.43 1.95 270
ENS 102.2 68.50 2.01 324
PRB/coal 102.2 41.78 1.59 320
Coal 100.7 43.00 1.49 34 1
29 566 382 210 163 5.23 3.90
32 638 447 226 165 4.33 3.88
21 558 394 197 116 3.89 4.41
25 462 346 189 143 3.97 3.52
39 563 408 207 149 3.75 2.18
27 581 403 196 147 3.72 2.10
2.38
2.46
2.08
1.71
1.68
1.59
44.15
43.07
31.24
22.90
40.31
41.38
FL UE GAS
Flue gas production rate from coal was higher than that from biofuels, and the coal flue gas had lower moisture contents (Table 4), partly reflecting the low moisture content of the fuel (Table 1).
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Particulate emissions ranged from 9 mg per dry standard cubic meter (mg/dsm3) of the flue gas volume (corrected to 12% C 0 2 )for wood offcuts to 191 mg/dsm3 for SRF E. nitens. Following the lowest value of 9 mg/dsm3, coal combustion produced the second lowest particulate emission of 40 mg/dsm3. T h s variation was associated with the density of the fuels being fired, and the overall proportion of fines in the feedstock as fed into the combustor. However, wood offcuts with sawdust was an exception for its relatively low density and h g h proportion of fine particles.
Table 4 Combustion products, ash fouling and slagging tendencies. ENS Fuel RPB RPO RPA Flue gas, Nm3/kg - fuel as fired ~ r flue y 4.56 4.31 4.30 4.40 Moisture in flue 0.97 1.02 1.04 1.04 Total 5.53 4.33 5.34 5.44 Particulate emission Conc. @ 12% CO, 91 9 51 191 mg/dsm3 8.11 0.74 2.17 10.6 Rate, i# Ash, g Bottom >1.676mm 0.0 0.0 529 0.0 Bottom 4.676mm 1085 85 1214 495 Fly 1027 499 429 44 Deposit on fouling 0.0 0.4 2.4 0.9 Total filter 34.0 3.O 9.8 57 Total solids 2246 687 2284 697 Expected ash, 2065 225 2252 471 Unburnt C, YOmass In bottom ash 3 14 2 1 In fly ash 17 51 16 13 Ash fouling/slagging tendency, fraction of total ash, % Fly ash 47.9 84.9 19.6 7.4 Deposit on probe 0.0 0.1 0.1 0.2 Filter ash 1.6 0.5 0.4 9.5 79.8 82.9 50.6 14.5 Bottom ash 0 Bottom ash >1.676mm 0 30 0 100 100 70 100 Bottom ash < I .676mm
PRB/coal
Coal
7.57 1.07 8.64
8.76 0.91 9.67
75
40
4.65
2.62
1249 672 87 1.2 87 2196 1957
1398 747 92 1.7 0.0 2330 2179
4 22
3 14
4.2 0.1 4.2 91.6 65 35 _-
4.1 0.1 0.0 95.8 65 35
Complete combustion is obtained when flue gas analysis shows no carbon monoxide, hydrogen or methane and when the percentage of carbon dioxide is at a maximum. The presence of CO and CH4 in the flue gas from combustion of some biofuels indicates incomplete combustion suggesting lower process efficiencies. The exceptionally higher methane content from the combustion of forest arisings most ldcely resulted from the higher feedstock moisture content reducing the firebox temperatures (Table 5) (6). The firebox temperature also affects the formation of NO, As reported in the literature increasing firebox temperatures leads to increased concentrations of hydrogen and hydroxy radicals in the gas phase. Both of these species very effectively reduce N20 and N2 (7).
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The trials for forest arising (RPA) and E. nitens (ENS), with high excess air ratio, had the highest oxygen content in the flue gas. The flue gas composition from the different fuels was significantly different in SO2, CO and CH4 contents with biofuels having lower SO2 emissions but higher CO and CH4 emissions compared to that for coal. Biofbels combustion resulted in the generation of significantly lower SO2 compared to those from coal partly due to the low sulfur content of biomass compared to that of coal. In the case of the biofueVcoa1 blend, it can be expected that the lime (CaO), generated in the combustion of biohels, also reduces the SO2 emissions from the coal combustion according to the sulfur capture reaction below (6,8): CaO + SOz + 0.5 0 2 = CaS04 Table 5 Flue gas characteristics.
Fuel RPB RPO RPA ENS PFWcoal Coal Moisture, % wet basis 9.9 2.3 7.5 11.2 5.9 7.3 13.4" 13.8" 10.5" 10.6' 12.3b 13.1" COz, % dry basis 6.7' 10.5" 7.9cd 7.0b 02,% dry basis 7.2dC 10.2" Nz,% dry basis 79.4ab 79Sab 79.4ab 79.0b 79.8a 79.6ab 0" SO2,ppm dry basis 4.0' 0" 2.3" 86.7b 128" CO, ppm dry basis 879' 185' 4015" 2450b OC 0' CH4 ppm dry basis 4b Ob 439a 29b Ob Ob 1 7a 3b NzO, ppm dry basis 2b 17" 0" 0' 378" 76' nd NO, ppm dry basis nd 147b 0 NO2,ppm dry basis nd 3" 3" nd lb nd Note: (1) nd means not determined. (2) Values of each composition with the same alphabetical notation (a, b, c) are not significantly different. ASH DEPOSITION The results showed that a greater proportion of bottom ash was retained on the 1.7 mm screen for the coal trials (Table 4). The ash softening temperature of coal ashes (1120°C and 1140°C) was not higher than those for the woody biomass (Table 2). However, the heat output, reflected by combustion temperature, was higher for coal when adjusted to the same combustion rate with woody biomass. These three factors indicate that combustion of coal and barkkoal mixture both have a greater tendency to cause slagging than that for biohels. Biofuels have a lower tendency for slag formation. From a mineralogical viewpoint, combustion of different fuels will lead to different behaviour in slagging. Short rotation E nitens ash was rich in CaO (Table 2) which can react with alumina and silica to produce slag deposits. However, the low content of alumina and silica in the E. nitens ash and the possible low fire-box temperatures (less than the ash fusion temperatures) tended to preclude the formation of slag complexes. In contrast bark had higher contents of albite (Na20.A1203.6Si02)whch may melt around 1050" C to cause glassy deposits (6). Most sodium-silica alumina complexes do not melt at such low temperatures. Evidence from the combustion trials indicates that firebox temperatures were low enough to prevent slag formation, however, h s may not be so under other combustion conditions.
764
CONCLUSION Biomass fuels are lughly variable in their characteristics. Information is critical on fuel quality, combustion performance and ash characteristics to optimise combustion system design. In comparison with coal, combustion of biomass fuels and fuel mixture of biomass and coal has advantages with lower sulphur emissions and less tendency for slag formation.
ACKNOWLEDGMENT T h s project was funded by the Public Good Science Fund (PGSF), Foundation for Research, Science and Technology, Wellington, New Zealand.
REFERENCES 1. Ministry of Commerce (1997) New Zealand Energy Outlook, February. Energy Modelling and Statistics Unit, Energy and Resource Division, Ministry of Commerce, Wellington, New Zealand. 2. EECNCAE (1996) New and Emerging Renewable Energy Opportunities in New Zealand. Energy Efficiency and Conservation Authority and the Centre for Advanced Engineering, University of Canterbury, Christchurch, New Zealand. 3. Gifford J., Senelwa K., Hall P., Clemens A. and McGimpsey N. (1999) Forestry byproducts co-firing with coal: opportunities in New Zealand. Proceedings, New Zealand Coal Conference. “Coal - Powering into the 2 1St Century” 20-2 1 October 1999. Wellington, New Zealand. 4. Senelwa K., Gifford J., Hall P., Robertson K., Hooper G. and McGimpsey N. (In Press). Power production and energy related greenhouse gas mitigation options from forestry by-products in New Zealand. Biomass and Bioenergy. 5. ASTM, American Society for Testing and Materials (1992) Annual Book of Standards. Phdladelplua. 6. Dare P., Gifford J., Hooper R.J., Clemens A.H., Damiano L.F., Gong D. and Matheson T.W. (2000) Combustion performance of biomass residue and purpose grown species. Submitted to Biomass and Bioenergy. 7. Kilpinen P. and Hupa M. (1991) Homogenous N 2 0 chemistry at fluidised bed combustion conditions. A kinetic modelling study. Combustion and Flame 85: 95. 8. Someshwar A.V. and Jain A.K. (1993) Sulfur capture in combination bark boilers. TAPPI Journal 76: 179.
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Operating Parameters for the Circulating Fluidized Bed (CFB) Combustion of Biomass K. Smolders (l), D. Honsbein (2) and J. Baeyens (3) (I) University of Antwerp, Institutefor Environmental Studies, Universiteitsplein 1, 261 0 Wilrijk, Belgium. (2) Aston University, Bio-energy Research Group, Aston Triangle, Birmingham, England ( 3 ) University of Leuven, Dept. Chem. Engrg., de Croylaan 46, 3001 Heverlee, Belgium.
ABSTRACT The use of a CFB combustor and associated power generation could make small scale operations of biomass combustion technically and economically viable. In the 6rst section of this paper, a design strategy of the CFB for combustion of biomass is developed. The main operating parameters are the solids flux and gas velocity in the riser of the CFB. Isothermal conditions will be guaranteed by circulating inert sand particles of 300 pm at a solids flux of 20 kg/m2s.The gas velocity in the riser is determined at 7.9 m / s , exceeding the transport velocity (=6.6 m/s) of the inert sand bed by 20%. Besides the design of the CFB for biomass combustion, the paper will moreover determine the net combustion capacity (=4.1 MW,/m2 riser)of the CFB combustor. Other CFB applications as combustion of coal and sludge will also be illustrated. INTRODUCTION Growing concern about the global climate change and pollution is driving the development of renewable and sustainable energy sources. Biomass fuels meet this objective, but being a dispersed and diluted resource, it was for a long time considered as necessary to implement large projects to appreciate the significance and full value of biomass benefits. The circulating fluidized bed (CFB) combustion and associated power generation could make small scale operations technically and economically viable. The CFB has been applied for over 20 years in a variety of processes, mainly involving gadsolid reactions such as coal combustion, coal gasification, minerals’ calcination or roasting, etc. The combustion of biomass could yet be another success story for the CFB principle. If and whenever possible, desk design should define the critical operating parameters and preliminary layout.
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THE CIRCULATING FLUIDIZED BED
Industrial processes often involve the interaction of a gas and a solid phase. Low gas velocity processes are characterised by limited carry-over and are performed in packed, fluidized or slugging beds. When the gas velocity through the fluidized bed is increased above a critical velocity, entrainment becomes so important that the bed can no longer be maintained in the column unless entrained particles are collected and returned to the bed. Reactors with an external recycling system have a means to control the solids hold-up and are referred to as ‘Circulating Fluidized Beds’ (CFB). Figure 1 illustrates a CFB coal combustor.
CONVFCIIVE
B A G HOUSE
ASH
Figure I Circulating fluidized bed combustor. The main goal of the CFB design is the determination of the reactor (riser) dimensions to firstly fblfil requirements of conversion (and thus of residence time) and secondly to achieve the heat exchange rate, if so required. The recycle loop is of additional importance and its design can be based on correlations found in literature (e.g. Smolders et al. [11) and is not reported in this paper. THE DESIGN STRATEGY FOR-A CFB BIOMASS-COMBUSTOR
Figure 2 illustrates the successive steps in the design strategy of a CFB reactor, which will be applied for the design of a biomass-combustor with a capacity of 2.5 tonne*. DATA GATHERING Principles
The design procedure starts with gathering data concerning production capacity, particle and gas characteristics, reaction kinetics and heat exchange necessities.
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Kinetics reaction model reaction rate law specific reaction rate
needs of heat exchange?
DATA GATHERING
I Heat balance 4 riser r3 heat
temperature exchange area
Calculate transport velocity
gas velocity gas flow rate 4 riser diameter
4
4
Determine residence time for required conversion 4 riser height 4 4
recycle ratio solids circulation rate
I Final selection on economic grounds
Figure 2 Design sequence of the CFB riser. 768
Particle characteristics include the particle size distribution and average size, specific gravities and bulk density, sphericity and heat capacity. Important gas characteristics include gas density, viscosity and heat capacity, which are all temperature dependent. Moreover, the composition of the gas mixture is also an important parameter in reactor design. The reaction kinetics and thermodynamics can be defined by literature data or by experimental work.
Applied to the bio-mass combustor As is the case for bubbling fluidized bed combustors, CFB-combustors are operated with an inert (sand) bed. At start-up the CFB-combustor circulates sand (pp = 2600 kg/m3) , which is gradually supplemented and replaced by shale and combustion ash (pp = 2500 kg/m3). The bed characteristics should therefore relate to ash, rather than to sand. The average particle size is approximately 300 pm. Biomass is represented by the idealised formula of (C6HI2O5),,.The Combustion process is also idealised as: C,H,,O, + 6.5 0, + 6 CO, + 6 H,O with an exothermic reaction heat of 17.5 MJkg Calculations will be made for wood particles between 0.1-2 cm with a particle density of approximately 600 kg/m3. Wood particles contain 40 - 60 wt% moisture, 50 - 30 wt% volatiles, 1 wt% ash and approx. lOwt% fixed carbon (Boroduyla et al. [2]). Kinetics of the devolatilisation and combustion reactions are reported below.
THE HE4 T BALANCE DETEMZNES THE NET ENERGY YIELD AND FLOWRATES Principles A heat balance over the riser determines the net energy yield. At sufficiently high solids circulation rates the riser can be assumed to be isothermal (Baeyens [3]). Considerations of the reaction kinetics and heat exchange aspects result in an optimum temperature. Systems for exchange of heat have to be chosen and the exchange surface area has to be fixed for imbedded tubes or for a wall-mounted jacket. Once the operating temperature in the riser is known, the gas characteristics in the riser can be determined.
Applied to the bio-mass combustor The composition of the biomass (C6H1205),, determines the flow rate of both the needed air and the flue gases and the composition of the last one. Considering an excess air flow rate of 15%, 4.86 Nm3 airkg biomass is required, while the flow rate of the flue gases is 5.63 Nm3 flue gadkg biomass. The total flow rates for a 2.5 tph combustor are then 12,150 Nm3/hrof air and 14, I 0 0 Nm%r of flue gases. Biomass combustion occurs mostly at a temperature around 850°C. Heat is produced by the combustion reaction (17.5 MJkg) while part of it is consumed by heating the biomass and the air to the reactor temperature. A heat balance around the combustor results in the net energy of 1 1 MJkg biomass. Burning 2500 k g h biomass
769
thus correspondsto a 7.6 MWthcombustor. Heat can be released by imbedded tubes or by a wall-mounted jacket in the riser or in the recycle loop. THE TRANSPORT VELOCITYDETERMINES GAS VELOCITY AND RISER DHMETER Principles The transpot? velocity is an important design parameter since it determines the minimum gas velocity required to filly develop a stable fast fluidization mode in the riser. Smolders et al. [4] gathered literature data on the transport velocity of different particles ($=24 - 1000 pm, pp = 600 -3000 kg/m3)in CFBs of different size (D = 0.08 - 1 m) They compared these experimental data with the predictions by 8 correlations of several authors. They concluded that the transport velocity can best be calculated by the correlation of Adhez [5] for small particles: Re,
= 2.078Ar0.463
Ar<90
(1)
or by the correlation of Perales [6] for larger particles: Re,,
= 1.458Ar0.484
Ar> 90
(2)
To include a safety margin, the operating gas velocity will be chosen to exceed 1.2 Urn. As the minimum gas flow in the riser is determined by the reaction stoichiometry, the gas velocityand the riser diameter cannot be chosen independently.
Applied to the bio-mass combustor CFB-combustors for coal are operated with sand and ash of average particle size 300 pm (Anderson et al. [7]). In view of the small difference of density between both bed materials, transport velocities hardly differ. The biomass combustor will be operated with the same particle size. The transport velocity and operating gas velocity are then: Urn
urn,
= 6.56 m =
/~ 1.2urn=7.9m/s
(3) (4)
As the the flow rate of incoming air and that of the flue gases are very similar, the bed diameter will be calculated at the average of the two flow rates: 13,000 N m 3 h or 53,500 am% at 850°C. With known gas flow rate and required gas velocity, the riser cross-sectionalarea A is fixed at 1.88 m2 which results in a diameter of 1.55 m. THE RESIDENCE TIME DETERMINES RISER HEIGHT AND RECYCLE RA TI0 Principles The average residence time of the solids, required for the specified conversion can be estimated by applying the reaction rate law. In a riser of reasonable height, this
770
residence time cannot be achieved by a single pass of the solids and particles have to be recycled. Figure 3 illustrates the riser system with recycle.
MF: mass feed rate of solids o
M, = MI; + MR P, = PF + PR = (1+ N) PF
Figure 3 Solid flows in the circulating fluidized bed.
The recycling of solids causes a spread in the residence time, and hence in the conversion of the particles as some rise only once through the riser while others are recycled several times. Partly unreacted feed can hence be discharged, causing insufficient conversion. It can be demonstrated (Smolders et al. [8]) that the conversion in the product stream ( X A ), is equal to:
with Xqi the conversion of a particle that has a total residence time of i times the average residence time for one pass and with N the recycle ratio and equal to: N=
number of particles in recycle stream - PR number of particles in feed PF 77 1
If the converted particle is the main reaction product (e.g. for calcination reactions), the recycle ratio N and the height of the reactor determine the average conversion of the discharged particles Mp (and hence the product quality) as illustrated by eqn. (3). Smolders et al. [8] illustrated these principles for a CFB limestone calciner. In a combustor the mass of products withdrawn fiom the reactor Mp, represents only the small stream of valueless ‘ashes’. The characteristics of this stream are hence less critical in the design of the CFB combustor.
Applkd to the bio-mass combustor Kinetics of the biomass-combustion
In the incinerator, devolatilisation and combustion of volatiles and char take place. Release of vapour fiom the core cools the particle’s surfice and keeps its temperature low, so that it can be assumed that char combustion starts only after devolatilisation is completed. Boroduyla et al. [2] investigated the combustion of wood particles (dp = 3.5 to 21.5mm) in an fluidized bed of sand particles at temperatures fiom 1023 to 1 123 K The biofiel particle size was found to decrease insignificantly during drying and devolatilisation, with these processes mainly coinciding in time. Furthermore, the volatile extinction time of wood can be estimated fiom: t, = 12.1x lo3 x d y
(7)
To calculate the burnout time of char particles, a reaction model is required. With the assumptions that the reaction is mass transfer limited and that it occurs in two succesive steps with CO as intermediate, the burnout time t of a particle can be theoretically derived as (Smolders [9], Baeyens et al. [lo]): t* = P M 4
‘i
DO,
co,
Biomass combustion in the CFB
When wood particles are introduced in the CFB, they stay in the bottom part of the bed, considered as the dense bed, until their terminal velocity becomes small enough to cany them up the riser with the gas stream. The transport of large particles in the riser occurs not only due to the drag force of the gas but also due to the impact of the fine sand particles on the large wood particles. Geldart et al. [l 11 demonstrated that particles are elutriated when the gas velocity in the riser is larger than the terminal velocity calculated with an ‘effective’ gas density. This ‘effective’ density can be calculated as: Peff = P g
+P
(9)
This approach was developed using solids ranging in density fiom 706 to 5000 kg/m3and in mean size 60m 62 to 219 pm. To guarantee isothermal conditions in
772
the riser, the circulation rate is set at 20 kg/m2 s, corresponding to an ‘effective’ gas density of 3.8 kg/m3. The density of the wood particles entering the bed is 600 kg/m3 and will be reduced during devolatilisation. Since the average fixed carbon kaction is approximately 10 wt?!, the density of a completely devolatilised particle is 60 kg/m3. It is assumed that char combustion starts only after the end of devolatilisation. The terminal velocity at 850°C for different combinations of particle sues and densities are given in Table 1. Table I Terminal velocities d, (cm) pp(kg/m3) U, (m/s)
0.1 600 1.5
0.5 600 4.9
1 600 7.5
1.1 600 7.9
1.5 420 7.9
2 290 7.9
Particles up to 1.1 cm can be elutriated from the bed without any devolatilisation. Particles with a size between 1.1 - 2 cm will be partly devolatilised before leaving the bed. The required time for devolatilisation and the burnout time of the char particles can be calculated by eqn. (7), respectively eqn. (8). The concentration of O2 varies through the riser corn 1.03 mole/m3 after combustion of the volatiles to 0.296 mole/m3 in the flue gases. The concentration of O2 is assumed to vary linearly during the char combustion. Results are given in Table 2. Table 2 Theoretical reactions times d, (cm) tw (4 tch (S) ttotd(s)
0.1 1.51 2.26 3.77
0.5 12.3 56.5 68.7
1 30 226 256
1.1 34 273 307
1.5 51 508 559
2 74 904 978
For particles exceeding 1.1 cm,part of the devolatilisation takes place in the dense bed. Taking this into account, the following residence times in the riser are required for complete devolatilisation and char combustion: Table 3 Required reactions times in the riser d, (cm) tw (s) tch (s) ttotal (s)
0. I 1.51 2.26 3.77
0.5 12.3 56.5 68.7
1 30 226 256
1.1 34 273 307
1.5 34 508 542
2 32 904 936
If the velocity of the particles is estimated as the difference between the gas velocity (7.9 m / s ) and their terminal velocity at each moment, the height required for devolatilisation is equal to: Table 4 Required riser height for devolatilisation
4 (cm) H-h)
0.1 10.4
0.5 55
773
1
1.1
1.5
82
82
80
2 71
The most commonly encountered height for industrial CFB combustors is 20 m. For one pass, the fiaction of devolatilisation is given in Table 5. Table 5 Fraction of devolatilisation in the first pass
d, (cm)
0.1 0.5 1 1.1 1.5 2 100% 45% 48% 53% 69% 80%
%
devolatised in the first pass The recycle of wood is thus evidently needed. Since the suspension leaving the riser is quenched in the external steam generating bed, further devolatilisation will be low. Since combustion of volatiles is a fast reaction, volatiles will be completely burned in the outlet duct after the cyclone. After devolatilisation is completed, char starts to burn. Assuming the char starts to burn at the bottom of the riser, the fiaction of char combusted during the first pass in the riser is given in Table 6. Table 6 Fraction of char combusted in a single pass pass of 20 m
d, (cm)
0.1
0.5
1
1.1
1.5
2
100%
5.4%
1.5%
1.3%
0.8%
0.4%
%
combustion in single pass
The recycle of the char is thus evidently needed. Since the suspension leaving the riser is quenched in the external steam generating bed where 02-levels are very low, it is reasonable to accept that no combustion occurs in this recycle loop. In the CFB, char particles of all sizes are found. At equilibrium, the mass flow of char in the recycle loop varies from 375 kg/hr to 29,500 kgihr when feeding at a constant rate of 2500 kgihr of the specified particle size. Results are detailed in the Table 7. Table 7 Char combustion in the riser
d, (cm) total height required for char combustion (m)
0.1 0.5 17.4 384
numberofpassesbefore complete burn-out (+ number of passes for devolatilisation)
1 19 70 (+I) (+3) (+4)
total of unburned char in recycle loop (kgihr)
375 2750 9250 10750 18375 29500
774
1 1.1 1390 1645
82 (+4)
1.5 2865
2 4695
143 (+4)
235 (+4)
The circulation rate of the inert bed is set at 20 kg/m2 s or 135,000 k g h . The fraction of unconverted char in the recycle loop is hence max. 22 w% (when feeding 2 m m wood particles). Only a small stream leaves the combustor as solid i.e. the ashes corresponding to their weight fiaction of the biomass (i.e. 1% for wood). This means that 25 kg/hr ash or max. 5.5 kg/hr unburned char (=.2 % of the fed char) is lost in the ash stream thus stressing that this loss should not be accounted for in the heat balance and associated calculations. COMPARISON WITH OTHER COMBUSTION APPLICATIONS
Coal combustion is one of the major applications for CFB reactors. Quoted capacities range from 1 MW* (labscale) to 700 MW* (Soprolif-EDF). It is announced by EDF that capacities will exceed 500-600 MW, after the year 2000. The proposed design procedure was used to design a 12 MW& CFB coalcombustor [8]. Comparison of relevant predicted data and typical commercial data demonstrated the validity of the design procedure. Sludge has significantly higher contents of nitrogen, volatile matter, ash and (for wet sludge) moisture than typical coals and has furthermore a very low fixed carbon content. It has been reported that 11,25 and 55% of the sludge produced respectively in the European Union, USA and Japan is incinerated. Because of the limitations currently fiicing the other sludge disposal outlets (landfill and use as fertiliser), it is expected that incineration will gain importance in the next decade. Because of the high moisture content of most sludges, autothermal incineration is only possible if the combustion air is preheated by the flue gases. Sludge combustors are hence mainly built for sludge disposal, rather than for energy production purposes. The proposed procedure was applied for the design of a CFB combustor for a sludge flow of 2 tonnes DS& and 4 tonnes waterh [8]. It was assumed that 50% of the solids were organics with an overall composition of CsH702Nand a heat content of 22.1 MJkg organic DS. Table 8 illustrates the predicted data for the combustion of biomass, coal and sludge. The particle size of the inert bed, solids circulation rate and gas velocity are the same for the 3 combustors. For the same thermal capacity, the coal and biomass CFBcombustors are comparable in size. The specific combustion load for sludge combustors is lower due to the large amount of water vapour.
775
Table 8 Comparison of CFB combustors for biomass, coal and sludge.
biomass
coal
sludge
2500 kg /hr
2200 kg /hr
6000 kg /hr (DS 33%)
1.88 mz 20 m
2.56 mz 20 m
1.63 mz 20m
particle size of bed material gas velocity solids circulation rate
300 pm 7.9 m / s 20 kg/mzs
300 pm 7.9 m / s 20 kg/m2s
300 pm 7.9 m / s 20 kg/m2s
combustion capacity net capacity
12.2 MWh 7.6 MWth
17.7 MWh 12 MWh
6.1 MWth
feeding rate surface area of riser height of riser
specific combustion load specific load (net)
6.5 MWt,,/mz 6.9 MWdmZ 3.8MW,dm2 4.1 MWth/rn2 4.7 MWth/m2
CONCLUSIONS A design strategy for a CFB combustor for biomass was presented and illustrated for a 2500 kg/hr combustor. The main operating parameters are the solids flux and gas velocity in the riser of the CFB. Isothermal conditions will be guaranteed by circulating inert sand particles of 300 pm at a solids flux of 20 kg/m2s. The gas velocity in the riser is determined at 7.9 m/s, exceeding the transport velocity (=6.6 m/s) of the inert sand bed by 20%. About 2% of the char will be lost in the ash stream. The net combustion capacity of the riser is equal to 4.1 MWh per mz of the riser, comparable with a coal CFB com bustor.
776
REFERENCES 1.
2.
3. 4. 5.
6 7 8 9 10
11
Smolders K. & Baeyens J. (1995) The operation of L-valves to control standpipe flow. Adv. Powder Technology,6,163- 176. Boroduyla V.A., Dikalenko V.I., Palchonok G.I. & Stanchits L.K. (1995) Fluidized bed combustion of solid biological waste and low-grade fuels: experiment and modelling. In: Fluidization V N , Oyland, Florida (U.S.A.), May 1995. Baeyens J. (1998) Heat transfer in fluidized beds, CPA-course, Amsterdam, November 1998. Smolders K. & Baeyens J. (2000) Gas fluidized beds at high gas velocities: a critical review of occurring regimes, submitted to Powder Technology. Adhnez J, de Diego L.F. & Gayan P. (1993) Transport velocities of coal and sand particles. Powder Technology,77,61-68. Perales J.F., Coll T., Llop M.F., Puigjaner L., Amaldos J. & Casal J. (1991) On the transition fkom bubbling to fast fluidization regimes. In: Circulating Fluidized Bed Technology ttt,(Ed. by P. Basu, M. Horio & M. Hasatani), 4- 1- 1, Pergamon Press. Andersson B.-A. & Leckner B. (1992) Experimental methods of estimating heat transfer in circulating fluidized beds. tnt. J. Heat/Mms Transfir, 35,3353-3362. Smolders K. & Baeyens J. (1999) Design strategy for a gadsolid Circulating Fluidized Bed Reactor. Powder Handling & Processing, 11,257 -264. Smolders K. (1 999) Hydro&namics in Circulating Fluidized Beds. Ph.Dissertation, KUleuven, Belgium. Baeyens J. & Geldart D. (1978) Fluidized bed incineration: a design approach for complete combustion of hydrocarbons. In: Fluidization, (Ed. by J.F. Davidson & D.L. Keams, 264-269, Cambridge University Press. Geladart D., Cullinan J., Georghiades S., Gilvray D. & Pope D.J. (1979) The effect of fines on entrainment fiom gas fluidized beds. Trans. tnst. Chem. Eng. 57,269275.
777
NOMENCLATURE
riser surfbce area Archimedes number
(= pg 6,- pg )gd; 1 p’)
concentration of O2 diffusion coefficient for 0 2 particle size gravitational constant height required for devolatilisation mass flow rate of product withdrawal recycle ratio number of particles in the feed number of particles in the riser radius of coal particle Reynolds number for the transport velocity (=pe Urn 4 /p) burn-out time for char total bum-out time of wood particles volatiles extinction time superficial gas velocity in the riser transport velocity conversion average conversion gas viscosity bulk density of fines in the riser ‘effective’ gas density gas density molar density of carbon particle density
778
Agglomeration and The Content of Amorphous Material in FB Combustion. A Full-scale Boiler Test J. H. A. Daavitsainen, L. H. Nuutinen, M. S . Tiainen, and R. S. Laitinen, Department of Chemistry, P. 0. BOX 3000, FN-90014 University of Oulu, Finland
ABSTRACT: The agglomeration tendency of the bed material in FB boilers can be studied with both XRD and SEM-EDS. XRD gives information on the identity of the crystalline phases as well as on the content of amorphous material in the sample. SEMEDS can be utilized to reveal the compositional distribution of the particles and their coating layers. The combination of these two techniques facilitates the better prediction of the bed behaviour during the combustion. The relative amount of amorphous material is connected to the deposition and agglomeration behaviour during combustion. The present paper discusses two boiler tests, where sodiurn-richplywood was combusted in quartz-free bed material GR GRANULE. The unused bed was hlly amorphous. After initial increase of the crystallinity of the bed material the content of amorphous material started again to grow as the tests proceeded. The content of amorphous material was generally hgher in the test I than in the test I1 and the SEMEDS studies showed the presence of small agglomerates in the bed of test I. All agglomerates were found to be formed around quartz impurities that were probably left in the boiler as a residue from a previous silica sand bed. The presence of a non-crystalline or a microcrystalline phase in a sample produces a broad hump (s.c. amorphous halo) in the diffraction pattern. The area under the halo is proportional to the content of the amorphous material. The complex ash samples often produce several partially overlapping halos in the diffraction pattern. This, and the presence of diffraction peaks due 9 crystalline phases have presented a challenge for the integration. The matrix effects need also be taken into account in the calibration. In this work we report results from a full-scale FB test. Selected samples were analysed with XRD and results were compared to the SEM-EDS data of the same samples. It was found that the content of amorphous material in the bed increased as the test proceeded.
779
INTRODUCTION
Agglomeration during fluidised bed combustion is a difficult problem, which can lead to defluidisation of the bed and to unscheduled shutdown of the boiler. In agglomeration, bed particles stick to each other with adhesive material. The adhesive material is a result of chemical reactions of ash during combustion. The mechanisms leading to agglomeration are assumed to include viscous flow sintering, reactive liquid sintering, and chemical reaction sintering.1-6 Glass formation may be connected to these sintering mechanisms. The latest studies show that alkali metal-aluminosilicates play an essential role in agglomeration, because they form the adhesive material of the agglomerate^.^-^ In order to predict ash-related problems it is important to be able to estimate the content of harmful substances in the process. With scanning electron microscope and energy dispersive X-ray spectrometer (SEM-EDS), the chemical composition of the problematic material can be determined, but it does not give much dormation about crystallinity. The determination of amorphous content in a sample has been difficult due to lack of suitable methods. Powder X-ray diffraction (XRD) offers a possibility to determine the content of amorphous material based on a character in a diffraction pattern called amorphous halo. The properties of the halo depend on the amount and chemical composition of the amorphous material present in a sample. In this paper we have studied changes in the amount of amorphous material content in bed samples.
EXPERIMENTAL
Two different tests were carried out with quartz free bed material using fuel with high alkali metal content. Plywood and sawdust were co-combusted during the test I in different mixing ratios in the 6 MW boiler. In test I1 with 5 MW boiler only plywood was used as fuel. The tests are described more detailed elsewhere.l0> The bed material samples were collected from bottom ash collector of the boiler. The selected samples were studied with SEM-EDS to determinate the chemical compositions. XRD was used to study the amount of amorphous material present in the samples. For the SEM-EDS-analysis samples were mounted with resin. Mounts were crosssectioned, polished and coated with thin carbon layer. The samples were characterised by using a Jeol JSM-6400 scanning electron microscope combined with a Link ISIS energy dispersive X-ray analyser. The acceleration voltage of 15 kV and a beam current of 120*10-* A were used for the SEM-EDS-analysis. The sample distance was 15 mm. The magnification used for bed samples was x50. Utilizing automated image processing, ca. 1000 domains were analysed for sodium, magnesium, aluminium, phosphorus, silicon, sulfur, potassium, calcium, titanium, and iron contents.* The image processing was performed with IMQuant software incorporated in Link ISIS. The SEM-EDS-results were visualised as quasiternary diagrams using a Konpad software package that is locally designed especially for this p u r p ~ s e . ~ For XRD measurements the samples were ground in a WC/Co -mortar for one minute and homogenized manually in an agate mortar. X-ray diffractograms were collected with the Siemens D5000 diffractometer with 28 goniometer. The MoKa molybdenum anode X-ray tube was used at 30 kV and 40 mA with zirconium absorber.
780
Step size was 0,02" and the counting time was 0.5 seconds per each step. Diffraction patterns were measured in 28 range 5"-60". The crystalline phases have been identified by the use of the database of ICDD.12 The amorphous material induces a broad hump or halo, in to the diffraction pattern. The width and the position of the halo indicate the distribution of interatomic distances in the structure of the material. The area under halo depends on the content of amorphous material, and therefore it is possible to semiquantitatively determine the content of amorphous material from the x-ray diffraction pattern.l3? l4 Mixtures of synthetic ash and common glass were used as calibration samples.
RESULTS AND DISCUSSION The main crystalline phases were identified (see Figs. 1 & 2) and the contents of amorphous material were determined from selected bed samples by XRD. Main crystalline phases are Ca9(Al6Ol8) calcium aluminium oxide, Ca8,2sNal,5(Al&8) calcium sodium aluminium oxide, calcium magnesium silicate (diopside), calcium phosphlde, and magnesium silicate (enstatite). It is clear that the relative contents of the crystalline phases change as the tests proceed. The'amount of first three phases initially increases in both tests, whle amount of latter two phases decreases. The decrease is less significant in test I1 than in test I. It corresponds well to the more pronounced growth of Ca-Mg- and Ca-P-containing coating layers on the bed particle surfaces that was observed with SEM in the case of test 11. Also, the increase of the content of the three first phases, especially diopside, can be connected to the growth of the Ca-rich coating layers. The composition of these coating layers is discussed below. The powder diffraction data show that the unused bed was almost 100% amorphous [see Figs. l(a) & 2(a)]. The content of amorphous material decreased rapidly at the early stages during both tests, which is probably due to the initial heating of the bed material effecting the crystallization. As the tests proceeded, the content of amorphous material started to increase again. In test I the increase was 54-67% and in test I1 2750%. The test I clearly indicated higher amorphous content through whole test period than test 11. This can be explained as follows: The bed material in test I is known to contain silica sand as impurity, probably due to the residue from the earlier bed in the boiler. It has previously been fbserved that the agglomeration takes place in the vicinity of the quartz particles. It is likely that the higher content of amorphous adhesive material is connected with the presence of small agglomerates formed around the quartz impurities. It is also possible, however, that the increase of the amount of amorphous material correlates with the growth of the coating layers on the particle surfaces. Agglomerates were not found in test 11, which explains the generally lower content of amorphous material. The amorphous content in test 11, however, also increases together with the growth of the coating layers indicating that the coating is at least partly amorphous.
781
Fig. 1 Diffractionpatterns of the bed samples collected during the test I: (a) unused bed, (b) day 7, (c) day 8, (d) day 15, and (e) day 18. The dotted line illustrates the amorphous halo. The crystalline phases have been identified by the use of the database of ICDD.
782
I
120
....
61,
*norphouh*
100
ao
2001
00
150
40
100
I
E. YgSIO,
50
20
I 0
10
20
30
I0
20
90
40 2 0 (de(l)
W
60
W
60
0
10
30
20
40
KI
60
300 250
200
'
.
150 100
'
50.
4a
Fig.2 Diffraction patterns of the bed samples collected during the test 11: (a) unused bed, (b) day 17, (c) day 20 and, (d) day 3 1, The dotted line illustrates the amorphous halo. The crystalline phases have been identified by the use of the database of ICDD.
100
100
I
1
80 60
%
Ye
40
20 0 0
7
8
15
18
0
Test day
17
20
31
Test Day
Fig. 3 The content of amorphous material in bed samples of a) test I and b) test I1
The unused bed is quite homogenous, as seen in Fig. 4 for test I. As the test proceeds, two main trends in the compositional distributions are apparent. The distribution spreads out towards the Al+Si/ Na+P+S+K-edge on one hand, and to the Mg+Ca+Ti+Fe-corner on the other. The first trend is connected to the formation of sodium-rich adhesive material in small agglomerates formed around quartz impurities 783
in the l o The second trend indicates the formation of several Ca- and Mg-rich coating structures on the surfaces of bed particles. These coating structures are discussed in more detail elsewhere.
Na+P+S+K
Na+P+S+K
Fig. 4 Time dependence in the compositional distributions of bed samples from test I. The compositional distribution in the adhesive material of the agglomerates is shown in Fig. 5, and its average elemental composition is illustrated in Fig. 6 . The selection in the diagram corresponds to medium grey area in the backscattered electron (BE) image of the sample (see Fig. 5). Two types of particles are visible in the BE image. Those appearing as lighter grey are quartz-free particles that form the main component in GR GRANULE. The darker grey core is quartz. The compositional distribution of the agglomerate in the vicinity of the quartz particle is also indicated in Fig. 5. Its average composition is 39% of silicon, 20% of calcium, 13% of sodium, 12% of aluminium, and 9% of potassium (see Fig. 6 ) . The amorphous phases that are probably alkali metal silicate glasses are found in these regions. It is well established that the adhesion of the bed particles in agglomerates involves the molten alkali metal-silicate phase 3, that may well be amorphous.
784
Fig. 5 The compositional distribution of the adhesive material of the agglomerate is shown in the quasiternary diagram and correlated with BE image.
1
Na Mg Al Si P
S
K Ca Ti Fe 1
Fig. 6 The average chemical composition of the adhesive material of the agglomerate in the vicinity of a quartz particle (see Fig. 5).
The samples from test I1 show one of the two trends observed for test I. The compositional distribution spreads out towards the Mg+Ca+Ti+Fe-corner due to the formation of coating structures, as discussed above (see Fig. 7). The chemical compositions of different coating layers that have been determined by the use of point analyses are illustrated in Fig. 8. The comparison of Figs. 7 and 8 shows that the content of calcium- and phosphorus-rich material increases as test proceeds. The presence of t h s type of phases is also indicated by the XRD results. It has previously been deduced that this type of coating layers protects the bed from agglomeration. 785
Indeed, test I1 where the bed contained no quartz impurities did not indicate a tendency towards agglomeration, and consequently no shift of the compositional distribution towards the Al+Si/Na+P+S+K-edgewas detected.
i
N8+P+S+K
Na+P+S+K
AI+SI
Al
Na+P+S+K
Fig. 7 The time dependence in the compositional distributions of the bed samples from test 11.
g+Ca+ +Fe
Na+P+S+K Fig. 8 Point analyses from coating layers as illustrated in a quasiternary diagram.
CONCLUSIONS T h ~ swork is a part of a systematic investigation with a main objective to establish a connection between the content of amorphous material in the bed and the agglomeration tendency during the FB combustion. The present paper discusses two
786
boiler tests, where sodium-rich plywood was combusted in quartz-free bed material GR GRANULE. The presence of amorphous material can be observed in the X-ray diffraction pattern as a halo. The area under the halo can be used to estimate the ratio of amorphous and crystalline phases. The unused bed was fully amorphous. During the combustion the crystallinity of the bed material initially increased, but as the tests preceded, the content of amorphous material started again to increase. The content of amorphous material was generally higher in the test I, which indicates a higher agglomeration tendency. Indeed, SEMEDS studies showed that there were small agglomerates in the bed samples taken during test I. All agglomerates were found to be formed around quartz impurities that were probably left in the boiler as a residue from a previous silica sand bed. The agglomeration tendency can be studied with both XRD and SEM-EDS. XRD gives information on the identity of the crystalline phases as well as on the content of amorphous material in the sample. SEM-EDS can be utilized to reveal the compositional distribution of the particles and their coating layers. The combination of these two techniques facilitates the better prediction of the bed behaviour during the combustion.
ACKNOWLEDGEMENTS Financial support from Technology Development Centre Finland, Putkimaa Oy and Vapo Oy, and programming support from Yrjo Leino at Centre of Scientific Computing are gratefully acknowledged.
REFERENCES 1. Nordin A,, et al. (1996) Mechanism of Defluidisation and Bed Sintering during FBC of an Agricultural Biomass Fuel. In: Applications of Advanced Technology to Ashrelated Problems in Boilers, (Ed. by B. L. and D. R.) 353-366. Plenum Press, New York. 2. Lin W., et al. (1997) Agglomeration phenomena in fluidized bed combustion of straw. In: Proceedings in 14th International Conference on Fluidized Bed Combustion, Vancouver, Canada, vii-xvi . 3. Ergudenler A., et al. (1993) Agglomeration of silica sand in fluidized bed gasifier operating on wheat straw, Biomass & Bioenergy 4, 135-147. 4. Manzoori A. R., et al. (1992) The fate of organically bound inorganic elements and sodium chloride during fluidizdh bed combustion of high sodium, hlgh sulphur low rank coals, Fuel 71, 513-522. 5 . Manzoori A. R., et al. (1993) The role of inorganic matter in coal in the formation of agglomerates in circulating fluid bed combustors, Fuel 72, 1069-1075. 6 . Skrifiars B.-J., et al. (1997) Mechanisms of bed material agglomeration in the cyclone and the return leg of a petroleum coke fired circulating fluidized bed boiler. In: Proceedings in 14th International Conference on Fluidized Bed Combustion, Vancouver, Canada, 8 19-829 , 7. Nuutinen L., et al. (2000) Role of quartz sand in the agglomeration during the FBcombustion using fuel with high sodium content. In: Proceedings in Effects of coal
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quality on power plant pe#ormance: Ash problems, Management, Solutions, 08 May, Park City, Utah, . 8. Virtanen M. E., et al. (1999) A Novel Approach to Use CCSEM when Studying Agglomeration in Fluidized Bed Combustion. In: Impact of Mineral impurities in Solid Fuel Combustion, (Ed. by R. P. Gupta, Wall, T.F., and Baxter, L.L.) 147-154. Kluwer Academic / Plenum Publishers, New York. 9. Virtanen M. E., et al. (2000) SEM-EDS image analysis in the characterisation of coatings and adhesive material in quartz bed. In: Proceedings in 5th International Conference on industrial furnaces and boilers, 11 April, Porto, Portugal, . 10. Nuutinen L., et al. (2000) An improved bed material for the BFB-boilers. Case 1.: Co-combustion of sawdust and plywood waste. In: Proceedings in 5th European conference on industrialfurnaces and boilers, 11 April, Porto, Portugal, . 11. Laitinen R., et al. (2000) An improved bed material for the BFB-boilers. Case 2.: Combustion of fuel with high sodium content. In: Proceedings in 5th European conference on industrialfurnaces and boilers, 1 1 April, Porto, Portugal, . 12. ICDD (1999) Powder diffraction file 13. Tiainen M. S., et al. (1999) Determination of Amorphous Material in Peat Ash by X-ray Diffraction. In: Impact of mineral impurities in solid fule combustion, (Ed. by R. P. Gupta, T. F. Wall and L. Baxter) 217-224. Kluwer Academic Plenum Publisher, New York. 14. Nakamura T., et al. (1989) Determination of amorphous phase in quartz porwder by X-Ray powder diffiactometry, Powder DiJjFction 4, 9- 13. 15. Dawson M. R., et al. (1992) Bed material cohesion and loss fluidization during fluidized bed combustion of midwestern coal, Fuel 71,585-592. 16. Lin W., et al. (1999) Agglomeration in fluidized bed combustion of biomass mechanisms and co-firing with coal. In: Proceedings in 25th International Conference on Fluidized Bed Combustion,May 16-19, Savannah, Georgia, 12 . 17. Skrifvars B. J., et al. (1997) Ash Behavior in a CFB Boiler during Combustion of Salix, Energy and Fuels 11,843-848. 18. Skrifvars B. J., et al. (1998) Characterization of the sintering tendency of ten biomass ashes in FBC conditions by a laboratory test and by phase equilibrium calculations, Fuel Processing TechnologV 56, 55-67. 19. Skrifvars B. J., et al. (1999) Predicting Bed Agglomeration Tendencies for Biomass Fuels Fired in FBC Boilers: A Comparison of Three Different Prediction Methods, Energy and Fuels 13,359-363. 20. Valmari T., et al. (1999) Field Study on Ash Behavior during Circulating FluidizedBed Combustion of Biomass. 1. Ash Formation, Energy and Fuels 13,379-389. 21. Visser H. J. M., et al. (2000) Biomass ash - bed material interactions leading to agglomeration in fluidised bed combustion and gasification. In: Proceedings inpjlh international conference on thermochemical biomass conversion, Tyrol, Austria, . 22. Werther J., et al. (2000) Combustion of agricultural residues, Progress in Energy and Combustion Science 26, 1-27. 23. Zevenhoven-Onderwater M., et al. (2000) The prediction of behaviour of ashes from five different solid fuels in fluidised bed combustion, Fuel 79, 1353-1361.
788
Co-Combustion of Different Waste Wood Species With Lignite in an Industrial Steam Boiler with a Moving Stoker Firing System P. Grammelis, P. Vourliotis and E. Kakaras Laborato y of Steam Boilers and Thermal Plants, Mechanical Engineering Department, National Technical University of Athens, 9 Heroon Polytechniou Ave., 15780 Zografou, Greece
ABSTRACT: The concept of co-firing coal with waste material is the most promising option to contribute to the thermal exploitation of significant waste quantities. The utilization of this fuel mixture in the existing combustion systems may have serious effects on the fuel preparation, the combustion behaviour, the emissions and the ash composition. This paper describes the results from the co-combustion application in a 13.8 MWm steam boiler, located in North-western Greece. Several trials with various waste wood species and Greek ligmte were performed in the combustion facility, which comprises a moving grate system and a multi-fuel burner. The waste wood species were uncontaminated wood, Medium Density Fibreboard (MDF) and power poles. Before the co-combustion trials, a theoretical estimation of the ash deposition tendency using several indices was accomplished. During the experimental measurements, O2 concentration in the flue gas and the emissions of CO, SO2 and NO were continuously monitored. The operation data of the boiler were recorded and ash samples from the combustion chamber and the multi-cyclone were collected. The results showed that co-combustion of waste wood with Greek lignite is technically feasible in boilers with a moving stoker firing system. Additionally, MDF use instead of uncontaminated wood in the fuel blend resulted in improved combustion efficiency. The burning of a MDFIuncontaminated waste woodlignite blend for a six-month period indicated that no additional maintenance costs of the mechanical equipment of the boiler or costly waste gas scrubbing are needed. INTRODUCTION ‘Waste wood’ is every wooden product that has been used at least for once as commercial product and remains now as waste for disposal or recyclingheuse (/lo.It mainly consists of residues from the wood processing industry, demolition and construction wood, old or destroyed fkmture, boxes and pallets, wood material included in municipal solid waste, old power poles and railway sleepers. Each type of waste wood has been treated with different mechanical processes and chemical
789
substances. Thus, its disposal or recycling practices present different risks for the human health and the environment. Since the waste wood potential remains significant in Europe (/2/, /3/) and, thus, its disposal is a growing problem, attention has been focused, in recent years, on the material and energy recovery from this waste material. In this way, waste wood can be thermally recycled either in an incinerator or in industrial boilers substituting part of fossil fuels. The latter option is considered to be an economically attractive alternative for the energy exploitation of great waste wood quantities, offering, in parallel, several rewards, such as the environmental benefits, i.e. reduced net COZ emissions, and possible financial gains from the minimisation of the waste disposal and reduction of the dependence on fuel imports. Within the above mentioned concept, the feasibility of thermally exploiting waste wood in combination with solid fossil fuels was investigated and the results from the co-combustion of different waste wood species with Greek lignite in an industrial steam boiler are presented in this paper. The installation belongs to the Medium Density Fibreboard (MDF) producing industry PINDOS SA, located in north-western Greece. The main goal of these trials was to prove that already existing moving stoker firing systems mostly in Southeast Europe, can thermally recycle a fraction of waste wood at a substantial percentage together with locally available solid fuels. The objectives of this work were (a) to investigate the operation of an industrial scale boiler during co-combustion of waste wood and lignite of low quality, (b) to determine the CO, SO2, NO, PCDD/F and heavy metal emissions, (c) to compare them with emissions from co-combustion with uncontaminated wood, and (d) to correlate gas emissions with the fuel blend properties. BOILER CONFIGURATION The steam boiler has total thermal capacity 13.8 MW*, with steam production being 16 todh at nominal pressure of 21 bar. It comprises a two-compartment combustion chamber, an economiser a multi-cyclone and an ash removal system (Figure 1). The first compartment employs a moving grate, which consists of seventeen series of stairs, each of them being alternatively constant or moving. The moving stoker system operates with sawdust and wood chips and covers about 50 - 60% of the total thermal input. Primary air is distributed through four ducts below the moving grate and it is also used as the cooling medium of the stairs, while the secondary air is fed above the fuel supply. The flame position is arranged to cover about 1/3 to 2/3 of the moving grate surface and there is enough area for the fuel drying and the ash removal, at the beginning and the end of the grate, respectively. A multi-fuel burner is placed at the top of the second compartment, which can operate with heavy oil, light oil, sawdust or light oil in combination with sawdust. The ignition of the burner is realised with light oil and air is regulated automatically, according to the used fuel. Beyond the increase of the produced thermal load, the role of this burner is also to improve the combustion efficiency of the installation.
790
CO,NO,SOI
+to rtncli
Emissions recording : 4,6 : 1,2,4,6 Temperature measurement Heavy metals, dioxins and furans sampling :
Differential pressure Ash sampling 6
: :
1-6
3,5
Fig. 1 Measuring sites at the moving stoker boiler of “PINDOS”
FLTEL TEST MATRIX AND METHODOLOGY OF MEASUREMENTS Greek lignite from the Ptolemais reserve and various waste wood species, i.e. uncontaminated wood, MDF residues and power poles were used in the cocombustion experiments. Demolition wood was not included in the examined waste wood types, due to the small percentage of wood that is contained in demolition waste - less than 10 % of the total demolition material - and the difficult and uneconomic separation of wood from the rest of the materials, m a d y the metal objects. As power poles are not considered to be conventional waste wood and their burning needs special requirements (/4/), they were added only in 20 (%wt) in the fuel blend. The fuel test matrix that was elaborated for the evaluation of the co-firing behaviour is shown in Table 1. Prior to testing, a theoretical estimation of ash deposition tendency was obtained, using appropriate indices. Following the tests, a demonstration of the co-combustion operation mode with the optimum fuel blend was performed. During the co-combustion tests and the demonstration phase, a multi-component gas analyser (MULTOR 610) was used to monitor the oxygen concentration and CO, SO2 and NO emissions in the flue gas. Oxygen measurement was based on the paramagnetic principle, while the emitted pollutants were determined through an NDIR measuring system. The recorded values as well as the operation data of the boiler were retrieved through a P/C. Additionally, ash samples were collected from the first compartment of the combustion chamber and the base of the multi-cyclone
79 1
and were analysed for their unbumt and heavy metals content. Gaseous samples were collected from the flue gas duct after the multi-cyclone unit and before the air draft fan, in order to investigate the presence of PCDDiF, particulates and heavy metals in the flue gases. The results of these measurements and the heavy metals concentration in the ash samples are presented elsewhere (154.All the measuring sites at the boiler are well depicted in Figure 1. Table 1. Fuel test matrix of the co-combustion trials at the moving stoker. Fuel blend Symbol (a) Uncontaminated wood - Lignite Uncontaminated wood - Lignite (b) Uncontaminated wood (c) (4 MDF MDF - Lignite (el Uncontaminated wood - Lignite (f) Power poles (g) Uncontaminated wood - Lignite MDF MDF - Lignite - Power poles (h)
(% weight) 80 120
(% thermal input) 88.4 I 11.6
60 I40 100 100 80 I 20 60 I 20 I 20
74.1 125.9 100 100 90.7 19.3 64.1 I 11.2 I 24.7
60 120 120
62.6 I 10.9 126.5
60 120 120
69.4 19.5 121.1
FUEL AND ASH CHEMICAL ANALYSES
The raw material characterisation for each of the fuel samples, including the proximate, ultimate and ash analyses, as well as the calorific value measurement was accomplished according to the ASTM standard methods. As shown from the results depicted in Table 2, the main difference is detected in the moisture content of the fuels, as Ptolemais lignite has the highest value of 58.8 (%wt), while the waste wood species and especially MDF residues have much lower moisture contents. This fact seriously influences the combustibles content and the calorific value of the fuels. Combustibles are almost 82 (%wt, as received) for the uncontaminated wood and over 90 (%wt, as received) for the contaminated wood species, rising up to 99 (“hwt)in a dry basis and indicating the good quality of these fuels. As a result, all samples appear to have similar net calorific values, varying from 14081 to 18716 kJ/kg, in as received basis. In relation to the biomass samples, Ptolemais lignite has significantly lower content of combustibles, around 25 (% wt, as received) and, thus, much lower net calorific value, equal to 6314 kJkg, in as received basis. All biomass samples have very h g h volatile content, while lignite has by far the lowest volatile content and a comparatively high fixed carbon concentration. Moreover, the waste wood samples are characterized by their .low ash content that varies from 0.5 to 1.1 (%wt, as received), which are typical values for biomass samples. The sample of Ptolemais lignite has a much higher ash content of 5.97 (%wt), which is rather low for this reserve and it is attributed to the high quality of the selected lignite. According to the ultimate analyses (Table 2) and compared to the lignite, waste wood has significantly lower sulphur content in a dry basis, being negligible in the case of uncontaminated wood. MDF concentrates the greater amount of nitrogen, whch is one order of magnitude greater than the other two waste wood species. Also, biomass appears to have significantly higher oxygen content, which is expected to accelerate the combustion process. 792
The ash chemical analyses of lignite and waste wood species in their oxides are given in Table 3, as these were used for the calculation of the ash deposition indices.
Table 2.Proximate and ultimate analyses (%wt) and calorific values of the fuels used during the co-combustion trials, at the moving stoker boiler. Fuel
Ptolemais Uncontaminated Lignite wood PROXIMATE ANALYSIS [%wt, as received] 17.30 58.80 Moisture 75.69 Volatiles 25.01 Fixed Carbon 10.22 6.29 81.97 Combustible 35.23 0.73 Ash 5.97 ULTIMATE ANALYSIS [%wt, dry] Carbon, C 47.78 42.42 Hydrogen, H 4.54 5.33 Sulfur, s 0.94 0.0 0.18 Nitrogen, N 1.44 51.19 Oxygen, 0 * 30.80 CALORIFIC VALUE [as received] 8160 Gross Calorific 15475 Value, Ha [kJkg] 63 14 Net Calorific 14081 Value, Hu [Wkg] * Estimated by difference
MDF
Power Poles
6.60 90.80 2.05 92.86 0.54
8.60 82.63 7.68 90.30 1.10
45.07 5.92 0.25 3.16 45.02
50.02 6.14 0.27 0.18 42.19
18519
20160
17142
18716
Table 3. Ash chemical anabses (!?!?wt)of the fuels.
Ash analysis % Ptolemais weight Lignite Si02 32.08 A1203 9.30 Fez03 7.18 MgO 6.67 0.36 K20 Na20 0.00 CaO 40.00 p2os 0.64 2.06 so3 Rest 1.71
Uncontaminated wood 14.45 2.71 1.61 8.00 10.04 0.17 5 1.30 2.82
3.01 1.59 0.00 10.00 1.69 4.50 63.50 4.50
Power Poles 13.43 2.50 4.50 5.60 1.75 0.74 57.45 0.88
8.90
11.21
13.15
MDF
CaO was found to have the hghest concentration in all the samples, exceeding 50(%wt) for the waste wood and S O 2 was also intensively present mainly in the
lignite. Increased percentages of alkaline oxides were detected in the waste wood samples, rising up to 10.04 (%wt) of KzO in uncontaminated wood and 4.5 (“hwt)of Na20 in MDF. The total amount of alkaline oxides did not exceed 1 (“hwt)in the lignite ash. This distribution of the alkali metals and the lower percentages of silica
793
and alumina compounds are anticipated to reduce the ash softening temperature of waste wood (161). QUALITATIVE EVALUATION OF THE ASH DEPOSITION TENDENCY
According to the ash chemical analyses for the fuels used in the co-combustion trials (Table 2), various indices (/ti/)were calculated, in order to qualitatively determine the ash tendency for the deposit formation. The expressions of the most important ash deposition indices are: Base / Acid ratio of ash components (B/A): Fe, 0,+CaO+MgO+K ,O+Na 0 ( B l A)= All 0,+SiO, +TiO, Total Alkalis (TA):
(2)
TA=Na,O+K,O
The results are shown in Table 4. Although two or more fuels may have the same (BIA) ratio, their major constituents mamly determine the slagging characteristics. In the most cases, the base/acid ratio in the range of 0.4 to 0.7 reflects a potential for the ash containing metals to combine in the combustion process for the production of low melting salts and, thus, higher slagging potential. The opposite occurs for the extreme values of the ratio at either end. The values of all the ash samples are higher than the upper value of 0.7, indicating low slagging tendency. However, the alkaline metals, sodium and potassium, can form combinations with low fusibility temperatures, altering significantly the base/acid influences. The increase of the total alkalis implies low-melting ash and, consequently, the combustion of waste wood is expected to provoke more serious ash deposition problems than the lignite burning. Table 4 Ash deposition indices of the fuels used in the co-combustion trials. Ash deposition indices Lignite Uncontaminated MDF Power poles Base I Acid Ratio Total Alkalis
1.310
wood 4.145
0.360
10.210
4.397 2.490
17.324 6.190
EVALUATION OF THE RESULTS FROM THE CO-COMBUSTION TRIALS OPERATING CONDITIONS OF THE BOILER In each co-combustion trial, waste wood was mixed with lignite and the blend was unloaded in the fuel tank. The mixture was supplied on the moving grate through a spinning wheel, a chain conveyor and a fuel damper. Two radioelements detecting the fuel level in the intermediate reception hopper and on the moving grate, respectively, were used for the appropriate regulations of the fuel supply. Emitted pollutants and steam generation were monitored and when the boiler operation was under steady state conditions, the experimental measurements started. The fuel supply in the 1'' compartment of combustion chamber and sawdust feeding in the burner were continuously measured. The thermal input was kept almost constant in all the experiments and equal to 4 MWm for the moving grate and 6 MWh for the burner. Among others, the temperature in the first and second compartments of combustion
794
chamber, before the multi-cyclone and at the flue gas exit, as well as the differential pressure between the first compartment and the chimney were recorded. The results of all these measuring parameters are shown in Table 5. No significant variations were detected for the different test cases and only a slight increase of the temperature in the first compartment was observed, when MDF was used as base fuel. The improvement of the combustion evolution, the faster achievement of constant operation conditions and the increase of the useful heat output were evident when MDF was burned, either alone or in conjunction with other fuels.
Table 5 Operation data of the boiler, as measured during the co-combustion trials.
Fuel
To ("C)
(a)
19 12.5 13 17.5 15 14.5 16 21
(b) (c)
(d)
(el (f) (g)
Ti ("C) 725 700 710 720 720 700 690 730
T2 ("C)
T3("C)
AP (mbar)
640 590 640 640 640 650 650 640
188 183 198 193 190 192 188 189
3.2 4.5 3.7 4.1 3.5 4.8 3.9 3.6
(h) Where: To: Ambient temperature T,: Temperature in the 1'' compartment of combustion chamber Tz: Temperature in the 2"d compartment of combustion chamber T3:Temperature of flue gases AF? Differential pressure between the first compartment and the chimney
As far as the multi-fuel burner in the second compartment is concerned, measurements were obtained when firing only sawdust and without any fuel supply on the grate. The purpose of t h ~ stest was to estimate the burner's influence on the experimental results. The emissions recorded are given in Table 6 and with reference the results of the co-combustion tests (see Fig. 2), NO emission is increased, because of the significantly higher excess air ratio. Table 6 Emissions during the test with only the multi-&el burner operation using sawdust. co SO2 NO mg/m3~ dry,6% 0 2 311 54 785
EMISSIONS AND COMBUSTION EFFICIENCY The evaluation of the co-combustion Fehaviour was based on the emissions of CO, NO and SOz and the unburnt fuel content of the ash samples. The emissions of CO, NO and SO2 measured at the stack, were converted at [mg/Nm3,dry, 6% OZ],and are shown in Figure 2, as a function of the excess air ratio. The corresponding results for the unburnt fuel content of the ash samples collected from the first compartment of the combustion chamber and the cyclone are shown in Figure 3. CO emission was maintained at similar levels during the combustion of the different fuel blends. The increase of the lignite percentage in the fuel blend (b) up to 40 (%wt) did not seriously influence the CO emission (Figure 2), while the unburnt fuel content was decreased, in the same test case (Figure 3). T h ~ scould be attributed 795
to the higher residence time required for lignite combustion on the grate, due to its high moisture content. Moreover, comparing the CO emission values and unburnt fuel content for the blends (a) and (e), it was proved that the use of MDF instead of uncontaminated wood in the fuel blend brings about a slight improvement of the combustion efticiency (Figure 2). The same result raised from the comparison of the multi-he1 blends (0, (8) and (h), when MDF was used at the highest percentage of 60 (%wt).
The burnout in the ash samples of all the fuel blends was high enough and greater than 97 (%wt) for the ash samples in the lst combustion chamber, indicating the increased combustion efficiency. As it was found in all similar cases (17/), the maximum values of the unburnt fuel content were detected in the fly ash particles of the cyclone. This is mainly caused by the faster pyrolysis of the waste wood particles that certainly affects their residence time in the combustion chamber. Also, the high excess air ratio coming from the multi-he1 burner operation seriously effected the combustion efficiency, as the sample collected when ftring only sawdust and without any fuel supply on the grate was found to have the highest unburnt fuel content of all the other samples, equal to 6.06%wt. Generally, NO emissions are dependent upon the fuel nitrogen content, but in all test cases, NO were mostly influenced by the operation conditions and especially the excess air ratio. This observation is valid even when MDF (fuel blend (d)) was burnt, which had the highest nitrogen content ( / 5 4 The highest values of NO were observed in the fuel blends (c) and (g), where the excess air ratio was at its maximum (Figure 2). The lowest NO values were obtained for the fuel blends (a) and (0, following the trend of the excess air variation. Furthermore, low SO2 emissions were measured in the co-combustion trials (Figure 2). The highest SOz value was observed during combustion of the mixture with the highest lignite content (b), due to the high sulphur content of raw lignite.
DEMONSTRATION CO-COMBUSTION MODE Based on the results of the experimental measurements, a long-term operation mode with the co-combustion of the fuel blend of MDF / Lignite / Uncontaminated wood in the proportion of 60 / 20 I 20 (%weight) was carried out on site. Power poles were not included in the optimum fuel blend mainly due to the increased operating and maintenance cost of the wood shredding system existing at PINDOS SA. The emissions of CO, SO2 and NO were recorded in intermediate intervals and the mean values converted at [mg/Nm3,dry,6% 04 were 249,40 and 600, respectively. These are below the legislative limits and do not deviate much fiom the values of emissions attained during the co-combustion trials with the various fuel blends, leading to the conclusion that in case of the systematical co-combustion performance, there will be no need for costly waste gas scrubbing. Additionally, the monitoring of the boiler operation and the inspection of the heat exchange surfaces and the water tubes of the economizer did not show any significant changes in the ash deposit formation. Consequently, no additional maintenance costs of the mechanical equipment of the boiler will be obtained if the burning of this fuel blend is chosen as permanent solution for the installation.
796
800
Fi
-
-
0 700--
s
-
600
: m
0
__--
---
- 0 -
.a-.-. . ,A. o--.-.=. * -. ,. ' '.
\
c3.
-
-
v
--2.25
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--
I
g
a- - - - -$ - - - + . .o.. .
~
200-1
.
O
/--&.
,
0- loo-.
u
r
,
0,' ~
'
,
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l
.-.=/-'
,
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h - - - -& -
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".o . . . - - . - . o - . - - - - - - o - - . .O........O.~...... ....~~ ,
z .-0
-
500:; E 400
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2.50
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.
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l
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Fig. 3 Unburnt fuel content of the ash samples that were collected from the first compartment of combustion chamber and the cyclone, respectively, during the cocombustion tests. CONCLUSIONS The assessment of the results from the experimental trials and the demonstration operation mode showed that co-firing of waste wood species with lignite is technically feasible in a boiler with a moving grate furnace. Especially, MDF use instead of uncontaminated wood, in combination with lignite, brought about a slight improvement on the emissions and the combustion efficiency, under the same operating conditions. Also, better burnout was achieved during the lignite cocombustion with waste wood, as the residence time of lignite on the moving grate was higher. However, one should always keep in mind that the utilisation of Greek lignite at high percentages in the fuel blend leads to reduction of the usefbl heat output,
797
because of its hgh moisture content. In all co-combustion trials, PCDD/F emissions and metal elements both in the flue gases and the solid residues were lower than the legislative limits (154. Taking into consideration all the above mentioned and the Greek Energy policy’s trend to promote the development of decentralised energy production units supplied with biomass and various solid waste types, it is concluded that waste wood is a promising option for industrial and district heating boilers. This is particularly true for wood processing industries, such as PINDOS SA. Further development of waste wood thermal exploitation will be accelerated, if solid waste management companies, which will collect and transform the wood waste into an easy to handle by the boiler operators’ form, will be established (/8/). ACKNOWLEDGEMENTS
The financial support by the European Union (Thermie Action A, Contract No SF/0261//97) is gratefully acknowledged. REFERENCES
1. Marutzky R. (1997), Energiegewinnung aus Rest- und Gebrauchtholz, “Moderne Feuerungstechnik zur energetischen Venvertung von Holz und Holzaffallen Springer-VDI-Verlag,pp. 1-21. 2. Meyrahn H., (2000), Low emissions co-combustion of different waste wood species and lignite derived products in industrial power plants, THERMIE Type A-Action, Final Report to EC DG-TREN, pp. 9-18. 3. Nussbaumer, T., (1996), Anlagetechnik und Wirtschafilichhit der thermischer Venvertung von Altholz, BWK ,vo1 48, Nr 11/12, NovemberKIezernber,pp. 6165. 4. Verordnung uber Verbrennungsanlagen fiir Abfalle und M i c h e brennbare Stoffe - 17. BimSchV. 5 . Kakaras E., P. Vourliotis, P. Grammelis, G. Sakellaropoulos, P. Samaras, G. Skodras (1999), Co-combustion of biomasdwaste wood /lignite blends in an industrial steam boiler, THERMIE type A Workshop on ” Co-combustion of Waste Wood, Biomass and Lignite in Industrial Boilers - PossibilitiesPerspectives”, Thessaloniki, GREECE, November 1999, pp. 89-97. 6. Singer G. (Editor). (1981), Combustion fossil Power Systems, 3‘‘ edn., Combustion Engineering Inc., Windsor, pp. 3-1 to 3-33. 7. Meschgbiz A,, Kakaras E., Krumbeck M. (1995), Combined Combustion of Biomass and Brown Coal in a Pulverized Fuel and Fluidized Bed Combustion Plant, 31d European Conference on Industrial Furnaces and Boilers, LisbodPortugal. 8. Kakaras E., Pavloudakis F., Karlopoulos E. (26-29 June 2000), The use of wood waste for energyproduction - Technical aspects and management issues, Renewable Technologies for Sustainable Development, Madeird Portugal. ’I,
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Biomass and Waste-to-Energy Conversion in the Netherlands by Means of (1n)direct CoCombustion: Status, projects and future applications in the Dutch Utility sector A.J.A. Konings, R. Meijer, C .M. Rozendaal, W.J.A. Ruijgrok, R. de Vries KEMA Power Generation & Sustainables, P.O. Box 9035, 6800 ET Arnhem, The Netherlands
ABSTRACT Production of energy from biomass and waste streams is a major topic for the Dutch electricity production sector. The main driving forces for these activities are the C02 emission reduction targets, tax incentives and new legislation with respect to waste disposal imposed by the Dutch government. In the Netherlands a goal to produce 120 PJ of electricity from sustainable sources in the year 2020 is set by the government. This corresponds to 10% of the total energy production in the Netherlands, of which 50 PJ is designated to come from (in)direct co-combustion and stand-alone bioenergy plants. Direct and indirect co-combustion of biomass and waste streams with coal are considered costeffective alternatives to achieve this target. The Netherlands have imposed fierce regulations on biomadwaste-to-energy regarding emissions into the air. It should also be guaranteed that the by-products of coal combustion, such as fly ash, bottom ash and gypsum, have to maintain their quality to ensure a 100% utilisation and application in road construction or in the building industry. In this paper an overview of the recent activities by the Dutch utilities will be presented, together with a view on the technically and economically most viable options to be applied in the Netherlands in the near future. INTRODUCTION The conversion of biomass and waste into electricity and heat is increasingly performed by the coal-fired power plants in the Netherlands. The main objective is the reduction of COz emissions from the combustion of fossil fuels, and subsequently the reduction of the contribution of coal-fired power plants to the potential hazard of global warming. Since the signing of the Kyoto-protocol in 1997, the Netherlands is obliged to reduce the greenhouse gas emissions by 6% in 2008 to 2012 compared to the level 799
of 1990. For ratification of this agreement by the Dutch parliament,a set of conditions should be fulfilled, such as the ratification by other countries. Secondly, the introduction of energy taxes on a European level and the possibility to use flexible instruments as Joint Implementation and emissions trading should be stipulated. One of the means to achieve the challenging task of meeting the Kyoto commitments is the increase of the contribution of the renewable energy sources from 5 per cent in 2005 to 10 per cent in 2010. This translates into replacing 270 PJ/a of fossil energy production of which 75 PJ/a must come from biomass. Table 1 shows that in 2020 a total of 50 PJ/a is foreseen to be delivered fiom either (in)direct cocombustion in large-scale power plants or in small-scale stand alone electricity production. The total target of 120 PJ requires 4 million tons of dry biomass per annum. Of this roughly 50% is foreseen to be delivered from contractible biomass and waste streams in the Netherlands, 10% from energy crops and 40% from imported biofuels.
Table I
Scenario for energy from biomass and waste (in PJ/a) (Novem, 1998a).
Technology Waste incineration Households Industry (1n)direct co-combustion Small-scale stand alone Landfill gas/ digestion Other Total
1995 14.5 8 5 0.1
2000 30 8 5 3
2007 40 8 18
2020 45 8 5 20
5
2 6
6
30
5
8
32.6
54
85
8 4 120
DUTCH CLIMATE FOR BIOMASS AND WASTE-TO-ENERGY CONCEPTS Energy from biomass is hardly commercially viable for any technological option at the current price levels of competing conventional alternatives in the Netherlands. Representative price levels for the production and supply of these alternatives are presented in Table 2. Due to the current process of liberalisation of both the electricity and the gas markets, prices are expected to drop substantially. Regarding electricity the decrease may lie in the order of 20 - 25%, whdst the gas wholesale prices may become 10 - 15% lower. Therefore, utilisation of renewable electricity from biomass is stimulated by financial and fiscal incentives to achieve the Dutch C02-targets. These measures should be placed in the above-mentioned framework of the Dutch policy to achieve an improvement of 10% energy supply from renewables in 2010.
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Table 2 Price levels for energy production in the Netherlands. Price level Electricity production Electricity spot market Heat production Natural gas (large consumers)
65 - 100
26 - 65 4-8 0.20 - 0.45
NLG / MWh (including fuel tax) NLG / MWh (import not taxed for fuel) NLG / GJ (including fuel tax) NLG / m3 (including fuel tax and Environmental levy)
Stimulation of the use of renewable energy by the Dutch authorities focuses on four aspects: Stimulation of demand of renewable energy by proper marketing of renewable electricity and introduction and reinforcement of taxes. Stimulation of production of renewable energy through solving political and administrative barriers and establishing emissions standards for biomass conversion. At present the various emission legislation acts and the emission legislation imposed by local authorities result in confusion and not in a level playing field. Development of a trading system for green labels. Green labels are granted for the production of a certain amount of renewable electricity. Import of renewable energy.
FISCAL MEASURES AND FINANCIAL D M N G FORCES To support the introduction of renewables the Dutch government has established the following fiscal measures, which aim at narrowing the price gap between "conventional" energy and renewable energy: A tax on the use of electricity and natural gas, which is largely paid by small and average consumers. An exemption of this energy tax when consumers buy renewable energy; in this way the extra cost of renewable pricing is lowered for the consumer. A support rate for producers of renewable energy, which is paid fiom the revenues of the energy tax. A scheme for corporate tax reduction based on an instant depreciation of the entire investment and a special deduction to a value of 40% of the investments. For energetic utilisation of thermal technologies a criterion of 40% energy efficiency is applied. An income tax exemption scheme for green funds that create the possibility to obtain loans for investments at a rate that is on average 0.5% lower than standard commercial loans. Of the present fiscal measures the energy tax, the exemption for renewable energy and the production support rate for renewables offer the strongest and most important incentives for narrowing the price gap with conventional energy. Especially the increases the Dutch government will introduce in 2000 and 2001 on the energy tax on electricity and gas may push the consumer to choose for 80 1
renewable energy. At the tax levels proposed the price difference between conventional and renewable electricity will vanish. Besides the arrangements for financial support of renewables, there are policy plans to introduce additional measures to stimulate the use of renewable sources. These plans include: 0
0
A tradable certificate system for renewable energy, which may come into operation in 200 1. A free, fully liberalised market for selling renewable electricity to all consumers by 200 1.
BIOMASS AND WASTE STREAMS Thermal conversion of biomasslwaste streams imposes demands on the fuel quality. The basic physical properties of the biomass and waste streams, such as moisture content, ash content and melting temperature, particle size (distribution), density and calorific value are important properties, which determine the design specifications of a new installation to a large extent. The fact that biomass and waste streams are usually ill defined, leads to a significant spread in physical and chemical data. The amount of biomasslwaste that is contractible for energy recovery is extremely sensitive to surrounding economical and political influences and is in most cases only a fraction of the potential amount. Some streams show a high potential, but due to the use in other markets, re-use applications or a very dispersed origin, the availability may be very low. Moreover, quality demands, changes in quality and the market value of available biomass/waste streams further reduce the amount that can be contracted for energy recovery. Tables 3a and 3b speclfy the amounts of several waste and biomass streams in the Netherlands. They show that fiom a potential of about 23,000 ktoda, representing an energy production of 313 PJIa at the moment, only approximately 20%, i.e. 17,000 G W a may be regarded as contractible. Table 3a Biomass and waste streams in the Netherlands (Novem, 9916; AOONVAV, 1998).
Stream
Potential Available Contractible kton PJIa kton PJla PJla GWa db/a dbla Wood 3,112 62 1,320 24 19 5,278 Agricultural 1,256 18 730 12 3 833 Manure 9,122 96 547 6 2 556 Chicken litter (7,632) (77) Cattlelpig litter ( 1,490) ( 19) Sludge 955 19 319 6 5 1,389 Waste from FDI 7,65 1 101 383 9 3 833 Waste from VFG 654 10 0 0 0 Other 373 7 282 5 2 556 Total 23,123 313 3,581 62 34 9,444 FDI = Food and Drinks Industry; VFG = Vegetable, Fruit and Garden 802
Table 36 Several waste streams in the Netherlands (GAVE, 1999). Stream Industrial waste Shredder waste Papedplastic rejects Residues demolition waste Total
Available kton db/a 2,000 170 150 700 3,020
PJ/a 26 3 3 12 44
Contractible PJ/a 22 3 2 ? 27
GWa 6,111 833 556 ? 7,500
In addition to availability the most important factor for using a biomass/waste stream for energy recovery is purely economical. Studies have shown that the most critical and influential parameter in the profitability of a biomasdwaste-to-energy system is the (negative) he1 price. Due to the rapid growing market and the relatively small amounts of contractible biomass in the Netherlands, this is a very competitive issue and subsequently little mformation is available. In the current Dutch economic situation (interest, depreciation time, maintenance & operating costs etc.) the llnk between the level of specific investment costs and maximum allowable fuel costs can be made (table 4).
Table 4 Maximum allowable fuel price as a function of the specific investment costs for thermal wood processing on a 20 to 30 MWe scale. Specific investment costs NLG/kWe 1,000 3,000 5,000
Maximum allowable fuel price NLG/GJ 4 -1 -7
NLG/ton 80 -20 -140
NLGMWh 14.4 -3.6 -25.2
CO-COMBUSTION IN THE NETHERLANDS The production of renewable energy in the electricity sector should be considered in relation to the privatisation of the electricity production and distribution companies and the liberalisation of the electricity market in the Netherlands and for a large part in Europe. Despite this uncertain future a large number of initiatives and projects have been developed to increqse the production of renewable electricity and heat in the coal-fired power plants id the Netherlands. The objectives of the initiator can be divided into two categories: business economics and environmental aspects. Co-combustion should not result in an increased unavailability or a reduction of the quality aspects of the by-products, such as fly ash, that could reduce the value or even the utilisation pqssibilities of the byproducts and lead to lower electricity production costs. The environmental aspects often regarded are the reduction of C 0 2 emissions from fossil hels, the reduction of waste disposal and compliance with emission standards and constraints. During the 80s the Dutch environmental policy became more and more integrated. In addition to the legislation of environmental compartments of bottom,
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water and air the general environmental law was finally (90s) introduced, covering all the different separate compartments. Simultaneously the preferred sequence of waste management was introduced by the so-called "Ladder of Lansink" (1979). This implies first of all waste prevention followed by applications of (material) reuse. Concerning the remaining waste streams energy recovery should have priority over landfill. According to the most recent National Environmental Policy Plan ("3) the use of waste as a fuel is also recognised as an option of re-use. The material options of re-use are still given priority to energetic options of re-use. Thls priority order ensures that co-combustion of waste is only possible in case the waste cannot be re-used or when formation cannot be avoided. Each plan for cocombustion of waste in coal-fired power plants is therefore tested on its order in priority. An Environmental Impact Study is often mandated or desired by the national administration and/or regional authorities, who discuss and compare the possibilities for reutilization and alternative processing routes (for example by means of a Life Cycle Analysis) (KEMA, 1999). Table 5 summarises the thermal conversion techniques, applied in the Netherlands, classified according to stand-alone applications as well as to integrated applications. In addition, the techniques for which the prospects are considered highly promising for application within the next 10 years, are also mentioned. Traditional thermal techniques for stand-alone waste combustion in the Netherlands can be divided into grate firing technologies and fluidised bed technologies. Direct co-combustion involves either the direct mixing of coal with biomass upstream the pulverisers or a separate grinding and transportation to supply the biomass to separate biomass burners. Indirect co-combustion involves either biomass combustors or gasifiers that are integrated with the coal-fired plant. Compared with direct co-combustion, indirect co-combustion offers the advantage of separated ash removal. In the Netherlands only few stand-alone bubbling fluidised bed combustion applications are in operation to process specific wood, sludge and industrial waste streams. The wood- fired bio-energy-plant at Cuyk represents the current state of the art of BFBC, handling fresh wood, saw residues, park-/garden residues with steam parameters of 100 bar and 520°C. Stand-alone FBC, like the one at Cuyk, offers approximately 30% (net) electrical efficiency. This value does not offer the energetic advantage (40% electrical efficiency) co-combustion does. The Dutch government stimulates the introduction of new waste-to-energy technologies. The objective is to realise a substantial improvement of the electrical efficiency compared to grate-fired Municipal Solid Waste Incineration (MSWI), processing contaminated types of waste. New technologies should at least offer 30% electrical efficiency, processing wastes from both municipal and industrial sources. If improved concepts of circulating FBC can come up to the supposed expectations, the Dutch market of energy recovery from specific waste streams can certainly be boosted.
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Table 5 Applied thermal techmques including hlghly promising techmques withm the next 10 years (2000 - 2010). Type of technology Stand-alone
Thermal technique Grate firing
Fluidised bed combustion Co-combustion
Direct Indirect New promising Pyrolysis technologies Hydro T h e m 1 Upgrading Pyrovac Co-combustion Co-gasification
Application Type of Biomass I Waste Municipal waste; RDF Industrial waste Biomass Industrial waste (monostreams) Biomass; RDF Coal mixed with biomass Biomass Wet Biomass Biomass Grate (in PC-boiler) Entrained flow gasification of coal and biomass streams
Direct Co-Cornbustion Experience with energy recovery from specific monostreams of biomasslwaste, in the Netherlands concentrates on direct co-combustion in coal-fired power stations (Beekes et al., 1998). In the year 2000 the Dutch coal-fired power plants will be responsible for utilising electricity from 900 ktodannum biomass/waste. At present different routes for direct co-combustion in coal-fired power stations are in operation or in preparation (figure 1).
[ COAL
SECONDARY FUEL@)
BURNERS
Flue gas
Physical pre-treatment pulverising / drying pelletising I mixing
Fig. I Proven and reviewed types of direct co-combustion in the Netherlands. The total installed capacity of coal-fired installations in the Netherlands is about 4,000 W e , divided over 7 units ranging from 400 MWe to 645 MWe. The main characteristics of the coal-fired units are listed in table 6 . The main fuel for all 805
units is imported bituminous coal. For strategic and commercial reasons the coals are mostly burned in blends, in such a way that the quality of the blend guarantees a smooth and economic operation of the plant. However, the dynamics of the coal market, the vast amount of suppliers and sources and the intention to buy cheap coals result in sometimes substantial variations in quantity as well as quality of the feedstock delivered to the power plants. Power plant operators are therefore used to frequently changing coal qualities. In the Netherlands a 250 MWe coal-gasification plant is located at Buggenum. Th~spressurised entrained flow oxygen-blown gasifier, equipped with extensive cold-gas cleaning, is exceptionally well suited for the co-gasification of relatively contaminated biomasdwaste streams. Several biomass and waste streams have been reviewed regarding suitability and feasibility. Currently preparations are underway for a 111-scale demonstration of co-gasification of chicken manure.
OVERVIEW OF DIRECT CO-COMBUSTIONACTIVITIES
In the coal-fired power stations there are a number of co-combustion projects, some of which are in an engineering phase (see table 6). More detailed information on the individual projects is given in short in the following sections. Table 6 Current status of the main co-combustion projects in the Netherlands.
Gelderland
Capacity MWe 602
Borssele 12
420
EPZ
h e r8 Amer 9
645
600
EPZ EPZ
Maasvlakte 1
520
EZH
Maasvlakte 2 Bupqenum
520
275
EZH Demkolec
Power plant
." 13
Operator EPON
Co-combustion fuel Waste wood Phosphor oven gas Waste wood Paper sludge cocoa shells Paper sludge Waste wood (gasifier) Liquid organic waste Biomass pellets Chicken rnanure
Status since 1995 since 1996 since 1998 since 1998 since 1998 since 1997 Commissioning since 1995 since 1998 Demonstration
POWER PLANT GELDERLAND 13 The main co-combustion activity engaged at the power plant Gelderland 13 (CG13) in Nijmegen is the co-combustion of waste wood (originating from demolition activities). The waste wood (chips) is delivered to the plant, stored in large bins on a specifically designed waste wood storage yard, pulverised in a Pallmann hammer mill and four Fuller MicroMills, and pneumatically transported with cold air to the wood burners. Two wood burners are situated in either sidewall of the fiunace below the main pulverised coal burners. The installation has been in operation since 806
1995 and can handle about 60,000 tons of waste wood (substitution of 45,000 tons of coal and reduction of 110,000 tons of COz emissions) in its current state. At present EPON is studying the feeding of wood particles directly into the pulverised coal transport lines after the pulverisers, since the injection with the separate wood burners was subject to some problems, such as the occurrence of large pieces of unburned wood in the bottom ash. The first experiences with this method of feeding observed no negative implications. Recently, the power plant applied for a permit to expand the wood co-combustion activities from 3 to 10 per cent of the coal input, to be done with the direct feeding of pulverised wood into the pulverised coal transport lines. In addition, EPON was granted a permit from the authorities to conduct a number of full-scale test trials with the co-combustion of biomass, such as waste of sunflowers, olive and apricot kernels, and coffee ground. In thls permit it is foreseen to conduct five tests with 200 tons of biomass and five tests with 50 tons of biomass. When the experiments have proved successful the co-combustion of the biomass fuels will also be addressed in the application of the revised permit for the power plant as a whole.
BORSSELE POWER PLANT UNIT 12 The phosphor production plant located near the power station of Borssele, produces phosphor oven gas (mainly consisting of CO and Hz), whch was partially used on t hs industrial site and partly flared. By creating the possibility of co-combustion in the coal-fired power plant a reduction in atmospheric emissions of COz, SO2 and NO, was obtained for this industrial region. The gas cannot be stored and must therefore be used at the power station the moment it cannot be applied at the phosphor production plant. In practice, this means that co-combustion m a d y takes place during nighttime, when the unit is operated at half load. A concern at the start of the project was whether the quality of the fly ash for concrete production could be influenced by substantial amounts of phosphate. The residual amounts of phosphor in the delivered gas however, turned out to be so low that no adverse effects were observed. Borssele power plant also has the intention to co-combust 120 ktons per annum of biomass. Examples of the types of biomass have already been listed in table 6 (sewage sludge and cocoa shells), but paper sludge, clean and untreated waste wood, waste pellets and charcoal may also be used. Most of the fuels will be blended with tHe coal prior to pulverisation, although the possibility of installing separate pu4verisers for the biomass fuels is subject of study.
AMER POWER PLANT UNITS 8 AND 9 The Amer power plant consists of two units, Amer unit 8 and unit 9. A waste wood gasifier producing fuel gas that is combusted in unit 9 will be in operation in the middle of 2000. Since 199Y1996 both h e r units have been suitable for the co-combustion of paper sludge (150 (wet) ktons per annum), although the paper sludge is mainly combusted in Amer unit 8. The quality of the type of paper sludge is good with respect to the level of contaminants. The large amount of moisture (50wt%) gave some handling problems during the demonstration tests that were performed in the
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winter of 1995196. If more than 150 ktons of wet paper sludge can be contracted, a sludge drier for dqmg at least part of the sludge is intended to be built.
MAASVLAKTE POWER PLANT A major activity at Maasvlakte Power Plant is the co-combustion of biomass pellets that are produced from sewage sludge, waste wood (untreated wood) and paper sludge of the biomass plant in the direct vicinity of the power plant site. The biomass pellets are transported by conveyor belts to the power plant and are blended with the raw coal in the raw coal bunkers. In the new Environmental Impact Study many other %els" were also considered. The fuels can be subdivided into three categories: 0 biomass (pellets, bone meal, animal fat, coffee ground) 0 carbon waste products (char coal, coal residues) fluids (by-products from the chemical industry, low halogen containing solvents). A total of 288,000 tons of wastehiomass - corresponding to 10 wt% of the annual coal consumption - is foreseen to be contracted, although a larger amount will be possible. The solid fuels will be blended with the raw coal and pulverised in the existing pulverisers, although the use of a separate milling system for some of the biomass products is also considered. The fluids will be stored in bins and pumped to the boilers. In both units fluid burners are mounted in two of the four burner boxes.
HEMWEG POWER PLANT In 1997 Hemweg Power Plant applied for a license to co-combust 75,000 tons of dried sewage sludge on a yearly basis. It was intended to blend the sewage sludge with the coal before the pulverisers, followed by an injection into the boiler through the existing coal burners. Although the license was granted, the decision to be taken by the new owner to start co-combustion on a commercial scale is pending.
Indirect co-cornbustion Monostreams that are not suitable for grinding or pulverisation before being fed into the burners should be pre-treated thermally: indirect co-combustion. Figure 2 summarises the possible types of thermal pre-treatment, including the necessary physical pre-treatment, taking into account all possible types of indirect cocombustion. The main (in)direct combustion or gasification projects in the engineering or planning phase are the following: 0
UNA is investigating the installation of pyrolysis units from Pyrovac, with a total capacity of 120 ktoda at the Hemweg power plant. With h s so-called Pyrocycling process the main objectives are to produce char and oil, which can be co-combusted directly with the pulverised coal or, if desired, be stored and co-combusted later at Hemweg or at another power plant.
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0
0
EZH is considering building a CFB combustion facility integrated with the Maasvlakte power plant. EPON has plans to integrate biomass and waste gasifiers with respectively their CG-13 and Eems power plant (Euroforum 1998).
EPZ has constructed a 90 MWth CFB gasifier at h e r 9 power plant for the gasification of 150 ktoda of waste B-wood (Euroform 1998, 1999). The fuel gas produced is cooled, cleaned and co-combusted in the 600 MWe boiler of Amer-9. With an efficiency of 35% it replaces 70 ktoda of coal and delivers 205 G W a with a yearly COz reduction of 170 ktoda. Start of the operation is planned for mid2000.
Fig. 2 Indirect co-combustion by thermal pre-treatment (pyrolysis, gasification or combustion). LIMITATIONS ON CO-COMBUSTION The replacement of coal by biomass in the existing boilers is not unlimited. The potential replacement depends on many factors, technical limitations as well as limitations imposed by the authorities, and financial and legal provisions. Focusing on the technical limitations it becomes obvious that the potential replacement of coal strongly depends on the quality (for instance the lower heating value) of the biomass, if the nominal load of the plant remains the same. Without intending to cover all aspects and limitations some examples of technical limitations are listed below: Grindability of the biomass/coal blend. The pulverisation of a coal biomass blend in the existing pulverisers is primarily limited by the grinding behaviour and the moisture content of the biomass. Most often biomass is soft or fibrous, whereas bituminous coal is hard. 809
Capacity of the unit components, such as the flue gas fans, bottom ash discharge system and air heaters. Very often the biomass has a high moisture content, as a result of which the flue gas flows significantly increase when a large proportion of coal is replaced by (wet) biomass. Severe slagging and fouling. The ash melting temperatures of some types of biomass may be low because of a high calcium and iron content. This may cause severe problems with the slagging of the burners, furnace walls and superheaters. A high sodium and potassium content may result in fouling problems. Corrosion of the h a c e walls and superheaters. Sulphur and chlorine may enhance the corrosion rate of the furnace walls. Severe slagging and fouling of the superheaters may locally result in deteriorating conditions resulting in corrosion. Erosion. A h g h ash content may increase the wear due to erosion, especially if the flue gas flows are increased because of high moisture content. Bed agglomeration, tar condensation (indirect applications). Other technical aspects that should be considered when maximising the biomass input in existing coal-fired power plants are the operating flexibility, the maximum load, and unavailability. All of these aspects are extremely important in a liberalised electricity market. Special attention should also be paid to the quality of the by-products, such as fly ash and bottom ash. The by-products have to be utilised in for instance the cement and concrete industry, and can not be used for landfill. The quality of the by-products and its application should be according to corporate and civil laws. European legislation is becoming increasingly stringent in its member countries. The insufficient affiliation of the European and domestic legislation results in a complex analysis of the status of by-products from co-combustion activities. At present in the Netherlands, fly ash from co-combustion activities, which contains less than 10 per cent (on a mass basis) of biomass, is still regarded as fly ash from coal. It is still uncertain what status fly ash fiom more than 10 per cent of co-combustion activities, will enjoy. Another important aspect is the market’s acceptance of reutilization of fly ash originating fiom co-combustion activities. With respect to the legislation of NO,, SO2,dust and heavy metals emissions, the situation is very complex, since the regulations, both on a national and a European level, are currently subject to revision.
CONCLUSIONS The contribution of biomass and waste to electricity production in the Dutch coalfired power plants is very rapidly increasing due to both political and economical incentives and is stimulated by fiscal measures. In order to achieve the objective of 10% energy supply by renewables in the Netherlands stimulation by offering financial and fiscal incentives is at present crucial due to the current price levels of competing (fossil) alternatives. In the year 2000 the coal-fired power plants in the Netherlands are producing sustainable electricity by means of direct co-combustion of more than 900 kton of biomass and waste. Stand-alone installations (for example Cuyk, h e r ) contribute 810
a further 400 kton. However, in order to meet the challenging objectives set by the Dutch government with respect to the contribution of renewable energy and the severe regulations concerning for example emissions and the use of by-products, many technical problems still have to be solved. These problems include the technical boundaries of increasing direct cocombustion and exploring different (in)direct thermal conversion routes, such as pyrolysis, gasification and combustion, in order to handle less cleaner fuels, whtle maintaining to fulfil all the requirements.
REFERENCES 1. 2. 3.
4. 5.
6.
7.
8.
9. 10. 11. 12.
AOONVAV (1998). Waste processing in the Netherlands, data 1998 (in Dutch) Beekes M.L. et a1 (1998). Co-combustion of biomass in pulverised coal-fired boilers in the Netherlands Euroforum (1998, 1999). The market for energy from waste and biomass, November 1998, Rotterdam and November 1999, Nijmegen Faay A.P.C. (1997). "Energy from biomass and waste", PhD thesis University of Utrecht GAVE (1999). "Availability of waste and biomass for energy production in the Netherlands", final report (in Dutch) KEMA (1999). (Kok, W.C., Boudewijn, R. and Lindeman, J.H.W.), Evaluation of environmental permits Dutch biomass projects. KEMA-report 56034-KPSMEC 99-305 1, December 1999 (in Dutch) Naber J.E. et al. "The HTU@ Process for Biomass Liquifaction; R&D Strategy and Potential Business Development", Proceedings of the Fourth Biomass Conference of the Americas. Oakland, Canada August 29 - September 2,1999 Novem, 9916. "Bio-masterclass, an overview of streams and prices", September 1999 (in Dutch) Novem, 1998a, (E. van de Heuvel, I. Gigler). "Is there ground for biomass Droduction in the Netherlands?", April 1998 -Novem, 1998b. "Biomass for sustainable energy", April 1998 (in Dutch) STORK (1999). Parallel Combustion of Biomass, summary report. Stork Engineering and Consultancy, 3 1 May 1999 W A V , 1999. From Waste to Clean Energy. Brochure published by the Dutch Waste Processing Association (WAV), 1999
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Whole Tree Energy Power Plant K. W. Ragland', L. D. Ostlie2 and D. A. Berg' I Dept. of Mechanical Engineering, University of Wisconsin, Madison, W,USA Energy Performance Systems,Inc., Minneapolis, MN, USA R.W. Beck, Inc., St. Paul, MN, USA
ABSTRACT A planned 50 MW Whole Tree Energy (WTETM)biomass-fired power plant and previous pilot-scale and small-scale testing is described. Hybrid poplar trees are planted and grown on land withm 80 km of the power plant site. The time from planting to harvest is five growing seasons, and the projected harvest yield is 56 dry tha. The harvested trees are trucked to the power plant site and dried in a drymg dome, which utilizes waste heat from the power plant. The whole (not chipped) trees are then fed by conveyor into a deep, futed-bed furnace which provides heat for a high pressure steam boiler-steam turbine-electrical generator. Flyash is removed from the stack gas with a wet electrostatic precipitator, and the ash is pelletized and used as fertilizer for the trees. Nitrogen oxides are controlled with extended over-fire air. A 50 MW power plant requires 18,400 ha of tree farms, which is about 1% of the land within a radius of 80 km. The required land is leased for a 15 year period. The first WTETMpower plant (50 MW) will be built in St Peter, MN, with projected startup in mid-2004. The average present value of electricity from a second 50 MW WTETM power plant is projected to be $O.O49kWh, and for 150 MW MW WTETMpower plant, $O.O36kWh (year 2000 basis, Federal tax credit for closed loop biomass included). INTRODUCTION This article describes a renewable, sustainable, closed-loop system for generating electric power from biomass - the Whole Tree Energy (WTETM)system'-6. Farm grown hybrid poplar trees are burned whole (not chipped) in a deep, fixed-bed fumace to generate superheated steam that powers a steam turbine-generator. Exhaust heat is used for air and feedwater heating and fuel drymg. The he1 supply, land requirements, power plant technology, equipment testing to date, project economics, and project status are discussed. The first WTETMsystem is a 50 MW power plant near St Peter, MN, with startup scheduled for mid-2004.
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FUEL SUPPLY The fuel supply is hybrid poplar and cottonwood trees grown as short rotation woody farm crops. The land is leased from the landowner. Each year for five years 4,000 ha is planted. Over the last 15 years various clones of hybrid poplars have been bred to be fast growing and disease resistant. Improvements in the genetic stock are ongoing. Hybrid poplars are beginning to be widely used in the forest and paper industries. Cottonwood, whch is in the same genus, is also very fast growing and disease resistant. Poplars will grow almost anywhere but best yield for energy crops is obtained by following proper agricultural practice. The trees grow tall and straight with relatively small branches. In the several years of operation before the energy crop is established, waste wood, over-aged stands, and natural gas will be used as the fuel. PLANTING AND FARMING
Cuttings from two-year-old stool bed trees are used to establish the tree farms. The two-year-old trees are harvested in winter and cut into 25 cm lengths with diameters of 8 mm to 25 mm. The cuttings are stored in boxes in an industrial freezer. The dormant cuttiags are planted in May when the soil temperature reaches 10°C. A rapid planting machme inserts a cutting into the soil such that one bud is exposed to the air. The cuttings are planted on 1.7 m centers, and a total of up to 4400 cuttings per hectare are required. The most successful clone developed to date for the upper Midwest is hybrid poplar NM-6 (Nigra X Maximawitzii), and this will be the main cutting. Other fast growing and disease resistant clones such as the eastern cottonwood varieties will also be used. Cuttings will be purchased from nurseries in Wisconsin, Minnesota, Michigan, North Dakota and Oregon. A pre-emergent herbicide is applied after planting, which provides weed control for most of the season. Young hybrid poplars cannot out-compete weeds. One tilling may be required in the first year for weed control, but not in the following years. A single final application of pre-emergent herbicide is applied in the second year. The clones selected have good pest resistance but some pest management may be required. Nitrogen, phosphorous and potassium levels in the soil must be maintained for optimum growth. The inorganic requirements are only about 10% of that required for a corn crop. The main source of nitrogen is from decomposition of fallen leaves. Flyash pellets from the biomass-fired power plant are applied to the fields once during the five year growing cycle to recycle the phosphorous, potassium and other trace elements. HARVESTING AND TRANSPORTING
The trees are ready for harvest beginning in the fall of the fifth growing season. The trees are typically 15-20 cm diameter and 12-14 m tall. A special harvesting machme is being designed by EPS with a grant from the U.S. Department of Energy. This machine, which is mounted on four rubber tracks, grabs the tree, cuts it off at the base, and holds the tree upright in an accumulator, while continuing to move down the row at up to 9 km/h. Near the end of the row the trees are loaded into a trailer carrying a 25 ton load of whole trees. In applications where wider row spacing and longer
813
harvest cycles are used, trucks may be directly loaded when a single row produces more than 25 tons. A 50 MW plant requires 52 truck loads per day of whole trees. After harvesting, the poplar trees sprout from the stumps and grow vigorously with little attention except fertilization with ash pellets from the power plant. Harvesting is done every fifth year.
PROJECT SITE AND LAND REQUIREMENTS A 30 ha site for the power plant is needed which has good road access, is close to an electric transmission line, has a natural gas supply for startup, and a makeup water supply. Approximately 6 ha are needed for the drymg dome and power plant, while the remaining 24 ha is to be planted with trees to provide a noise barrier and be aesthetically pleasing. Water loss is due to cooling tower evaporation and boiler water blowdown. The dedicated farmland should be within 80 km of the power plant site. The best soils for poplar trees are loams, sandy loams, and clay loam. Relatively fertile soil is recommended. Planting in very sandy soil is not recommended unless a sufEciently high water table is present. Areas prone to summer flooding should be avoided but poplars can tolerate standing water for short periods of time. Soils should be void of a hardpan layer. The soil pH should be in the range 5.5-7.5, but up to 8.4 can be tolerated by some cottonwood clones. The slope should be less than 14 degrees (or 25%) for efficient harvesting because of mechanical constraints. Five growing seasons are required to harvest. The yield is a function of the soil quality, tree spacing, clone type, fertilization, cultivation, pesticide application, and most importantly water availability. The design yield is 56 dry tons per hectare after five growing seasons. For a 50 MW power plant 20,000 ha are required, which is 1% of the land within an 80 ktn radius. If each tree f m averaged 32 ha then a total of 625 fields would be needed and 125 fields would be planted each year.
POWER PLANT TECHNOLOGY The key power plant technologies are the drying and storage dome, the fiunace, the boiler, steam turbine, generator, and emissions control equipment. The general layout of the plant is shown in Fig. I.
DRYmG AND STORAGE DOME Whole trees are delivered to the drying and storage dome (Fig. I) at the rate of 52 trucks per day, and the trees are removed from the trailer truck by a tower crane with grapple. A 30 day supply of wood totaling 19,000 oven dry tons is maintained in various stages of drymg. The pile is 24 m high by 85 m diameter (this is based on a pile density of 272 kg/m3 at 44% moisture). The dome, which is a pressurized twolayer fabric facility similar to those used to cover sports stadiums, is 36 m high by 150 m diameter. Heated air from heat exchangers, which transfer heat fiom the furnace flue gas to the drymg air, is circulated by means of zoned ducting underneath the pile and flows up through the pile and out through an opening in the top of the dome. The drying air enters the dqmg dome at 55OC and exits at 23°C. The delivered wood 814
typically has 44-50% as-received moisture, and after 30 days it is dned to 20-25% moisture. The whole trees are removed from the stack in the dome on a first-in, firstout basis by the overhead crane with a specially designed grapple and placed on a conveyor in batches for delivery to the fumacehoiler. There is a possibility to plan for greenhouse space between the outside of the tree drylns pile and the dome walls, pending consideration of safety issues. The greenhouse area available in the drymg dome is up to 9300 m2. FURNACE Whole tree batches are transported on the conveyor to the furnace charge chamber (Fig. 11). While on the conveyor a heavy duty sectioning saw cuts the batch to length and the batch (up to 4.5 ton) is pushed into the charge chamber from the top. The top door closes, sealing the chamber and a charge ram forces the wood through a furnace entry door and onto the top of the fuel bed. The fuel bed, which is typically 3 - 4 m deep, is supported by a patented water cooled grate with controlled circulation. Preheated air from a second heat exchanger, which transfers heat from the flue gas to the furnace air, flows upward under the grate and also above the fned fuel bed (overfire air). The furnace design heat release rate of 7.8 MW/m2 based on pilot scale test results (see below). For a 50 MW (electrical) power plant the required grate size is approximately 6 m long by 3 m wide. The combustion process involves a patented three-stage process:
(1) Char in the bottom 30-60 cm of the deep bed burns on the grate where available oxygen is consumed. The hot gases from the burning char flow upward through the fued bed of whole trees and drive the volatiles from the wood (pyrolysis). As the char at the bottom is consumed, the bed partially subsides and another batch of wood is fed from the charge chamber above the bed. (2) Above bed over-fire air is strategically mixed with the volatiles, which consist primarily of CO, COz, H2, CHI, HzO, N2 and tars. Combustion of the volatiles proceeds, the tars are burned out, and heat is transferred to the boiler walls and convective tubes. Reactions above the bed occur first in a reducing environment, and then excess air is added gradually to the upper combustion zone. (3) Any char that falls through the openings of the grate at the bottom of the bed is collected on a lower grate and burns out below the bed.
815
Figure. Z Top and side views of drymghtorage dome and power plant.
Figure IZ. Schematic showing furnace, ram feeder, steam generator, flue gas heat exchanger, and drylng dome. 816
BOILER-STEAM TURBINE-GENERA TOR The boiler produces 150,000 k g h of steam at 13.8 MPa and 540°C exhausted to a pressure of 6 kPa absolute. The steam turbine is a dual casing, single axis machine with six extraction points and one reheat at 540°C. The turbine efficiency is 43.8%Btu. The generator is rated at 57,800 kVA, 3600 rpm,3-phase, 60 Hz, 13.8 kV. The generator has a gross capacity of 52.4 MW to provide for 4.8% auxiliary power for boiler feedwater, scrubber and cooling tower pumps; combustion air and wood drying air fans; crane, conveyor, and ram feeder power; and miscellaneous power. The expected boiler efficiency is 83%. The overall efficiency of the power plant6 is expected to be 33% (based on the higher heating value) for the first plant.
EMISSIONS CONSIDERATIONS Particulate control is by means of either a wet scrubber or a wet electrostatic precipitator located after the condensing heat exchanger. The choice of particulate control equipment depends on the degree of control required. The particulate slurry collected by the scrubber or precipitator is circulated through a pug mill that de-waters and pelletizes the ash. The ash pellets are spread on the tree fields as a fertilizer. The ash content of the wood is less than 1% so that a maximum of 5 t/day are collected and pelletized. The particulate emission standard to be met for new wood fired power plants in the State of Wisconsin, for example, is based on best available control technology (BACT) and can be expected to be about 21 g/ lo6 kJ. Nitrogen oxide control is by means of over-fire air jets whch are located at strategic positions above the reducing zone of the fuel bed. It is expected that nitrogen oxide formation from fuel nitrogen will be small because is of the extended reducing zone of the fuel bed. The nitrogen oxide emission standard for wood fired power plants is often based on BACT and is set on a case by case evaluation. Sulfur dioxide and carbon monoxide emissions are very low and emission standards do not apply. However, CO is used as a surrogate for products of incomplete combustion (PICs), and it is likely that the CO emissions would be limited to 300 ppm (dry volume basis) at 7% oxygen.
TECHNOLOGY TESTING TO DATE
Planting Tests: A prototype machine has been developed to plant 25 cm long tree cuttings at a rapid rate. Approximately 180,000 cuttings were planted at the rate of 1 per second in a single row. In the operational version of the planter six rows at a time will be planted. The planter can be used on unprepared Conservation Reserve Program land and difficult-to-farm land also. Harvesting. Loading, and Transportation Tests: Tree harvesting at the rate of 45 green t/h has been demonstrated using a single harvesting "train" of one feller-buncher, one
skidder, and one loader. Transportation of 27 ton whole tree loads on logging roads and public highways was done. Crowns and limbs did not pose a problem with loading or transporting whole trees.
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Stacking Tests': Whole hardwood trees were stacked to a height in excess of 30 m. Each layer of trees was placed at 90" to the layer below in a 21 m by 2 1 m wide array. The weight of the trees tend to compress a notch in the wood where other trees made contact, whch enhances the stability of the stack. As a test, a 90 kN side load, which was applied to trees near the top of the stack, had no visible affect on the tree stack. In another test a stack with a 110" angle of repose overhang was created without any apparent instability. Thus the lateral stability of a large stack of whole trees (tops and limbs intact) is exceptional. Dryinn Tests': A 21 m by 21 m by 26 m tall stack of whole hardwood trees was supported by an air distribution manifold. Air was heated by a heat exchanger using a propane burner and ducted into the distribution manifold. The stack of trees was b e d with 58OC air for 30 days, and the average moisture level was reduced from 44% to 20%. There was no pressure drop through the stack of trees because of buoyancy; rather the pressure was -125 Pa at the base of the stack. These tests established the feasibility of drylng a large stack of green trees with waste heat from the flue gases. Combustion tests' were conducted at the Bay Front Unit No. 3 of Northern States Power Co. in Ashland, WI. The coal-fired underfeed stoker with boiler rated at 45,400 kg/h steam at 5.1 MPa drum pressure and 455°C was modified to receive 4.6 m long whole tree sections from a charge chamber by means of a ram feeder. A higher pressure over-fire air system was added. The tests successfully demonstrated the feasibility of replacing coal with logs on a grate without a loss in boiler performance. Pilot-scale combustion tests"' of a deep fvred bed of sectioned hardwood trees on a grate were conducted in a 1.4 m by 2.6 m by 6 m high test furnace located in Northern Minnesota (Fig. 111). Under-fire air was preheated to 275°C using a two-pass plate and frame heat exchanger. The air flow rate was set at 565 kg/min which gave an air velocity under the bed of 4.0 d s . The fuel bed, which was supported by an air cooled grate, was maintained at a depth of 3.7 m by feeding wood into the top of the furnace with an inclined chain conveyor. The tree sections were 2.5 m long, with log diameters up to 20 cm, and the average moisture level was 3 1.6%. The weight of each tree section was measured on a weighing table before being placed on the conveyor. The average void fraction of the bed was 0.65. There was no overfire air and combustiodpyrolysis products were vented to the atmosphere. Based on the wood feed rate required to maintain the level of the bed over a 2 hr period, the average effective heat release rate during the tests was 10 MW/m'. These tests demonstrated the intense nature of the combustion and pyrolysis in a deep bed of sectioned whole trees. From the measured wood consumption rate and data from single log tests (see below) a computer model of a deep fmed bed was developed and validated'. The model relates heat release rate to under-grate air flow and preheat, fuel moisture, size and void fraction, and bed height. Based on this, a design heat release rate of 7.8 MW was selected for the first 111-scale unit.
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.... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... .... ....
48m
I
8 cm od Saeel Tu be5 on 22 cm
18m
-
-
d
*
Fire Rick Insulating Wick %I Plate
....................... .... ........................ ....
Preheaed Air
Figure I11 (a) Pilot-scale, open-top furnace test rig; Preheated Air
-
E haust
BLlrner
Ekchangw
-Air
E-
Air
Figure I11 (b). Heat exchanger for preheated combustion air.
Sinale log combustion testsg were conducted at the U. S. Forest Products Laboratory in Madison, WI, in a specially designed furnace that was supported by load cells (Fig. IV). These single log tests established the burning rate of individual logs up to 20 cm in diameter in high temperature oxidizing and reducing environments and confiied the high heat release rate observed in the pilot scale tests. The detailed data obtained was used to develop a computer model' that simulates the performance of the deep fixed-bed furnace over a range of conditions.
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Natural Gas Air
Schematic of test fitmace setup. Figure IV. Furnace to obtain burning rate of single logs. Scrubber Tests: The wet scrubber was tested in 1978 at the Sherburne County Generating Station of Northern States Power on Unit 1, a 757 MW coal-fired unit. In ths unit the combustion products were ducted to 11 scrubbers each cleaning the equivalent of 68 MW of gas. The “multiple angle scrubber” was installed in one of the scrubbers replacing the existing rod scrubber. Particulate emissions averaged 174 g/106W on coal based on follow-up tests. L. D. Ostlie was granted US patent number 4,3 13,742 in 1982 for the multiple angle scrubber. If the wet scrubber is deemed to be inadequate, a wet electrostatic precipitator will be used.
ENVIRONMENTAL IMPACT Energy crops provide vegetative cover throughout the year, reducing soil erosion and improving wildlife cover, unlike annual row crops. Much lowered applications of agricultural chemicals and reduced tillage benefit water quality. Conversion of agncultural land from row crops to woody crops improves soil structure, organic matter content, and water quality. Woody crops develop an extensive root structure that adds organic matter to the soil, slows wind and water erosion, and helps to reduce soil compaction. Soil nitrogen and inorganic nutrients are maintained by means of a controlled combination of fertilizer, leaf litter, and fly ash pellets from the power plant. In regions with high levels of nitrate pollution in the ground water, planting hybrid poplar trees has been shown to reduce the nitrate levels by a factor of ten or more when the water table is 1.5 m or less below the surface”. Woody crop cover in agricultural areas benefits a wide variety of birds, small mammals, and deer. Woody crops provide edge effects and comdors with native habitats thereby improving wildlife diversity. Plots averaging about 32 hectares are envisioned so that landscape diversity is preserved. The Whole Tree Energy power plant is clean burning, and should meet all Federal and State emission standards. The power plant is carbon dioxide neutral since
820
the carbon emitted by the power plant is balanced by the carbon previously sequestered by the trees and roots.
PROJECT ECONOMICS Project costs consist of fuel costs, capital equipment and construction costs, fees, operations and maintenance, and financing costs. Farmland rent for trees is assumed to be $173/haJyear. The estimated cost to establish and tend the trees is $760/ha for the first year and $543lha for the second year; each year thereafter is $247/ha. Harvesting and transporting the trees to the power plant site at the end of the fifth growing season is $402/ha. For a new power plant one-fifth of the acreage is planted in each of the first five years; then the trees are cut in five-year periods. The cost of the fuel, excluding establishment, but including rent, tending, harvesting and transportation is $1.35/109Jbased on year 2000 costs. The amount of land required depends on the tree yield, power plant efficiency, and the capacity factor. The average biomass yield was assumed to be 11.3 dry tonihectare. The power plant efficiency is based on a dry higher heating value of 20,200 Hkg. A power plant annual capacity factor of 86.3% was assumed. Power plant sizes of 25, 50, and 150 MW were considered.
Table I . Fuel supply for 20 year period. Net Power Wood Growth Power Plant Plant EfficRate Dryton Size* iency MW idyr % 28.4 11.3 25 11.3 50 32.5 11.3 150 34.1
Total Land
hectare 10.520 18,400 52,610
Capital equipment includes fuel handling, storage, drylng and feeding systems; steam generator and support systems; turbine-generator and support systems; condensate and feedwater systems; circulating water and treatment systems; electrical and control systems; power plant facility; and power station construction. Engineering, procurement and construction costs on a turnkey basis are shown in Table 2 along with additional site costs, fees and contingencies.
Table 2. Capital costs (year 2000 basis). Power Plant EPC EPC Site Costs, Tree Farm Total Size (MW) Equipment Equipment Fees, Establish. Capital Cost Unit Costs Costs Contingen. ($ million) ($kW) ($ million) ($ million) ($ million) 46 18 21 85 25 1840 50 1500 75 26 37 138 150 1090 163 52 106 32 1
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To fund the project the assumed financial arrangement is 25% equity and 75% debt. Return on equity is assumed to be 25%, and interest on the debt is assumed to be 8.5%. Hence, the effective interest rate is 12.63%. The annual operating costs are $3.2 million, $3.9 million, and $8.6 million year respectively for the three power plant sizes on a year 2000 basis, excluding the fuel. The property tax rate is 2.1% in Wisconsin. For power plants that are over 50 MW and sell at least 95% to a power company that sells at retail, a gross receipts tax can be paid instead of a property tax, whch is 3.19% of the sales revenues and is a significant savings over a property tax. No tax moratorium was assumed. No federal capital contribution was assumed, but a closed loop biomass tax credit of 1.79 centskWh was taken for the year 2000 and escalated at 2.6% per year for 20 years. Year 2000 costs were calculated by discounting the revenues each year using 3.75% discount rate, whch was the rate used to escalate operating costs. The 20 year cumulative discounted revenues were divided by the 20 year cumulative value of electricity delivered. The projected 20 year average present value of electricity for the second WTETMpower plant is shown in Table 3.
Table 3. Average present value of electricity (year 2000 basis). Net Power Plant Size Cost of Electricity* (MW) (%/kW h) 25 0.064 50 0.049 150 0.036 *includes Federal tax credit for closed loop biomass CURRENT STATUS OF WHOLE TREE ENERGY
Research on poplar tree clones for bioenergy by Oak Ridge National Laboratory and its affiliated Universities, by the US Forest Service-North Central Forest Experiment Station, and by members of US. Poplar Council is extensive. As a result of this research, an EPS affiliated nursery is currently growing 8 ha of NM-6 hybrid poplar stool trees. Other commercial nurseries are also growing hybrid poplar and willow clones for cuttings. EPS holds patents in 30 countries in North and South America, Western and Eastern Europe, Asia, and Australia covering the Whole Tree Energy technology. A rapid harvesting machine has been designed and is being built and tested under a contract with the U.S. Department of Energy. A machine for harvesting whips (twoyear-old trees from which the 25 cm long cuttings are made) is being developed. A prototype machine for rapid planting of the tree cuttings also is being developed. The first Whole Tree Energy power plant is under development by EPS/Beck Power, which is a limited liability corporation formed by Energy Performance Systems, Inc. (EPS) and R. W. Beck, Inc. EPS develops and holds patents on the WTETMbioenergy technology, and R. W. Beck is an engineering consulting company. In January 2000 a power purchase agreement between Northern States Power and ESP/Beck Power for a 25 MW Whole Tree Energym power plant was approved by the Minnesota Public Utilities Commission. In late April NSP expanded the project to 50 MW.. Startup is scheduled for the mid-2004 at a site near St. Peter, MN. The 822
agreement covers 20 years of operation at a capacity factor of 86.3%. Private equity and debt arrangements are being negotiated. Individual long-term (15 yr) leases of land for tree farms are being sought with landowners within 80 km of the power plant site. ACKNOWLEDGEMENT The information collected for this article was made possible in part by a grant from the Wisconsin Energy Bureau.
REFERENCES EPRI (1993) Whole Tree EnergyTMDesign, Volume 1: Engineering Evaluation, TR-101564v.1, Electric Power Research Institute. 2. EPRI (1993) Whole Tree EnergyTMDesign, Volume 2: Program to Test Key Elements of WTE, TR-101564v.2, Electric Power Research Institute. 3. EPRI (1993) Whole Tree EnergyTMDesign, Volume 3: 100 MW Design, TR101564v.3, Electric Power Research Institute. 4. EPRI (1995) 100 MW Whole Tree EnergyTMPower Plant Feasibility Study, TR 104819, Electric Power Research Institute. 5. Lamarre, L. (1994 Jan-Feb) Electricity From Whole Trees, EPRI Journal, 16-24. 6. Ostlie, L. D. & Ragland, K. W., (1998) High Efficiency Bioenergy Steam Power Plant, in Bioenergy’98, pp. 772-78 1. 7. Withrow, K., Wichert, D. & Moran, D. (1999) The State of Wisconsin Energy Geographic Information System, Wisconsin Energy Bureau, Department of Administration. 8. Bryden, K. M. & Ragland, K. W. (1966) Numerical Modeling of a Deep, Fixed 269-275. Bed Combustor, Energy and Fuels, American Chemical Society, 9. K. M. Bryden & K. W. Ragland (1997) Combustion of a Single Log under Furnace Conditions, Developments in Thermochemical Biomass Conversion, (Ed by A.V. Bridgwater & D. B. G. Boocock), pp.1331-1345, Blackie Academic and Professional. 10. Licht, L. A., Schnoor, J. L., Nair, D. R. & Madison, M. F. (1992) Ecolotree Buffers for Controlling Non-point Sediment and Nitrate, ASAE paper 922626, Nashville, TN. 1.
a,
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Influence of Ash Composition on Slagging and Defluidisation in a Biomass Fired Commercial CFB Boiler Anncharlotte E. Tranvikl, Mehri Sanati2, Bjorn Zethraews2 and Mats Lyberg3. 1 VEAB,Kvarnvagen 35, S-352 41, Vtixjo,Sweden. 2 The Bioenergy Center at IBP, Vtixjo University, S-351 95 Vaxjo, Sweden. 3 Division of Physics at MSI, Vtixjo University, S-351 95 Vtixjo, Sweden.
ABSTRACT The purpose of this work is to determine the physical and chemical reactions, whch lead to the undesired agglomeration of bed material in a commercial circulating, fluidised bed (CFB). Samples of bed material were randomly collected and analysed using different methods. The analysis of the bulk chemical compositions (major, minor and trace elements) by x-ray fluorescence (XRF) revealed no significant variation in potassium content. But the zinc, copper and magnesium contents showed a maximum at the same sample size. This behaviour can indicate that alkali has a minor roll in agglomeration phenomena. Other techniques as Inductive Coupled Plasma (ICP-OES), Scanning Electron Microscopy (SEM) and Atomic Force Microscopy (AFM) have been used for characterisation purposes. On the basis of the AFM experimental results and evidence of the non-height difference in the topography of the agglomerated samples, it was concluded that there is a homogenous melt fonnation. BACKGROUND A growing interest, by public opinion and authorities, for environmental issues have resulted in increased research in topics concerning utilisation of biomass fuels. One of the important and as yet unsolved problems regardmg the utilisation of biomass fuels, is the appearance of large agglomerates in different parts of the boiler, for example the furnace and the separator. These slagging or sintering products can cause blocking of both material and gas flow, which results in unfavourable combustion conditions and can even result in unplanned shut down of the heat and energy production unit. During the combustion process, inorganic species bound to the organic molecules are released and in some cases volatilised. These species react and produce more or less crystalline inorganic compounds that adhere to any available surfaces.
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Different mechanisms for the agglomeration process have been suggested like the formation of eutectic melts, solid state sintering, condensation of vapours and viscous flow sintering of silicates. It has been suggested that high potassium content of the fuel result in bed agglomeration. Several commercial boilers have their bed replacement program determined by the potassium content of the bed material. In a fluidised bed boiler there is a large amount of sand present. This often has a high content of silica, e.g. quarts. Alkali silicates are thought to be responsible for the phenomena since their melting points are similar to temperatures used in the boiler EXPERIMENTAL COMB USTION CONDITIONS
In order to identify the chemical composition of the material acting as an adhesive, several different samples collected in a commercial 104 MWth CFB Co-generation Boiler (at VEAB) have been analysed using various techniques. The fuel mixture of the boiler consisted of 50 % forest residues and 50 % sawdust. Feldspar sand with a mean particle size of 0.28 mm has been used as bed material. The typical sand consumption was 6-8 tons/day at high loads. Different types of ash material, including nonagglomerated (free flowing) bottom ash, small agglomerates in the bed and larger agglomerates found in the lower parts of the riser have been analysed. The latter samples have been collected during shut down of the plant and it was not possible to determine under which conditions they have been formed. CHARACTERISATION OF ASH MATERIAL
In this study the results obtained fiom the following analytical techniques have been compared with each other: 1. 2. 3, 4.
Mechanical sieving Induced Coupled Plasma Optical Spectrometry Scanning Electron Microscopy Atomic Force Microscopy
Mechanical sieving A homogenised sample of 100-g bottom ash was sieved for ten minutes. The material
was separated on six different sieves with spacings of 0.125mm, 0 . 3 5 5 O~S O O m m , 0.71mm, 1.OOmm and 1.40 mm. The mechanical sieving of bottom ashes from different periods during combustion shows that when the sand addition decreases, the size of the particles in the bottom ash increases. Ths is most evident in the fractions around 0.71 mm. When these particles were crushed with a mortar and studied by optical microscopy, it was seen that they consisted of a kernel with a surrounding shell. The kernel was actually mini agglomerates of transparent small particles whilst the shell was more concrete like. The crushed samples were mechanically separated into shells and kernels. A comparative analysis of these was made with an ICP-OES instrument.
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ICP-OES Analysis
This was performed by an authorised analytic laboratory on an ICP-OES instrument, Optima Perkin Elmer 3000 DV. The results are show in Table 1. Table J Elemental Analysis by ICP-OES Elements Sodium Magnesium Aluminium Silica Phosphorus Potassium Calcium Titanium Manganese Iron Barium Arsenic Cadmium Cobalt Chrome Copper Nickel Lead Vanadium zinc Molybdenum
unit weight % weight % weight % weight % weight % weight % weight % weight % weight % weight % weight % Pdg Pdg Pglg Pdg Pdg PSIS Pdf3 Pdg
Sample 1 Sample 2 A [%] Significance
Pdg
pg/g
2.42 0.78 5.28 31.4 0.35 7.47 5.58 0.12 0.49 0.91 0.19 <20
2.4 0.57 5.06 33.9 0.32 7.15 4.68 0.05 0.44 0.36 0.16 <20
<5 16 118 7 10 10 3100 4 0
<5 10 89 4 6 10 2700
0.8 26.9 4.2 -8.0 8.6 4.3 16.1 58.3 10.2 60.4 15.8
37.5 24.6 42.9 40.0 0.0 12.9
+ -? +?
+ + + + +
+ + + + +
Sample 1 is from the whole particles and Sample 2 mainly consists of kernels. Silica is the only element that has a high concentration in the kernel. Most metals are of higher concentration in the whole particle, indicating that they are enriched in the shell. Iron, titan, nickel and lead are those with the largest differences in concentration in shells and kernels. However it should be pointed out that these results were obtained from only one randomly chosen sample and the concentrations were in some cases very low. Therefore, the test needs to be repeated to c o n f i i the above observations.
XRF analysis of different fractions The analysis of the elemental composition of particles from the bottom ash showed that the composition was strongly dependent on the particle size (see Figure I).
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3 ,
35
10 0.5
t 5
-
0
.-
0.4
0.5
I
+Ca
n
0.8
0.1
i
J
Average diameter [mm]
Figure 1 Elemental composition of particles of different size These analyses were made on a Phdlips PW 2400 XRF spectrometer, and showed that the 0.71mm fraction contained the maximum content of several metals, notably magnesium. Again thls is based on results with only one randomly chosen sample and the observations need to be verified with some more analysis. SEM analysis Samples was grinded, covered with carbon and analysed with a Scanning Electron Microscope (JEOL-JXA-840 SEM). An agglomerated sample is shown in Figure 2, where it can be seen that the grains are held together with an adhesive. Elemental analysis with EDX shows differences in composition between grain and adhesive a fact, which has also been reported earlier6.
Figure 2 SEM picture of an agglomerated sample
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The composition of the adhesive shows similarities with that of the kernels. However these analysis were made using different techniques, and comparison using the same technique is required. Atomic Force Microscopy (AF'M)
Some of the samples was analysed with a TopometrixBDiscoverer AFM unit. Two examples are shown in Figure 3 and Figure 4, which correspond to respectively nonagglomerated and agglomerated random samples taken at the summer shut down of the boiler in 1998. The upper parts of the pictures shows the atomic force micrograph and the lower parts the topography of the sample. The analysed area is 20*20 pwith a 400*400pixel resolution.
Figure 3 AFM image of a nonagglomerated sample. zmin OllUl zmax 6575.7 nm Scanrange 20p Resolution 400x400
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Analysis of non-agglomerated sample indicates an inhomogeneous surface consisting of several different chemical compounds. On the other hand, the agglomerated sample seems to have a homogeneous surface, indicating that it is covered by only one single chemical compound.
Figure 4 AFM image of an agglomerated sample. Zmin O m zmx 7941.8nm Scan range 20 P Resolution 400x400 RESULTS AND DISCUSSION
Many different mechanisms have been proposed in the literature for the agglomeration of bed particles. Viscous flow sintering of liquid silicates, reactive liquid sintering of molten salt systems, chemical reaction sintering by formation of new compounds, solid state sintering and vaporisation followed by re-condensation have been reported by several investigator^'-^^^. Chemical equilibrium models have
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been used to predict binary and ternary compounds, often silicates, with melting points in actual temperature interva11>8-10. As previously stated6 the results of this study are not in line with earlier work. The study of the assumed critical particle size 0.71-1.00 mm shows that the difference in potassium content, between kernel and shell is insignificant. However the difference is larger for other metals, such as magnesium, copper and zinc. This may indicate that some of the trace elements have a bigger role in the agglomeration mechanism. This will be further investigated with complementing analysis methods such as Thermo Gravimetric Analysis and X-Ray Diffraction. AFM analysis of a randomly chosen deposit indicates that the surface consist of one singe chemical compound. This is similar to the behaviour of glass-formic, amorphous compounds. The conclusion is that the original sand particles are covered with a homogenous layer of a single chemical compound. Whether this is the same as the compound covering the particle size in the range of 0.71-1.00 mm studied in this work is not clear. More work needs to be done to clarify this matter
REFERENCES 1.
2.
3.
4.
Blander, M.,Milne,T., Dayton,D., Backman, R., Blake, D., Kuhnel, V., Linak, W., Mann, M., Nordin, A., and Ljung, A. (1997) Equilibrium chemistry ofthe combustion of biomass: a round robin set of calculations using computer programs and data bases. Proceedings, Engineering Foundation Ash Conferance, Kona Hawaii. Brown, R. C., Dawson, M. R. and Smeeni, J. L. (1996) Bed Material Agglomeration During Fluidised Bed Combustion, Final Report, Iowa State University. Etiegni, L. And Champell, A. G. (1991) Physical and chemical characterisation of wood, Bioresearch. Technology 37, 173 Grubor, B., Oh, S., Ilic, M., Dakic, D. And Arsic, B. (1995), Biomass FBC combustion bed agglomeration problems, Proceedings 13'th International Conference on Fluidised Bed Combustion, May 7-10, Orlando, Florida, pp. 422515
5.
Kallner, P. And Zintl, F. (1997) Causes of deposits to superheaters and other heat exchange surjiaces when using biomass fuels and additives to reduce problems, Varmeforsk Report No 03-5 16, Stockholm ( in Swedish) 6. Tranvdc, A. E., Sanati, M., Zethraeus, B., and Lyberg, M. (1999), Study of the bed fouling problems in a biomass fired commercial CFB boiler, Proceedings, International Joint Power Generation Conference, Burlingame, California, USA 7. Manzoori, A. R. And Agawal, P. K. (1994), Agglomeration and defluidisation under simulating fluidised-bed combustion, Fuel 73,563-568. 8. Skrifvars, B. J. (1993, Sintering tendency of differentfuel ashes in combustion and gaslfication conditions, Thesis. Abo Academy University, Finland. 9. Nordin, A. (1993), On the combustion and gasification of biomass fuels, peat and waste: environmental aspects, Thesis, Department of Inorganic Chemistry, University of Umel, Sweden 10. &man, M. (1999), Experimental studies on bed agglomeration duringfluidised bed combustion of biomass fuels, Thesis, Department of Inorganic Chemistry, University of UmeA, Sweden
830
Utilisation Of Bagasse Residues In Power Production David Beckman a and Yrjo Solantausta a Zeton Inc., 5325 Harvester Road, Burlington, Canada L7L SK4 VTTEnergy, Biotonginkuja 3-5,02044 Espoo, Finland
ABSTRACT: Bagasse residues, produced fiom sugarcane processing, are the largest industrial biomass residues worldwide. Co-generation of heat and power is currently the only commercial large-scale use of this biomass waste. The Rankine power plant, which is the industrial technology, has a low power-to-heat ratio. Increasing the ratio would be desirable, as thls would potentially lead to lower cost of electricity, and a more efficient utilization of the bagasse. As part of the IEA Bioenergy Technoeconomic Assessment Task, alternatives to the Rankine cycle have been studied. Alternative bagasse energy concepts were studied by considering the current sugar mill bagasse utilization and energy requirements, and integrating bagasse pyrolysis and gasification to improve the mill's energy efficiency. In both cases combined heat and power cycles using gas turbine power plants were compared to the conventional Ranlune cycle. Both systems are under development and offer some advantages. The gasification system has an especially hgh efficiency, and the technical uncertainties are well known. The power plant fuelled with pyrolysis liquid may be operated for periods extending beyond the typical 180 days operation of a sugar mill by storing the pyrolysis oil. This concept has some technical uncertainties. Both advanced system have a higher efficiency and a power-to-heat ratio than the Rankine cycle. Estimated cost of electricity for cases under different operation modes are presented and compared.
INTRODUCTION Bagasse is a waste biomass from the sugarcane refining process. Bagasse residues represent the largest quantity industrial biomass waste available worldwide and therefore the most significant potential source of power produced from biomass. Figure 1 shows a comparison of the potential number of power plants by 2025 at 7 and 40 MW of electricity production for various biomass residues'.
83 1
Sugar Industry Residues
Residues from Pulping
Other Residues
Plantation
I
2,500
40 MWe
7 MWe 40 MWe
40 MWe
7 MWe 40 MWe
7mwe
7 MWe
Feedstock / Plant capacity
Figure 1.-Number of Potential Power Plants by 2025 for Different Industrial Biomass Residues As methods are considered to increase the usage of bioenergy globally as a way to reduce greenhouse gases, bagasse utilization will have a significant impact on the global bioenergy production. This study considers ways to improve and increase bagasse utilization by increasing the efficiency of electricity production through advanced bioenergy processes, and therefore replacing electricity produced by fossil fuels. Figure 1 shows that the majority of bagasse residues exist in Latin American countries. As an example, the country of Guatemala alone produces 28,600 tonnes of wet bagasse (50% moisture) daily during the sugar cane harvest period lasting for 150 to 180 days a year. Table 1 shows the typical sugar mill capacity and bagasse production for eight mills in Guatemala2. Significant quantities of bagasse are produced for consideration of integrating advanced bioenergy concepts at even the smallest capacity sugar mills. Table 1: Sugar Mill Plant Capacity and Bagasse Production in Guatemala (short tons, 50% moisture)
Facility
Daily Bagasse Produced (tons)
Daily Cane Capacity (tons)
832
Annual Sugar Produced (tons)
CURRENT BAGASSE UTILIZATION In the production of sugar from sugar cane, cane stalks are harvested and transported to sugar mills for processing. Cane stalks are crushed to release the sugar-laden juice, whch is clarified and then reduced by steam heated evaporation processes until the desired level of sugar concentration is reached, followed by various refining processes. The fibrous remains of the crushed cane, called bagasse, is burned wet in boilers to produce steam that is used in the various processes. The average sugar mill has a daily capacity of 5,000 to 25,000 tons of raw cane throughput per day. The harvest period is 150 to 180 consecutive days annually, and during this period the mills operate continuously (24 hours per day, 7 days per week). Many sugar mills also generate electricity for their own consumption using steam turbines, and sell any excess electricity through the local grid during the 150 to 180 day period of operation ofthe sugar mill3. The steam required by the sugar mill for cane processing is low pressure steam. In mills where electricity is produced a typical Ranlun steam cycle is used. In sugar mill operation the amount of bagasse available exceeds the amount required for energy to produce low pressure steam for cane processing. The amount of bagasse used to produce steam for sugar production can range from 60 to 90 percent of the total bagasse available, depending on the efficiency of the bagasse combustiodboiler system. For the purposes of this study a sugar mill steam demand equivalent to 75 percent of the bagasse available was assumed for a basis. In the mills that do not produce their own electricity by co-generation, t h s excess bagasse is a waste that must be disposed. It is either burned in open fields or given to nearby mills that have co-generation capabilities. For the purposes of t h ~ sstudy, the integration of advanced bioenergy processes into a sugar mill considered processing only the amount of excess bagasse available. This was viewed as a more realistic approach to the energy integration as it would allow a mill to continue to use its current system of existing boilers to produce steam for the mill. Bagasse is approximately 50% water by weight, and roughly 25% of the weight of raw cane milled becomes bagasse. For example, in a sugar mill with a 10,000 ton per day throughput, roughly 2,500 tons of bagasse will be produced daily. Typical properties of bagasse are shown in Table 2.
Table 2: Typical Physical Properties of Bagasse (moisture free basis)
ILower Heating Value (dry)
I
13.6MJkg
Basis of Study
The sugar mill operating conditions vary from one mill to the next. Each mill operates at different capacities for different periods of time and with different efficiencies of the existing bagasse combustion systems. Therefore average conditions were chosen to study the overall feasibility of the energy integration technologies. However, the 833
conditions particular to a specific mill can be inputted to the material and energy balance model that has been prepared to study a specific sugar mill application. Table 3 shows the typical sugar mill operating conditions that form the basis for the pyrolysis combined cycle concept. The actual amount of bagasse fed to the existing boilers is 70 percent of the total bagasse available. This is 5 percent less than the amount that would be used in the boilers if the bagasse pyrolysis system was not used. In this case the additional steam required by the sugar mill is produced in the heat recovery boiler in the combined cycle power section and passes to the mill after passing through the steam turbine. Table 3: Sugar Mill Operating Basis for Energy Integration
SUGAR MILL ENERGY INTEGRATION CONCEPT
PYROLYSIS COMBINED CYCLE Figure 2 shows a block diagram of the concept studied for the integration of a bagasse pyrolysis combined cycle process into a sugar mill. The integration of the pyrolysis concept into the sugar mill results in additional energy efficiencies beyond those realized by the pyrolysis combined cycle alone. The heat contained in the flue gas from the existing bagasse boilers, which is normally lost, is used in the bagasse dryer in the pyrolysis plant. This increases the energy efticiency of the pyrolysis plant since normally a small portion of the biomass feed to a pyrolysis plant is needed to supplement the energy required for drymg. Low pressure steam from the steam turbine is used to supplement the sugar mill steam demand, allowing more bagasse to be processed in the pyrolysis combined cycle to produce electricity. Wet bagasse is fed to the existing boilers, which is the current method of bagasse utilization in the sugar mills. A portion of the wet bagasse is dried to 10 percent moisture and fed to a fast pyrolysis plant. The char and gas by-products of the pyrolysis process are burned to produce the energy needed for the pyrolysis. Flue gases from the char combustion are sent to the dryer as a heat source.
834
"-
WET BAGASSE
FLUE GAS
Y
CHAR BURNER
3
L 2
-
EXISTING SUGAR MILL BOILER
STEAM TO PLANT
Figure 2: Sugar Mill Energy Integration by Bagasse Pyrolysis Combined Cycle Figure 3 shows a flowsheet of the pyrolysis section of the bagasse pyrolysis combined cycle concept. The vapours from the pyrolysis of the dry bagasse are condensed to produce a homogeneous bio-oil. The bio-oil can be stored and used in the power production section of the plant as required. Since wet bagasse cannot be stored for long periods of time, the ability to store bio-oil produced from bagasse allows the sugar mills to produce power with their surplus energy from the bagasse for longer periods of time. In the power section of the plant the bio-oil is combusted in a gas turbine that is connected to a generator to produce electricity. The hot flue gas from the turbine passes through a heat recovery boiler to produce steam. The steam powers a steam turbine, producing more power, and the low pressure steam from the steam turbine is sent to the sugar mill to meet the mill's steam requirement. The power section for the pyrolysis combined cycle consists of the same design as the power section shown in Figure 4 below for the gasification combined cycle concept. Cyclonc
I
Liquid recovery
Fe
Bio-oil Product I
Recycle Gas
Figure 3: Bio-oil Production by Bagasse Pyrolysis
835
The pyrolysis section of the system has a capacity of 600 t/d wet bagasse, whch corresponds to the typical capacity of wood pyrolysis plants that are foreseen by the process technology developers and that have been analyzed in previous technical economic studies4.
GASIFICATION COMBINED CYCLE Figure 4 shows a flow sheet of the bagasse gasification combined cycle concept. The gasification power section is integrated into the sugar mill in the same way as the pyrolysis concept shown in Figure 2 above. Since a gas is produced from the bagasse, the key difference from the pyrolysis case is that there can be no intermediate storage. The gas turbine must be run directly with the gasifier. Additional energy efficiencies beyond those of the gasification combined cycle alone can be gained by integrating bagasse gasification into the sugar mill. These include the use of flue gas from the existing bagasse boilers for bagasse drymg, and the use of low pressure steam from the steam turbine to supplement the sugar mill steam demand as explained above.
Figure 4: Power Production By Bagasse Gasijication Combined Cycle The gasification process is based on atmospheric pressure gasification with wet gas clean-up. The bagasse is dried to 10 % moisture in a dryer and fed to the gasifier. The gases pass through a gas clean-up system and are fed to the gas turbine to produce electricity. Steam, produced in the turbine flue gas heat recovery boiler, powers a steam turbine and the low pressure steam from the steam turbine is sent to the sugar mill to meet part of the mill’s steam requirement. The gasification combined cycle plant must be run in conjunction with the sugar mill, corresponding to the bagasse availability.
836
TECHNICAL UNCERTAINTIES Bagasse Pyrolysis Some of the companies that are developing biomass fast pyrolysis process technologies have tested bagasse in their pyrolysis pilot plants. However, the amount of pilot plant operating data and experience with bagasse pyrolysis is limited, and the largest scale of operational experience with bagasse pyrolysis is significantly less than the 25 to 30 t/d scale at which wood fast pyrolysis technology has been demonstrated. The limited experience and data means that some problems can be expected in the scale up of the technology to the 600 t/d wet bagasse feed capacity. The following are the key uncertainties related to the scale-up of bagasse fast pyrolysis processes: 0 0 0
0 0 0
Scale-up effects on liquid product yields. Feed particle size effects on liquid yields. Stability of the liquid product particularly at the high ambient temperatures present in sugar mills, including possible changes with time, temperature, water content and viscosity increase. Liquid product solids content. Pyrolysis plant safety - carbon monoxide, metal carbonyls, carcinogens. Combustion characteristics of the char by-product.
BIO-OIL AS A GAS TURBINE FUEL The use of bio-oil as fuel in gas turbines is a relatively recent development. Efforts to date indicate the feasibility of operating a gas turbine on bio-oil through combustion rig and full scale engine testing’ *. The technical challenges can be summarized by examining the key differences in fuel characteristics between biofuel and common distillate fossil fuels such as No. 2 Diesel, shown in Table 4. The utilization of bio-oil with gas turbines is acheved through the modification of current systems and processingkandling procedures to accommodate these unique properties. These are summarized in Table 4 with a detailed discussion found in reference 7. However, much of the requirements for bio-oil operation are analogous to those utilized in the combustion of heavy fuels for gas turbines. T h s therefore allows previously established technologies to be applied to bio-oil operation. This technology has been demonstrated through combustion rig and full scale gas turbine engine testing. Orenda alone has combusted over 10,000 kg of bio-oil in combustion rig tests and 15,000 kg in engine testing. These short term tests have demonstrated the applicability of bio-oil to turbine operation. Therefore, the next step, which is currently being carried out, is the development and testing of second generation bio-oil gas turbine systems for commercial level operation.
’
837
Table 4: Gas Turbine Design Modifications for Bio-oil Compared To Diesel Fuel
Gasification Combined Cycle
There is some uncertainty related to scaling up the gasification section of the plant. Bagasse has been gasified in pilot facilities, for example TPS in Swedeng,but there is no data available from larger facilities. A gas turbine has not been operated for long periods with gas derived from bagasse. However, the Sydkraft demonstration Gasification Combined Cycle plant at Varnamo, Sweden, has been succesfully operated with several biofuels for more than 8500 hours of the gasifier and 3600 hours of the gas turbine with wood gas fuel". The uncertainty related to the gas turbine operation may therefore considerd moderate. RESULTS PROCESS MODELING METHODS
A rigorous performance analysis is the key to a meaningful feasibility study. Performance analysis is essential in estimating system the costs. The design of industrial units, or more accurately, the prediction of industrial plant performance was the primary objective of the current work. Both the operating and investment costs, key elements in feasibility studies, were determined on the basis of mass and energy balances. A process analysis computer s o h a r e , AspenPlusTM,was used as the basic framework to perform material and energy balance calculations for the process cases. Aspen is a steady-state process analysis program extensively employed in chemical engineering process modelling. The performance models were used to supply: Mass and energy flows in the process
838
0
0
0
Energy demand and production in the process unitdequipment Performance of unit operations Sizing data for unit operations The data from the models are then used for: A cost and performance analysis of the process configurations T e c h c a l sensitivity studies, in which some critical process parameters are varied to study their effects on the overall performance or cost.
PERFORMANCE RESULTS The results of the material and energy balance for the bagasse pyrolysis combined cycle plant are shown in Table 5 . The annual overall efficiency is 25.2 percent. The plant would be operated in two sections: pyrolysis section and power section. The pyrolysis section consisting of the bagasse dryer and fast pyrolysis plant would be operated for 160 days per year, corresponding to the sugar mill operating period and the bagasse availability. Part of the bio-oil produced from the bagasse pyrolysis would be stored and the power section of the plant would be operated for 160 days per year with the steam at a power capacity of 6.2 MW. The benefit of the bagasse pyrolysis combined cycle concept is to allow the sugar mill to produce power for an additional 130 days after the sugar cane processing period.
Table 5: Pyrolysis Combined Cycle Power Plant: Process Performance Pyrolysis Section Bagasse feed to dryer (wet 50% moisture) Bagasse feed to pyrolysis (10% moisture)
600 t/d 333 t/d
Note 1: For the pyrolysis case the power section is run in two modes: Combined Heat and Power mode during the operation of the sugar mill to provide steam needed to supplement the mill’s steam demand, and Condensing mode during the period when the mill is not in operation. Table 6 shows the results of the material and energy balance for the bagasse gasification combined cycle plant. The gasification combined cycle plant would be
839
operated for 160 days per year corresponding to the sugar mill operating period and the bagasse availability. The overall efficiency of the plant is 38 percent. In addition to the pyrolysis and gasification cases, a conventional Rankine cycle case was analyzed for comparison. In this case the same amount of bagasse was combusted in a modem boiler with a steam turbine power cycle. The overall efficiency for this case is 25 percent.
Gasification Section Bagasse feed to dryer (wet 50% moisture) Bagasse feed to pyrolysis (10% moisture) Gas heating value, lhv Power and Heat Section Gas turbine power output Steam Turbine power output
600 tld 333 t’d 4.9 MJkg 11.6MWe 6.4 MWe 37.2 % 18.8 MWth 76.0 % 160 dyear
Power production efficiency Heat production Overall efficiency Operating period COSTS OF BAGASSE POWER PRODUCTION
The costs to produce electricity for each of the cases are compared in Table 7 and Figure 4. For each case an interest rate of 10 percent and plant operating life of 15 years were used. In Guatemala the cost for a sugar mill to purchase electricity is approximately US$O.O6/kW.
Power Production, MWe Annual Operating Period, d Power Production efficiency Cost of Electricity,
I I
Pyrolysis Combined Cvcle 6.6 29 1 28.4 % 0.085
IJS$/kWh
840
Gasification Ranlune Cycle Combined Cvcle 18.0 13.3 160 160 37.2 % 25 % 0.12
0.076
0.140
3 0.120 c Y,
5 0100 d .- 0.080 0
f om
2 0.040 0
g 0020
"0
.
m
m
1
2
pyrolysis
Gasification
3
Wne
Figure 4: Comparison of Cost of Electricity
CONCLUSIONS 1.
2.
3.
4.
5.
It appears that retrofitting a sugar mill to integrate advanced bioenergy processes for only the excess bagasse available is not competitive at current electricity costs. The situation may be different, if a power plant is built to process all of the bagasse available and integrated with the sugar mill's energy requirements. In this case, the overall energy system may be optimized both for sugar and power production, whereas now the sugar production is practically the only consideration. Such a project would almost certady have to involve both a sugar company and an utility. Further study to consider integrating a power plant to process all of the bagasse available at a sugar mill is recommended. This study would be based on the actual conditions at a specific sugar mill. At the small scale, 5 - 10 MWe, Pyrolysis combined-cycle is competitive with the Rankine cycle and could possibly be competitive with electricity produced by fossil fuels, depending on efficiencies and economic factors of a specific sugar mill. Pyrolysis combined cycle also offers the advantage of storing bagasse derived fuel and operate the power plant longer than the sugar mill operating period. At the 18 MWe capacity used in this study gasification combined cycle power plants are not competitive and will need to operate at larger capacities. Larger capacity operation of gasification combined cycle power plants are possible given the large amounts of bagasse that are available.
ACKNOWLEDGEMENTS We gratefully acknowledge the assistance of Dan Fuleki of Magellan, Orenda Aerospace Division, in providing performance data on gas turbines, and insight into the operating requirements of gas turbines with biomass derived fuels.
84 1
REFERENCES 1 2 3 4
5
6
7
8
9
10
Hall, D., House, J., Biomass: a modem and environmentally acceptable fuel. Solar En-ergy Materials and Solar Cells 38 (1995) 521-542. Factors Affecting the Commercial Potential of 'Solarwall' Technology in Drying Bagasse, Natural Resources Canada Report, Ottawa, Canada, March 2000. Ibid. Beckman, D., Pyrolysis Power Plant Using Rapid Thermal Processing and a Steam Turbine, and Bio-oil Production Plant Using Rapid Thermal Processing, Volume 3, IEA Bioenergy Techno-economic analysis activity, VTT Research Notes 1961, Espoo Finland, 1999. Wornat, M.J., Bradley, G. P. and Yang, N.Y.C, Singe Droplet Combustion of Biomass Pyrolysis Oils, Energy and Fuels, 1994, Vol. 8, No. 5, pp. 1131-1142 Jasas, G. Kaskper, J. and Trauth, R., "Gas Turbine Demonstration of PyrolysisDerived Fuels", Technical Report, DDE/ET/3333--T2, Report No. 1901, June 1983. Andrews, R.G., Fuleki, D., Zukowski S. and Patnadc, P.C., Results of Industrial Gas Turbine Tests Using a Biomass Derived Fuel, 31d biomass conference of the America's, pp. 425-435, 1997. Krumdiek, S.P. and Daily, J.W., Evaluating the Feasibility Of Biomass Pyrolysis Oil for Spray Combustion Applications, Combustion Science and Technology, Vol. 134, pp. 351-365, 1998. Waldheim, L. et al., Biomass Power Generation: Sugar Cane Bagasse and Trash, in proc. Progress in Thermochemical Biomass Conversion, September 2000. Stahl, K. et al., Final Report: Varnamo Demonstration Program, in these proceedings
842
Use of Thermo-Economic Analysis Based on Exergy Concepts to Evaluate the Cost of Electricity From Sugar Cane Bagasse in the Brazilian Sugar Cane Sector
’
S. T. Coelho’, J. R. Moreira2,D. Zylbers~tajn~ CENBIO - The National Reference Center on BiomassKJniversity of Sio Paul0 - Av. Professor Lucian0 Gualberto, 1289 - 0.5.508-900SZo Paul0 - Brazil CENBIO - BUN- Biomass Users Networkfrom Brazil ANP - The National Petroleum Agency, Brazil
ABSTRACT:Cogeneration from sugarcane bagasse (a by-product fiom sugarcane milling process) is already a reality in Brazilian sugar/alcoholplants. However, in most
cases the sugarcane bagasse is used in a very inefficient way - low-pressure boilers and low efficiency-steam turbines, almost 20 years old, despite the huge potential of electricity generation. Because bagasse production is quite high, the surplus not burned in boilers is sold to other industries. Previous studies have shown that the existing barriers against a large-scale cogeneration program are political, institutional and m a d y economic. Economic barriers are related to electricity sales in a competitive market, as happening now in Brazilian electric sector. Local utilities and cogenerators in most cases do not agree about the sale’s price and the recent legislation introduced by the Brazilian regulatory agency has not yet solved this question. Aiming to collaborate with this discussion, this paper proposes the use of the Thermoeconomic Analysis (exergy-based methodology) to evaluate the costs of steam and electricity produced in cogenerationprocess and analyses the case study of a real sugarcane mill in Silo Paul0 State. From this methodology, also electricity surplus costs are evaluated and discussed. In this study, it is proposed to develop thls cost partition using a combination of “equality” and “electricity as by-product” methods.
INTRODUCTION Cogeneration in all Brazilian sugar/alcohol plants is from sugar-cane bagasse, but almost all plants burn bagasse in this inefficient way, only to supply process needs.
843
Most Brazilian plants are self-sufficient and, in SPo Paulo State, around 10% of the plants sell an electricity surplus to the local utility but this surplus is minor when compared with state consumption (less of 1%). Eletrobras forecasts more than 4,000 MW as the potential of electricity generation from sugarcane origin, with commercially available technologies. However local utilities do not seem interested in electricity purchase from cogenerators. Until now, producers did not consider purchase prices offered attractive and existing legislation needs further revision to improve biomass-origin electricity generation. Aiming to collaborate with this discussion, this paper proposes the use of the Thermoeconomic Analysis to evaluate the electricity surplus costs, through a cost partition methodology using a combination of “equality” and “electricity as byproduct” methods. This method is applied to a real sugar/alcohol plant of Slo Paulo State, which already sells some electricity surplus and have significant potential to increase its generation, as several others in this state.
THE BRAZILIAN ALCOHOL PROGRAM Proalcool, the Brazilian Alcohol Program, presents the well-known environmental benefits, reduction on import expenditures and creation of jobs in rural areas, among other advantages. Social aspects are especially significant: current jobs in the sugaralcohol sector are one million direct jobs in rural area, more than 300,000 industrial jobs in private industrial units and sugarcane growers. Slo Paulo State is responsible for 50% of these jobs. Moreover, to create a job in sugar/alcohol industry is much cheaper than in other industrial sectors. Brazilian sugarcane production was 300 million (metic) tonnes in the 1999/00harvesting season. Sugar and alcohol production was 19 million metric tonnes and 12.7 billion litters respectively. In Sgo Paulo State, cane production was 194 million metric tonnes, corresponding to 13 million metric tomes of sugar and 8.5 billion litters of alcohol. These figures show the important role played by SPo Paulo State in this agribusiness sector. In the past, alcohol prices were not competitive when compared to gasoline prices and there were subsidies to support the difference. Two years ago the Federal Government liberated fuel prices and excluded alcohol subsidies, extinguishing also the centralised distribution system and creating a direct contact among alcohol producers and distributors. Despite the free market, the Alcohol Program is being reactivated because of its environmental and social role. Among several policies being discussed, one is especially Important: the implementation of a large-scale cogeneration program for Brazilian sugar/alcohol sector. Revenues from electricity sales could allow further reductions on alcohol production costs and there are also its positive environmental impacts due to its biomass-origin.
COGENERATION IN BRAZILIAN SUGAlUALCOHOL SECTOR Cogeneration in all Brazilian sugar/alcohol plants is from sugar-cane bagasse. Because bagasse production is quite high (50% wet-bagasse correspond to 30% of crushed 844
sugarcane) and it is burned in an inefficient process, the surplus not used is sold to other industries. Almost all plants burn bagasse in this inefficient way; 2 1 bar-boilers and back-pressure steam turbines are usually employed, with very low efficiency. Energy production is enough to supply the needs of production process: 500kg of low-pressure process steam (2.5 bar, 130 to 150’ C) per tonne of crushed cane (tc), around 20 to 25 k W t c (electric and mechanical energy). Most Brazilian plants are self-sufficient and, in Sgo Paul0 State, around 10% of the plants sell some electricity surplus to the local utility. Total installed power in the state is around 750 MW in 1999, but less than 50 MW of surplus is generated. Higher efficiencies in cogeneration processes are possible with conventional technologies already commercially available in Brazil. Eletrobris forecasts more than 4,000 MW the potential of electricity generation from sugarcane origin, with these technologies, Other existing evaluations estimate a wide range for electricity cogeneration potential in sugar/alcohol sector. From 1.O GW, with conventional technologies, while the sector’s installed power could rise up to 9-14 GW, with more advanced ones [l]. Electricity surplus is currently available only during harvesting season because, to generate electricity all over the year, it would be necessary to harvest green cane, through mechanical harvesting processes. Top and leaves of sugar cane (“barbojo”) could be stored and used for electricity generation off-season (more than 30% of harvested cane). T h s process is being implemented by some industries and cooperatives (Copersucar) [2], but there are still some difficulties regarding the best methodology for transportation of tops and leaves to the mill, as well as social problems to be solved (due to the high number of jobs in harvesting season). It must be observed that generation during the harvesting season is an advantage for local utilities because that is the dry season, when reservoirs of hydro plants present the lowest levels and around 95% of Brazilian electric installed capacity is from hydroelectric origin.However utilities do not seem too much interested in electricity purchase from cogenerators. Until now, prices offered by local utilities were not considered attractive by producers (around US$ 28/MWh, 2000 US$, for long term contracts and US$ 6,5/MWh for short-term contracts). There is not yet legislation to unprove biomassorigin electricity generation. The purchase of this energy by the utilities is not mandatory and special discounts (up to 100%) on wheeling tariffs are available only for electricity from small hydro. In February 2000, due to forecasts of deficit in Brazilian electricity supply, special policies were established by Federal Government to implement large naturalgas power plants. Government has established special prices for natural gas purchase by the plants, giving warranties of a power purchase agreement (PPA) for the electricity generated and financial support of Federal agencies to the investors. However these policies do not yet include biomass-origin electricity. Considering that the main discussion nowadays is about the prices of the electricity surplus from bagasse-origin cogeneration, this paper intends to evaluate these fijyes using Thermoeconomic Analysis concepts.
845
THERMOECONOMIC COGENERATION
ANALYSIS
OF
BAGASSE-ORIGIN
In a cogeneration plant, like a sugar and alcohol mill, the combination of the exergy analysis (combination of energy and entropy analysis of thermal systems) with cost partition methods allows rigorous evaluation of the production costs of electricity and steam. This is possible because the exergy concept [3] evaluates differently work (pure exergy) and heat. Exergy is the work obtained when a given system goes fiom a given state to a state of global equilibrium with the environment through reversible processes [3]. Following previous studies on this subject [4], this evaluation utilises the Thennoeconomic Analysis, based on the concept of exergy, to compare specific electricity and steam cogeneration costs in the chosen sugar/alcohol plant. This method was used in [4] for a theoretical plant, corresponding to an average Brazilian plant. In this paper the methodology is applied to a real (existing) plant in SZo Paul0 State, using figures obtained through local technical visits. The cost partition methods adopted are, as in [4], the “equality” and “electricity as by-product” [5] ones, with the following considerations: 0
0
“Equality” method: this method assumes that all products have the same specific (exergy-based) costs, so the investment (including steam turbine’s cost) is distributed equally into the products (steam and mechanicallelectric energy). This method is applied here for the evaluation of the base case (actual plant situation). Because the plant chosen for this study nowadays sells some surplus to the grid and some bagasse to other plants, the revenues from these sales are included in the exergy-cost balance. “Electricity as by-product” method: this method assumes that steam is the primary product and the electricity is the “by-product”. Then the steam cost must be evaluated through the best methodology available. This method is applied here when more efficient cogeneration technologies are introduced to generate electricity surplus. For this situation, it is assumed that specific steam (exergy-based) costsfor the alcoholplant are the same of the previous base case and then electricity costs are evaluated for the new configuration. After evaluating the electricity costs by this method, it assumed that specific electricity cost for the alcohol plant must also be equal to the one evaluated for the base case (actual situation mentioned above). Then, generation costs of electricity surplus are evaluated. In brief, surplus electricity is primarily responsible for the investment in the more effxient cogeneration technology. In this configuration B, investments in energy conservation in alcohol plant are assumed to reduce process steam consumption.
This methodology is adopted aiming to maintain the energy (exergy-based) costs for the alcohol plant equal to the actual ones, despite the investment to generate electricity surplus. This is important because alcohol production costs must be kept in low levels, to compete with gasoline costs, as discussed in [ 11. In both configurations, Thennoeconomic Analysis is based on the following cost rate balances from [6], [7]: 0 0
for boiler system: chps * Bbps= Cbag * Bbag+ CbS for steam turbine: c, * We + clPs * BiPs= chps*Bhps+ Cst 846
Where: Chps = specific cost of high pressure steam, on exergy basis; Bhps= exergy rate of high pressure steam; Cbag = specific bagasse-opportunity cost, on exergy basis; BbPg= exergy rate of bagasse; Cb, = capital cost for boiler system (including bagasse system and investment on energy conservation improvements; it includes also the revenue from bagasse and electricity sales, in actual situation); c, = electricity generation cost; We = generated power; clps= specific cost of low pressure steam, on exergy basis; BIp,= exergy rate of low pressure steam; C,, = capital cost of steam turbine and condenser (when needed).
THE STUDY CASE OF A REAL SUGAIUALCOHOL PLANT ENERGY FIGURES IN THE CHOSEN SUGAR/ALCOHOL MILL The chosen mill, located in Silo Paul0 State, crushed, in 1997/98 harvesting season, 5,5 million metric tonnes of sugarcane, being one of the largest sugarcane mills in Brazil. Its sugar and alcohol production was 550,000 metric tonnes and 250 million litters, respectively. During this season the mill sold 5 MW to the local utility, during 220 days (utilisation factor of 90%). No bagasse surplus was sold. Energy figures obtained during techmcal visits were as following: Steam consumption: 556 th @ 2.5 bar, 156O C Electricity consumption in sugar/alcoholprocess: 15.5 MW Mechanical energy consumption: 14.6 MW (preparation system, mills and feedwater pump are steam turbine driven) LHV (low heating value) of 50% wet bagasse = 7.44 MJkg Selling price: US$28/MWh (2000 US$) (exchange rate: R$ 1.SAJS$). Equipment: one medium pressure boiler (43 bar, 410' C, 115 th of steam, 53 t/h of bagasse); seven low-pressure boilers (22 bar, 280' C, 440 t/h of steam, 227 th of bagasse); two 43 bar-back pressure-turbo-generator sets (12 MW); four 22 bar-back pressure-turbo-generator sets (7.2 MW); 22 bar-back pressure-steam turbines for mechanical driving. Exhaust steam at 2.5 bar to the process.
EVALUATION OF ACTUAL ENERGY COSTS IN THE SUGAWALCOHOL MILL Because the 43 bar-cogeneration system is relatively new in the plant, two options (base cases) were considered when evaluating the actual energy costs for the plant:
Base case I: assuming the 43 bar system is already amortised Base case 11: assuming the 43 bar system is not amortised. In t h ~ ssituation investment is assumed according information from Brazilian equipment manufacturers. For the mentioned 43 bar existing equipment, currently commercialized in Brazil, the assumed specific investment is US$ 633/kW, according to local manufacturers. Other assumptions are as follows: Financial conditions: discount rate equal to 15% per year; amortisation period in 10 years 847
0
0
Opportunity costs for sugarcane bagasse: US$ 2 to 5 per metric tonne of 50%wet bagasse (0.27 to 0.67US$/GJ) Operation: 220 days for operation during harvesting season, 365 days for operation during season and off-season (90% of utilisation factor, UF).
Applying the “equality method” from Thermoeconomic Analysis to the equations (1) and (2), the actual specific (exergy-based) energy costs are evaluated for each of the above mentioned base cases I and 11. Two hypotheses are assumed in each case: the first one assumes specific (exergy-based) cost for boiler-feeding water equal to zero; the second one considers h s cost equal to process steam cost (exergy-based) [ 11. Results are shown in Table 1. Table 1: Energy costs for the chosen sugar/alcohol plant Base case Steam
Electrici US$/Mwh
Source: Authors ’ evaluation Notes: Process steam @ 2.5 bar, 156’ C. Bagasse: US$ 2 to 5 per m tric tonne of (50% wet bagasse); US$0.27 to 0.67/GJ. Base case 1 refers to 43-bar equipment already amortised; base case 2 refers to equipment not yet amortised. Financial conditions 15%p.y., 10 yr. Utilisation factor UF = 90%. EVALUATION OF ELECTRICITY SUGAWALCOHOL MILL
SURPLUS
COSTS
IN
THE
Among the several more efficient configurations available in Brazilian market, the following two possibilities are evaluated here, maintaining the new 43-bar system is maintained:
0
Configuration A: the existing 22 bar-boilers are assumed to be replaced by new 80 bar-boilers, assuming the same steam production (440th of steam; no investment on energy conservation in the process). Consequently the existing 22 bar back pressure turbo generators are replaced by new 80 bar back pressure turbo generators, with extraction at 22 bar (for the mechanical driving steam turbines) and exhaust steam at 2.5 bar (to the process). Energy balance for this configuration A shows an electricity generation of 43 MW by the 80 bar-turbo sets. So, when including the 12 MW already generated by the existing 43 bar system, there is a total of 55 MW produced during the harvesting season (40 MW of surplus, corresponding to a surplus of 35 kWh/tc). Specific investment for this configuration is US$ 900 per kW (electric) generated. Configuration B: similar to Configuration A but the new steam turbines are extraction-condensing units (CEST system). The existing 43-bar system (1 15 t/h of steam, 12 MW) is maintained in operation during the harvesting season and, off-season, only the 80 bar-CEST system operates. Steam extraction does not exist off-season and all steam produced by the 80 bar-boilers expands until condenser pressure, generating electricity.
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According to the energy balance for this configuration B, 80 MW of surplus is generated to be sold all over the year. Specific investment is US$ 1,500kW (electric) installed in thls situation, including investment on energy conservation to reduce steam consumption from 480 to 345 kg of process steam per metric tonne of crushed cane. Tables 2 and 3 shows the results obtained for Configurations A and B for bagasse-opportunity costs from US$2 to 5 per metric tonne of (50% wet bagasse).
Generation (Season) Electricity generated Electricity surplus US$/MWh US$/Mwh 23.65-24.72 35.14-38.82
Base case
1
Base case
1 2
Season US$/MWh 32.31-35-85 32.49-36.00
CEST-Season and Off-season Electricity generated Off season Surplus (Season) US$/MWh US$/Mwh 42.90-45.39 50.88 53.12-53.25 42.2 1-44.67
CONCLUSIONS From the obtained results, some remarks must be made. The influence of boiler feed water cost is quite reduced (less than 1%), because its specific exergy is low (conditions nearby environmental ones). However, bagasse cost is very important; Comparing base case 1 to base case 2, it can be observed the lnfluence of bagasse cost, mainly when existing system is already amortised, because in thls case bagasse is the only (operational) cost. When bagasse costs double, process energy costs triple. Off-season electricity is more expensive due to two factors: there is no cogeneration process to share the costs and biomass used in the boilers is more expensive (tops and leaves at US$ IO/dry metric tonne). Our analysis shows that both cases A - back pressure in the milling season, and case B - CEST year round 849
generation, have electricity costs that are higher than current utility electricity purchase tariffs. In fact case B offering the highest renewable energy contribution over the entire year is the higher cost, and is thus viewed as non-economic generation. These results indicate the necessity of special policies to implement cogeneration from sugarcane bagasse in Brazil. The Brazilian regulatory agency ANEEL - has introduced some mechanisms to incentivate cogeneration but further policies are necessary. For instance, ANEEL has defined the so-called Normative Values that correspond to the higher purchase tariff to be included in the final price for consumers. However, because these values are not mandatory, utilities do not seem interested. By now, there are proposals suggesting that the same special policies introduced to encourage natural gas-fired thermoelectric power plants can be adopted also for cogeneration and electricity generation from biomass. REFERENCES
Coelho, S.T. Mechanisms for the Implementation of Biomass-origin Cogeneration: A Model for the State of Sfo Paulo. Doctor Thesis (in Portuguese). Thesis advisor: Prof. Dr. David Zylbersztajn. University of SZO Paulo, 1999. Copersucar. GEF-Project “BRA/96/G31 : Energy Generation fiom Biomass, Sugarcane Bagasse and Residues”. Reports (in Portuguese) published by CENBZO NOTZCZAS, newsletter from The National Reference Center on Biomass, S2o Paulo, 1998, 1999, and 2000. Kotas, T.J. The Exergy Method of Thermal Plant Analysis. Buttenvorths, London, 1985. Coelho, S. T., Oliveira Jr., S., Zylbersztajn, D. “Thermo-economic Analysis of Electricity Cogeneration from Sugarcane Origin”. In: Third Biomass Conference of the Americas. Proceedings. Vol. 11, pp. 1631-1640, Montreal, 1997. Vertiola, S.R. and Oliveira Junior, S., “Thermo-economic analysis of the steam cycle of a Brazilian medium-sized sugar and alcohol mill”. In: XI International symposium on alcohol fuels”. Proceedings. V.2, pp. 415-422, Sun City, South Africa, 14-17 April 1996. Szargut J., Moms D. R., Steward F.R.,Exergy Analysis of Thermal, Chemical and Metallurgical Processes, Hemisphere Pub. Co., N.Y., 1988. Bejan, A., Tsatsaronis, G., Moran, M. Thermal Design and Optimisation. John Wiley & Sons, Inc., New York. 1996.
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Competitiveness Assessment of Applications of Thermochemical Biomass Conversion Technologies Maximilian Lauer' and Martin Pogoreutz2 Joanneum Research, Graz, Austria Graz University of Technology, Institute of Thermal Engineering, Graz, Austria
'
ABSTRACT A variety of technologies for the thermochemical biomass conversion is available or under development. The possibility for implementing a technology depends on the competitiveness of its specific application compared to alternatives such as fossil energy application. In t h s paper an approach to a comparative competitiveness assessment is demonstrated and results are presented for a variety of applications of biomass conversion technologies, ranging from small scale wood heating systems to biomass IGCC's and the use of pyrolysis liquids. With these results promising possibilities may be identified helping to put the right focus on RTD activities.
INTRODUCTION
In thls paper a differentiation is made between the technology and it's specific application. The realistic assessment of the chances of technologies to be implemented successfully is an important issue for putting the right focus on RTD activities. For realistic assessment of the chances the point of view of an investor has to be taken and the decisive process of the investor has to be understood. The basis for the investors decision can be divided into two groups of criteria: (1) Techno-economic criteria; Ths group of criteria describes the general applicability of a technology from a technical and an economical point of view. For the investor these criteria have to be met or the disadvantages have to be fully compensated by incentives, otherwise the decision will be negative. Technoeconomic criteria are the readiness for use and the competitiveness. (2) Non-techno-economic (other) criteria: Th~sgroup of criteria describe relevant issues for the application of a technology in a specific situation. The specific situation is characterised by social, economic, administrative and environmental boundary conditions relevant to the investor. These boundary conditions can be
85 1
very different depending on the country, the company, the specific site etc. As an example the existence of specific laws and standards, the general motivation of the company (e.g. image cultivation), hazards and risks in relation to the surroundings of the site etc. will be very lmportant for the decision of an investor. As the non-techno-economic criteria are depending on the very specific situation, they will not be discussed in this paper. As criteria generally applicable for the assessment of an application the readiness for use and the competitiveness are to be discussed at this place: (1) Readiness for use: A technology has to be ready for use, before it can be implemented. This means, that there is no substantial development work left to the investor except perhaps some adaptation to the specific application. The technology has to be proven, reliable and more or less free from technological risks. A "turn key " operation should be possible. If a technology is not ready for use (e.g. in the pilot phase or demonstration phase), the implementation will only be possible, if there are incentives existing for the investor equalising the disadvantage and the financial and the technological risk. Such incentives could be a positive outlook to the fume or financial subsidies given by authorities interested. (2) Competitiveness: It describes the attraction of the application compared to another possibility (different technology, fuel etc.) providing the same service. Normally the term competitiveness is used for specifying the economic attraction compared to other possibilities from the view of an investor.
As the assessment of the competitiveness of applications of different biomass conversion technologies is investigated in this paper, it is important to identify comparable possibilities for the investor. In most industrialised countries the investor can choose between the application of a biomass technology and a conventional alternative. The conventional alternative for the investor is normally to produce heat with a fossil oil/gas boiler and/or to buy electricity from utilities. This is specified as the standard situation in this paper.
APPROACH
Comparison of technologies are usually made by comparing the technology related data or by comparing the specific cost of the products. It will not be useful to compare the technology related data of different technologies in order to assess their competitive situation, because competitiveness depends very much on the situation technologies are used for and the technical, economical and social environment they are integrated in. Comparison of technology related data is only usefhl, if the technologies are similar and used for the same purpose in a similar size. In this case a comparison of technology related data as efficiencies, specific cost etc. can give good information. As biomass conversion systems are quite different in technology, size and purpose, an assessment of competitiveness by comparing technology related data is not possible. It seems also to be impossible to assess the competitiveness of technologies by comparing the specific cost of products as heat and power, because the products can
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have quite different values depending on the specific situation of the user. As an example a big power station can produce electricity at relatively low specific cost, but will earn only low prices for feeding power to the high voltage grid. If a small company produces electricity with a e.g. fixed bed biomass gasifier for its own needs, specific power generating cost will be much higher as the ones of the power station. But as the company has to pay retail prices for the electricity, the small gasifier in the example could be competitive from the company’s point of view. The assessment of competitiveness of technologies by comparing specific cost is applicable, if the overall situation of the applications considered is similar. If two different concepts for a power station on a specific site with similar power output are to be investigated in order to find the more competitive one, the comparison of specific cost is a very promising way. So, as the comparison of technology related data and the comparison of specific cost are not likely to describe competitiveness for different kinds of biomass conversion technologies, it seems to be necessary to choose the application of technologies as a basis for describing the competitive situation. The term application is used in this paper to describe a plant and its purpose, size, the fuel used, the annual operation time, the specific labour cost and the cost for the conventional alternatives given for.the investor. Competitiveness in this paper is described from the investors point of view as the relation of annual cost of two possible applications using different technologies but providing the same service. For easy comparison between a variety of applications a “competitiveness factor” CF is introduced using the relation: annual cost of conventional alternative CF
=
annual cost of biomass technology application As CF is a non-dimensional factor describing a cost relation, the economic competitiveness of different applications of different biomass conversion technologies can be compared to each other. The annual cost are calculated as an overall sum of cost related to investment, to operation and to consumption following the guideline VDI 2067 (1) edited by the association of German Engineers. Cost related to investment are calculated as annuity (constant annual settlement of the investment including the interest rates over a time period corresponding to the technical lifetime of the plant). Cost related to operation include personnel, service and maintenance cost. Cost related to investment and to operation are fmed cost. Cost related to consumption are variable cost and include cost for fuel, for transportation and all other cost related to the amount of consumption caused by production intensity. The annual cost of biomass application is generated using : (1) Technology related data (power range, efficiencies, specific investment, specific operation and maintenance cost, lifetime etc.). At the moment more than 40 different biomass to energy conversion technologies are listed in the technology database in the basis project (2).
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(2) Fuel related data (price, water content). At the moment 20 different fbels are integrated in the database of the basis project (2). The fuel specifications and prices are based on the Austrian situation. (3) Application related data (annual operation time, power output, labour cost, interest rate etc.). In total 46 different applications are listed in the application database of the basis project (2). The annual cost of the conventional alternative is calculated by using specific cost data for heat generation based on fossil fuels and for electric power purchased from utilities. These data are compiled for the different applications discussed depending on power demand, consumer specification (industry, utility, private) etc. DATA BASIS
The data used were mostly collected for the project “Sustainable Bioenergy Strategy for Austria” (2). Due to the limited length of this contribution, the technologies and their applications discussed in this paper represent a selection of the possibilities considered in the project “ Sustainable Bioenergy Strategy for Austria”. Data are based on the situation given in Austria in March 2000. By adapting the fuel related data and the application related data (and if necessary technology related data as specific investment cost, interest rates etc.) the calculation can easily be adapted to specific conditions in other countries. The technology related data, the fuel related data and the application related data used in this contribution for calculating the competitiveness are specified below. TECHNOLOGY RELATED DATA A wide variety of biomass conversion technologies, ranging from small scale wood heating systems to large applications for electricity production are available or under development. At the moment more than 40 different technologies and applications are listed in the technology database of the project “Sustainable Bioenergy Strategy for Austria” (2), containing a basic description of the technologies (function, power range, efficiencies), their operating parameters and conditions, state of the art and economical data (investment cost, operating cost). For this contribution a selection of technologies had to be made due to the limited length: (1) (2) (3) (4) (5) (6)
Small scale biomass boiler (room heating) for wood and pellets Small scale steam boiler Steam turbine (in combination with a steam boiler) Co-combustion Fixed bed gasification system IGCC (pressurised) (7) Steam piston engine (in combination with a steam boiler) (8) Organic Rankine Cycle (9) Pyrolysis oil used in large scale Diesel Engine In the following the selected technologies are discussed. As shown in Fig 1 functions for the specific investment cost over the power range are used. These
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functions are not reproduced in this paper for all the technologies described. For some of the technologies no function is used, as only data for one specific power output are available and used for the calculations. In these cases (Co-combustion, IGCC and Pyrolysis oil Diesel engine),.
Small-scale biomass boiler (room heating) Within the last two decades considerable development of automated room heating systems based on biomass took place in Austria, Wood-chip fired and biomass pellet fired room heating boilers have been successfully introduced to the room heating market. Wood chip boilers as well as pellet boilers automate as many functions as possible. The most significant one is fuel delivery, which is controlled by an auger or similar feed device that delivers regulated amounts of fuel from the hopper to the combustion chamber. The amount of air needed for optimum air-to-fuel ratio (optimum performance, combustion efficiency) is delivered automatically by a fan. Pellet fuel is made of wood residues which are left over from lumber production. The material is taken to a pellet mill where it is dried, compressed and formed into small, cylindrical pellets. Pellets are supplied directly to the storage device by specialised tanker trucks. The typical power range for wood c h p boilers is between 25 and 120 kW and for pellet boiler up to 40 kW. Wood chip fired room heating boilers reach an overall efficiency of up to 91% (under test conditions, practical efficiency 70-85%), pellet boilers even up to 93% under test conditions (practical efficiency 88-91%). Emission standards can easily be met by most of the systems in base and part load. Both types of room heating boilers are state of the art and comparable to oil or gas fired boilers (fully automatic operation, low maintenance needs, high efficiency in base-load and under part-load conditions, low emissions etc.). Fig I shows the dependency of specific investment cost of wood chip boilers (top) and pellet boiler (bottom) on the t h e m 1 power. Investment cost for fuel feeding system from the main storage to the hopper is not included, if used additional cost of 1.800 to 2.200 EUR per boiler have to be considered.
Small scale steam boiler Depending on the pressure and temperature of the produced hot water respectively steam two different types of boiler construction, fire-tube boiler and water-tube boiler, are distinguished. Within the following sections steam boilers of small and medium scale up to a heat capacity12 MW will be described. Fire-tube boilers consist of a water filled vessel which is crossed by the fire tubes. The hot fumes flow inside these tubes and their heat is transferred to the water respectively steam. The arrangement of superheat assemblies in the flue gas channel is basically possible. With water-tube steam generators the water vaporises inside the flue gas proof welded tubes located on the wall of the furnace. With this kind of construction notably higher steam pressure can be achieved as compared with fire-tube steam generators.
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20
30
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50
80
70
80
90
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100
IP
Thermal power, kW
I
10
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25
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35
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Thermal power, kW
Fig. 1 Specific investment cost of wood chip boiler (top) and pellet boiler (bottom).
For fire-tube boilers steam pressure typically lies within a range of 8-32 bar and a temperature between 170-450°C. Water-tube boilers of the given power range are designed for a steam pressure of 60 bar and superior with steam temperatures of up to
450°C. Both, fire-tube and water tube steam generators are technically well developed. Both systems provide a high availability, largely automated operation is possible. The approval for operation without full-time surveillance (BOSB)is state of the art. Steam Turbine Steam turbines are a well known technology for a wide range of power and applications, therefore only the most important facts for the analysis are given. Usually live steam parameters of small-scale turbines are 30 to 40 bar and temperatures about 400OC.Rotational speed is between 6,000 and 22,000 rpm which 856
makes the use of a gearbox necessary. The power range is between 500 and 5,000 kWe. Efficiency for single stage turbines is about 70%, multistage turbines reach 75 to 80% at considerable higher cost. Usually live steam parameters of big-scale turbines are up to 90 bar and temperature about 525°C.Efficiencies for big scale steam turbines reach 75 to 80%. Co-Combustion
Different technologies for the energetic use of biomass in a fossil fuel power station (co-combustion) are known. An interesting approach discussed here is a demonstration plant of Draukraft in ZeltwegIAustria where wood wastes and bark are pyrolysed and partially gasified in a fluidised bed gasifier. The gasifier produces a gas loaded with a fine dust of charcoal which is burned in the furnace of the coal-fired boiler. Since September 1999 the demonstration plant (10 MW wood gasifier co-feeding a coal fired power plant) has reached about 1.400 hours of operation. Expected performance data of the concept are in a power range of 10-100 MW thermal power, electric efficiency depending on the existing power plant between 35 and 42% and a ratio electric power to thermal power between 0.8-1.2. At the moment only the 10 MW demonstration plant is existing with data available.
Fixed-Bed Gasification Systems Different fmed-bed gasification systems are known (cocurrent, countercurrent, etc.). For the principle description of thls technology for combined heat and power production by using the producer gas in a gas engine a two-zone fixed-bed gasifier, installed at the Institute of Thermal Engineering, Graz University of Technology is taken as an example (see Fig 2). The fuel fed to the gasification reactor is converted to producer gas by a twozone gasification process. The major part of the fuel is gasified in the downdraft part of the reactor. Through the grate a secondary gasification air inlet secures complete fuel conversion. The producer gas is cooled and the gasification air preheated in a countercurrent heat-exchanger. Downstream of the cyclone for dust removal the gas is quenched and cleaned in a jet scrubber and a disintegrator. The conditioned producer gas is converted to power and heat in a Waukesha gas-engine (3). Typical lower heating values of producer gas from fixed-bed air gasifiers are given between 3.7 - 5.1 MJ/Nm3 (countercorrent) and 4.0 - 5.6 MJ/Nm3 (cocurrent and two zone). Beside the problems of gas cleaning and waste water treatment the upscaling of fixed-bed gasifiers is limited to app. 1 MWe. But even in this relatively small power range electrical efficiencies between 18 and 28%, depending on the efficiency of the IC engine and the cold gas efficiency of the gasifier, are possible. The ratio electrical power to heat output ((T)reaches 0,45 up to 0,65. Because long time experiences running a IC engine on woodgas are not available operation cost are estimated between 11 and 21 EUR/kWFa (full maintenance contract). At the moment there are about 40 gasifiers in operation in Europe, most of them for research, development or demonstration purposes. Although some of them have reached several 1000 hours of operation (with IC engines), this technology has not yet
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reached a commercial status and needs fUrther development of gas cleaning, waste water treatment and gas engines.
Fig.2 Fixed-bed gasifier with gas cooling, gas cleaning and gas engine (2). IGCC (pressurised) The technology of Integrated (fluidised bed) Gasification with a combined cycle process (gas turbine with steam turbine) is used for large power systems only (> 10 MWF). IGCC technology is represented by the pilot project of Varnamo/Sweden (see Fig 3). The preconditioned fuel (moisture content < 20%, size of particles < 8 mm) is pressurised in a lock hopper system by an inert gas (nitrogen) before being dosed via a pressurised container to the gasifier by means of a fuel screw. The bed material can be fed in together with the fuel. The gasification takes place at app. 20 bar and 950 1000°C using air as gasification medium (10% of air mass flow is extracted from the gas turbine compressor at about 10 bar, cooled and compressed and than fed to the bottom of the gasifier at 200 - 250°C). The product gas is cooled down to 350-400°C (generating saturated steam of app. 40 bar) and cleaned (dust removal in a ceramic hot gas filter), before it can be used in the gas turbine. The exhaust gas leaves the gas turbine with a temperature of about 470°C and is used in a single pressure heat recovery steam generator to generate live steam of about 40 bar and 450°C. The steam is used in a back pressure steam turbine to generate electricity and heat for a district heating system.
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Fig. 3 Simplified process diagram of the Bioflow-Concept in V h a m o (4).
Since autumn 1995 the demonstration plant of V h a m o has reached more than 4000 hours of operation, although Sydkraft AB reports the high complexity of the plant and the difficulty of achieving long, continuous periods of operation (4). With the demonstration plant of V h a m o it was shown that the biofuel-fired IGCC technology is a technically feasible option of heat and power production. Although not all of the technical problems have been solved yet, assessments of the next generation of fullscale plants have been made. Expected performance data of the Bioflow-concept are in a power range of 20-1 10 MW,electric efficiency about 32% and a ratio electric power to thermal power between 0.8-1.2. Investment cost of 1,380 EURkW, were calculated by Sydkraft for a capacity of 60 - 70 MW, power output. No estimations on operating cost are available. Steam engine
Steam engines are a well known as a reliable technology suitable for a wide range of industrial applications. Steam engines are available with power output up to 1600 kW for a single module. Live steam parameters between 6 bar up to 60 bar and 400°C and back pressure from 0.5 up to 25 bar are possible. Part load is realised through the' steam mass flow by adapting the stroke of the valve gear. To lubricate the piston motion an oil demand of app. 40 g per cylinder and hour is necessary. This oil has to be separated from the steam condensate before it is fed to the boiler again. The investment cost for a steam engine are between 1.268 EURkW,'(P, = 170 kW,) and 569 EUR/kW, (P, = 605 kW,). Operating cost including maintenance are app. 5 E m .
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Organic Rankine Cycle The Organic Rankine Cycle is similar to the cycle of a conventional steam power plant. The main difference is the working fluid, which is an organic fluid with a high molecular mass (i.e. Propane, Octane, Silicon oil, etc.) instead of waterlsteam. If the ORC is applied to biomass, it principally consists of the components (Fig.4) biomass fired thermal oil boiler, evaporator (heat exchanger thermal oiYworking fluid), turbine, optionally a regenerator to increase the efficiency, condenser (heat exchanger working fluidlocal district heating) and in order to close the cycle a feed water pump (5). Thermal Oil Cycle
Organic Rankine Cycle
I
Local District
Heating Fig. 4 Flow sheet of the ORC (left) and ORC module (right, (5)) The process is limited by the maximum temperature of the working fluid, which is between 250°C and 300°C. The evaporation pressure is about 10 bar. The turbine, which is characterised by a high efficiency (up to 85%) brings the pressure down to usually < 1 bar but the temperature is still high enough to enable district heating temperatures of 80 - 90°C. The overall electrical efficiency ranges between 12 and 15%. As the organic working fluid is easily inflammable, special precautions have to be taken (e.g. ventilated covers).
Pyrolysis oil used in large scale Diesel Engine Pyrolysis of biomass is a thermochemical conversion technology of solid biomass into a liquid (“pyrolysis oil”, “bio-oil”). This liquid can be used as a fuel with properties comparable to (crude) mineral oils. The pyrolysis process itself is not subject to this paper, the technology is described in (6), where there is also given a equation for calculating the cost of the pyrolysis oil depending on plant size and feedstock price. Pyrolysis oil in principal can be used as feedstock for processes in chemical production and as fuel for burners (similar to fossil oil burners), for special Diesel engines and for gas turbines etc. In this paper the use as fuel for Diesel engines is selected, as at the moment it seems to be the most promising way for economical reasons.
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Diesel engines for this purpose are likely to be low speed, heavy duty engines with a power output of 1 MW and more. These Diesel engines can be used for CHP production. As there are no data available up to now, all data for this technology are assumed in analogy to the data for burning fossil fuels based on information given by an engine producer (OMROD Diesels, UK). Pyrolysis oil input rate is assumed to be over 5000 kW with an electrical efficiency of 38 % and a thermal efficiency of 45 %. The specific maintenance cost are estimated to be 2 [€/kwha], consumption related operation cost 19 €/Mwh. Biomass pyrolysis technology is being developed very intensely in the last ten years and is now on the edge from the research phase to the demonstration phase. Up to now, some few tests have been done successfully with a low speed, heavy duty Diesel engine in the UK. In h s paper it is anticipated, that the further RTD work with pyrolysis oil Diesel engines will be successful. Overview on technology related data
Table I and Table 2 present a summary of the most important technology related data. Table I Summary of technical data of technologies presented. QF IMwl Pel QH pel[%] PH[%I Room heating- (wood chip) 0.03-0.13 0.025- 70-85 88 - 91 Room heating (pellets) 0.01-0.04 0.01-0.04 1 - 12 Steam boiler 1.25 - 13 80 - 92 Steam turbine (small scale) 12 - 18 20 - 62 4-33 0 . 5 - 5 2.5 - 6.5 Steam turbine (large scale) 21 - 2 5 25 - 30 20-120 4-30 5 - 30 Co-Combustion 4.3 - 53 34 - 42 43 - 53 10 - 100 3.4 - 42 Fixed-bed gasification 18 - 28 40 - 52 0.06 - 3.6 0.01 - 1 0.02-1.7 9 - 44 33 - 42 50 - 40 IGCC (pressurised) 18-110 6-46 Steam engine 0.6 - 0.4 12 - 20 70 - 73 0.8 - 8 0.1-1.6 12 - 15 75 - 82 Organic Rankine Cycle 0.5 - 10 0.03-1.4 0.4 - 8.2 Pyrolysis oil Diesel CHP 593 290 294 38 45 QF[hfw].QH [Mu7 Fuel input power (QF)and to thermal output power Qw p , [MW] ~ Electric power output pel PA] eflciencies electric and thermal
86 1
Table 2 Summary of economical data of technologies presented. Conversion Technology Room heating (wood chip) Room heating (pellets) Steam boiler Steam turbine (small scale)') Steam turbine (large scale) Co-Combustion Fixed-bed gasification IGCC (pressurised) Steam engine ') Organic Rankine Cycle Pyrolysis oil Diesel CHI'
ie [EWkW,]
200 - 1,200 2,000 - 5,000 354 - 1163 2,000 - 4,000 n.a. - 1,380 570 - 1,270 2,300 - 4,300 1,600
iF CF [EUR/~WF] [ E U R ~ W Fa] , 140 - 390 6 - 12 250 - 650 6-24 8-24 250 - 800 30 - 144 4-9 496 - 986 4 - 10 145 - 465 n.a. 560 - 1,046 11 -21 n.a. - 470 < 18 114 - 152 3-6 10- 15 345 - 516 600 2
specific investment related to electricpower &[EUR/kWj] specific investment related to fuel input power cF [E(IR/k?f$,, a] specific maintenance cost ') Investment cost without steam boiler i, /EUR/kWJ
FUEL RELATED DATA
In Austria the market for biofuels is established. Wood pellets, wood chips from forestry and sawmill residues have quite stable prices depending on supply and demand. The Price for pyrolysis oil is assumed to be the production cost calculated according (6). The price and the transport cost of the fuels used for the applications in this paper are listed in Table 3. The transport costs are calculated with a fixed and a variable component. The calculation is based on transportation by truck fiom the site of production to the site of use (storage of the biomass plant).
Table 3 Prices and transport cost for Biomass fiels.
Fuel price Biomass Wood pellets Wood chips extensive production Sawmill residue (coarse wood chips) Pyrolysis oil (production capacity 20 dth)
tYMWh 30,4 16,4 13,9 24,5
Transport cost fixed variable €IMWh €/GWhkm included included 097 22 190 30 093 19
APPLICATION RELATED DATA
.For the specification of the selected applications a variety of data has to be integrated. This data describe the situation of typical applications of biomass conversion plants in the detail needed for the calculation of the annual cost. In the following list the
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different application related data are specified and discussed. The data used for the calculations are shown in Tables 4 and 5. The kind of the biomass conversion technology (Table 4 ) used in the application gives information on the technology related data as efficiencies, specific cost etc. The nature of the investor (Table 4) is an important information in order to specify other information as personnel cost, the cost for conventional energy as an alternative to the investor etc. There are three kinds of investors considered, the private one (paying the applicable VAT), the industry (producing energy for its own use) and utility companies (producing energy for feeding into the grid and district heating systems). The purpose of the application (Table 4) is important for being able to specify the annual operation time, the personnel cost and the cost of the competing conventional energy. The specification of the plant size (Table 4 ) is important for the cost assessment. (the technology related data specify the possible size range). The specification of the fuel used (Table 4) has to be made in view of the investor, the size and the purpose of the application. In the applications selected for this paper we specify only four types of fuels in order to keep the applications comparable and to stick to the most realistic case given in Austria. The annual operating time (Table 5 ) is calculated as full load equivalent operation time per year and indicates the utilisation intensity of the plant. For room heating purposes in Austria it is relatively low (1500 Wa), for big industrial CHP plant it can be up to 7000 Wa (example biomass CHP in pulp mills in Austria). Fuel transport distance (Table 5 ) is estimated as a typical average distance considering the Austrian situation. Personnel cost (Table 5 ) is calculated by cost per hour and hours per year. Cost per hour are calculated as typical full cost for personnel for the lund of investor and the qualification of the personnel needed for operation and maintenance. Interest rates (Table 5 ) can have an important impact on cost calculation. In this paper no rise of the interest rates is included for taking into account increased risks of the investment. All applications in the example given are calculated as being state of the art. (10) Cost of conventional alternatives (Table 5 ) have been collected in spring 2000 for the situation in Austria and reflect the cost, that private and companies have to pay for producing heat with fossil oiVgas and purchasing power from the grid instead of operating a biomass application.
-
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Table 4 Specification of the selected applications.. Biomass conversion Investor technology Room heating (pellets) private Room heating (wood private chips) Fixed bed gasification industry CHP Steam piston engine industry CHP Organic Rankine Cycle industry CHP Pyrolysis oil Diesel CHP industry Co-combustion utility
IGCC CHP
utility
Steam turbine CHP
industry
Purpose
room heating room heating
Size QF Fuel used [MW] 24 Wood pellets 36 Woodchips
CHP for industry
2000
CHP for industry
4000
CHP for industry
5000
CHP for industry CHP for grid and district heating CHP for grid and district heating CHP for industry
5263 10000 10000 16000
Sawmill residues Sawmill residues Sawmill residues Pyrolysis oil Sawmill residues Sawmill residues. Sawmill residues
QF; biomass input rate in MW Table 5 Data used for further specification of selected applications. Personnel
Conventional alternative cost Biomass conversion Operating Fuel Cost timely Interest Heat El. Power technology time per transport [€/h] ear rate [€/MWh] [€/MWh] year dist-ance, [€/a] [“?/a] [Wa] [km] Room heating 1500 4 0.00 125 6 45.3 n.a (pellets) Room heating 1500 <5 0.00 125 n.a 6 45.3 (wood chips) Fixed bed 4000 10 18,17 3000 6 30,s 943 gasification CHP Steam piston engine 4000 15 18.17 8500 6 24.4 94.5 CHP Organic Rankine 5000 25 18.17 1800 6 21.8 94.5 Cycle CHP Pyrolysis oil Diesel 4000 50 15 700 6 30.5 94.5 CHP 4000 Co-combustion 25 26.16 6400 n.a 6 14.8 IGCC CHP 4000 50 23.26 36000 6 26.2 50.9 Steam turbine CHP 7000 50 21.80 32000 6 21.8 50.9 The cost of the competing conventional alternative for co-combustion is a special case in this calculation. As the gasifier produces gas for a coal power station, the competing conventional alternative is not heat or power, but the coal, that is substituted. Therefore the cost for the conventional alternative is the cost for the substituted coal.
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RESULTS In Table 7 results are given for the applications selected and specified before. The relation between overall cost of the biomass application and the cost of the conventional alternative gives the “competitiveness factor” CF as described in the chapter “Approach”.
Table 7 Results of the calculation of the “competitiveness factor” CF for the selected and specified biomass applications (figures rounded). Biomass application as specified in Tab 5 and Tab 6
Room heating (pellets) Room heating (wood chips) Fixed bed Gasification CHP Steam piston engine CHP Organic Rankine Cycle CHP Pyrolysis oil Diesel CHP Co-combustion IGCC CHP Steam turbine CHP
Heat output [kW] 21 25 1.046 2.353 3.481 2.368 9.500 44.000 35.000
Power output [kw
0 0 47 1 489 416 2.000 0 46.200 28.800
Overall Cost of conv. cost per year alternative [Euwa] [EWa] 2.800 2.700 385.400 738.000 743.500 1.040.000 1.280.000 16.300.000 20.690.000
1.493 1.710 305.400 4 14.500 576.100 I .040.000 562.000 14.005.000 15.596.000
CF 0,52 0,64 0,79 0,56 0,77 1 ,oo 0,44 0,86 0.75
CF-values <1 indicate that the biomass application is more expensive than the conventional alternative and therefore not very competitive. The hgher the CF-value is, the more competitive is the application investigated. Very low economical competitiveness is indicated for room heating boilers fuelled with wood chips (CF=0,64) and wood pellets (CF=0,52). Heating with this applications is significantly more expensive than using fossil fuels (spring 2000). But as the prices for fossil fuels are changing rapidly at the moment, the situation can change rapidly. The relatively good success of these applications on the market indicates, that at least private consumers are willing to pay a little bit more for room heating with biomass. CHP applications based on a state of the art technology show a realistic scenario. The industrial steam turbine CHP with a CF-value of 0,75 is much nearer to competitiveness as the steam piston engine CHP. (Existing Austrian industrial steam turbine CHP plants in pulp mills can use very low value or negative value biomass and therefore are competitive compared to fossil fuels). The co-combustion of biomass in coal power plants s e e m to be not competitive at the first look. In Austria a big size cocombustion different to the application investigated is analysed at the moment; first calculations show very good chances with CF-values of about 1. As these calculations are based on sensible company data, they are not integrated in tlus paper. Applications based on technologies to be further developed give a very encouraging outlook. If the development can be finished successfully, the IGCC could come very near to competitiveness. If cheap fuel can be provided or special prices for power generated by biomass can be achieved, biomass IGCC’s will be competitive in the near future. Industrial applications of fixed bed gasification CHP s e e m also to be quite near to competitiveness, if further development is successful. The same can be stated on ORC-CHP-applications. These applications are very near to a commercial
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status. Very promising for the future is the application of pyrolysis technology and the use of pyrolysis oil in Diesel engines for CHP production with a CF-value of 1. The method developed for assessing the competitiveness of biomass applications gives an excellent overview on the economical attractivity of different applications of biomass conversion technologies for investors. The method can be used for the assessment of the chances for implementation of technologies being subject to technical development and can help to put the right focus on RTD activities. REFERENCES 1.
2.
3.
4. 5.
6.
VDI- Richtlinie VDI 2067, Blatt 1 und folgende, VDI Gesellschft Technische Gebaudeausriistung 1983, Beuth Verlag, Berlin. Schaller, W. et al. (2000) Sustainable Bioenergy Strategy for Austria (‘“achhaltige Bioeenergiestrategie f i r dsterreich’y. Energy Research Association of the Assosciation of Austrian Utility Companies (Energieforschungsgemeinschafi des Verbands der Elektrizitiitswerke bsterreichs, VEb). Report on EFG project No 1.26, Vienna 2000 (to be published in September 2000. Hammerer D. & Pogoreutz M. (2000) Operation Experiences with a Two-Zone Fixed-Bed Gasijkation System for Biomass. Paper presented at the 1st World Conference and Exhibition on Biomass for Energy and Industry, Sevilla 5-9 June 2000. Stahl K. (1997) Varnamo Demonstration Plant - Construction and Commissioning.Sydkrafi AB, Malmo, Sweden, July 1997. Bini R. & Manciana E. (1999) High Eficiency Rankine Cycle for Biomass Powered Electricity Production. Prospect of the company Turboden, Brescia, Italy October 1999. Bridgwater A.V. (1 999) An Introduction to Fast Pyrolysis of Biomass for fuels and Chemicals. In: Fast Pyrolysis of Biomass; a Handbook. pp. 1-13. CPL Scientific Publishing Services Ltd. Newbury, UK, 1999.
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A Comparison of Using Wood Pellets and Fast Pyrolysis Liquid Industrially for Heat Production within Stockholm A. Ostman', E. K. Lindman2, Y. Solantausta3and D. Beckman4, Kemiinformation AB, Birkagatan 35, S-1 I 3 39 Stockholm, Sweden Birka Energi AB, Tegeluddsvagen 1, S-115 77 Stockholm, Sweden VTTEnergy, PO Box 1601, 02044 VTK Espoo, Finland 4 Zeton Inc., 5325 Harvester Road, Burlington, Ontario L7L SK4, Canada
'
ABSTRACT Wood pellets and tall oil pitch are currently used as renewable fuel for district heating in the Stockholm Metropolitan Area. Pyrolysis liquid, another biofuel, is also a potential substitute for petroleum fuel oil. A technical, economic, and environmental assessment for the whole utilisation chain from forest residues to heat has been camed out. It is concluded that development of the quality of pyrolysis liquid is necessary. Even-quality liquid is not available today, and the liquid cannot easily substitute conventional fuel oil. In that respect pellets are superior. There are a number of uncertain stages in the pyrolysis liquid utilisation chain, which needs to be addressed for the concept to reach an industrial stage. For example, it has to be demonstrated that the flue gas treatment from combustion of pyrolysis liquid is not too expensive. Preliminary results indicate that pyrolysis liquid may compete with wood pellets in heat production. The assessment from raw material to hot water and flue gases yields a small preference to pyrolysis liquid. However, this requires utilisation of the by-product steam in production stage. Otherwise pellet manufacture seems slightly more advantageous since the energy efficiency of pyrolysis is lower. INTRODUCTION Fast pyrolysis liquid may be envisioned to replace either light or heavy fuel oil. Typically these oils are used as boiler and engine power plant fuels. A successful introduction of a new liquid fuel on the energy market will be very difficult because of the low quality of the present pyrolysis liquids. The adverse properties of pyrolysis liquids as fuel are well known [l, 2, 3, 4, 51. Work is being camed out especially in Europe to improve its fuel quality. Of the pyrolysis liquid applications considered, technically closest to feasibility is to use pyrolysis liquid in existing boilers designed for heavy fuel oil [6]. This appli-
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cation appears feasible because of high taxation of heating fuels, for example, in Sweden, and may also be feasible in some other EU countries (Denmark, Italy). Birka Energi is presently using 100 000 t/a of tall oil pitch and wood pellets. The company is interested in increasing it use of renewable fuels, and is studying the competitiveness of pyrolysis liquids. As part of the International Energy Agency (IEA) Bioenergy Task, a techno-economic comparison was carried out. A preliminary test with 40 tomes of pyrolysis liquids has already been carried out in a 9 MWth boiler [7]. More tests were recently carried out in an wellinstrumented test boiler [8]. Results of these tests were encouraging. However, additional work is needed especially in handling and storage until continuous demonstration is possible. A technoeconomic assessment was carried out using base data from existing plants and activities in Sweden. Here, a process more or less under development is to be compared to existing plants. The proposed technology (pyrolysis) has several unit operations in common with the established, alternate process (pelletising). In fact, only the final unit operation has a significant difference for the alternatives. In effect, a sounder basis for comparison is assumed if actual data from existing plants are used and only the significant difference is estimated based on standard engineering practises. It should be emphasised that many figures used are average values and that the variations are large. For instance, a wood fuel (pellet) manufacturer located close to a pulp industry most certainly does not pay the raw material prices used in the assessment. Nor does a biofuel producer pay the power and steam prices when located close to a utility company (using wood fuel). In these cases, co-production is so intimate with backpressure steam, milling, etc., that it is hard to calculate even a rough figure on the specific costs. RENEWABLE FUEL ALTERNATIVESIN STOCKHOLM
Birka Energi is currently using two renewable fuels in its heat production: pellets and tall oil pitch. The alternatives are shown in Figure 1. Wood pellets are manufactured from selected mixtures of different wood residues. Raw wood is chipped and sieved before laid in raw material storage along with - for instance - sawdust. From the storage the material is milled and dried, in an optional order depending on process design and equipment. With a moisture content of 4 5 % the milled wood is fed to pelletisers for compressing to pellets of 6 - 12 (25) mm in diameter. The product is cooled in a cascade cooler and is delivered in bags or bulk. The bulk density of the product is 650 - 700 kg/m3. In pyrolysis the wood material is heated rapidly to about 500 “C at which temperature the wood decomposes to a maximum amount of liquid product. At lower temperatures more char is formed and less liquid and gas, and at higher temperatures the energy requirements are higher without producing noticeably more liquid. The pyrolysis process is carried out in a fluidised bed where milled material is fed into the bed and the product stream is condensed at temperatures between 30 and 60 “C. The char is usually separated before the condenser and used as fuel - along with the gas to provide heat to the fluidised bed. The fluidised bed may be bubbling or circulating. In both cases a “fast pyrolysis” is obtained in contrast to slow pyrolysis which usually yields lower amounts of liquids.
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Tall oil pitch, a by-product from wood pulping, is used as a reference in the study.
Fig. 1. Renewable fuel alternatives BASES FOR THE ASSESSMENT
For the pellet manufacture as well as the pyrolysis unit a raw material intake of 100.000 tons of dry substance wood per year is assumed. The amount is available within a transportation distance of 100 km in many places in Sweden. At present most of the wood fuel manufacturing is supplied with raw material from forest industries; pulp industry, saw mills, etc. Residues and waste products may be obtained in a range of prices and only a very minor part of the potential raw material from forestry is actually collected today. The total raw material cost for these materials is estimated at least 10 USDiMWh, including hauling of forest residues (thinnings and clearings, etc) in the
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forest to a roadside. The average cost for transports is 0.79 USD/ton DS km.Sawdust is locally available at prices between 10 and 13 USD/MWh in 1998 in Sweden. However, the supply of sawdust varies strongly with economic fluctuations, and sawdust is unllkely to be guaranteed in required volumes over several years. With reference to the fairly large consumption of raw material (100.000 tons of DS per year), an average price of 12.5 USD/MWh is assumed here. This cost is considerably higher than what can be obtained in Finland, Canada and USA - simply because the utilisation of wood residues in these countries is comparatively smaller and that surplus still exists. In fact, raw material (and wood fuel products) is at present exported fiom these countries to Sweden. For small, occasional amounts in Sweden, lower prices are also feasible but the present comparison in general refers to an industrial, long-term activity. As far as possible, the costs are quoted from data valid for today’s activities in Sweden. Some 20 wood pellet plants produce three hundred thousand tons of biofuel for commercial use mainly in utility companies. The reason for this is the C02 tax in Sweden laid on all fossil fuels. The total use of wood fuel in Sweden is about 1 million tons or 5 TWh, excluding the use of black liquor in the pulp industry and other similar utilisations. The issue about C02 taxation on fossil fuels in Sweden is here left out of the discussion although that, of course, forms the basis for the situation. A pyrolysis plant does not yet exist and for this unit the costs are estimated from previous evaluations [9, 101 and prestudies of proposed units. The pretreatment of the raw material - milling and d y n g - is the same or very similar to wood fuel manufacture and pyrolysis liquid manufacture; in pellet production about 10 15% moisture content is accepted whereas pyrolysis preferably should have a feed of 4 0 % moisture. Hence, the comparison of production costs basically can be reduced to the actual pellet production vs. the pyrolysis process. For the general cost level, however, the receiving and pretreatment are evaluated as well.
-
BIOFUEL PRODUCTION WOOD PELLET MANUFACTURE In Sweden, wood pellets are increasingly used as wood fuel. The pellets are more easy to handle in the transportation than the wood chips and the pellets may be crushed before firing in “powder firing”. One of the main arguments for pellets is the storage function. There are restrictions in the storage of wood chips due to degradation, freezing risks and mould problems. Further, pellets are more homogenous in quality. For smaller boilers and heaters (30 kW (in family houses) to 4 MW) pellets are considered significantly advantageous. Wood fuel production is often located in co-production with pulp industry or powerheating plants where steam from the other production is utilised in fuel production. Steam is required in the pelletising operation. In consequence, steam dryers are usually preferred and the low-pressure steam evolving fiom these may be used internally or for district heating (occasional flue gas dryers exist in plants not localised to other activities). A block flow diagram and mass flows of the pellet plant are shown in Figure 2.
870
50 000 tonnes of
wood material, 50 % moisture
Steam generation
200 000 tonnes of wood material, 50 % moisture
-
t Pretreatment
6
Low pressure steam
100 000 tonnes ds wood
Pelletizing
100 000 tomes ds pellets
Fig.2 Production of wood pellets PYROLYSIS LIQUID MANUFACTURE In the pyrolysis liquid production, steam is not used. Hence, flue gas driers are preferred and assessed in most evaluations. Flue gas driers - drum driers or fluidised beds - are somewhat cheaper in investment but require a flue gas treatment after the dryer. The flue gas treatment most likely makes the investments about equal. In addition to a scrubber, or similar device, a wastewater treatment has to be added. For comparison in this study, a steam h e r is assumed. To create analogous conditions, thls also implies that the pyrolysis liquid production is located where steam is available or at least low-pressure steam can be attributed a value. A block flow diagram and mass flows of the pyrolysis plant are shown in Figure 3.
COST ESTIMATES The cost evaluations for the total use of wood pellets and pyrolysis liquid may be summarised according to Table 1. Cost breakdown is shown in Figure 4. The total cost of using wood fuel as pellets in a 3 MW boiler for district heating is about 33 USD per MWh. The corresponding cost of pyrolysis liquid is estimated a little lower (=30 USDMWh).
87 1
200 000 tonnes of wood material, 50 % moisture
10 000 tonnes of
wood material, 50 %moisture
i
,......... ".."........-...-.. 1 Combustion i:.............T"".J
1
t Pretreatment
Low pressure steam
Flue gases separation,
70 000 tonnes pyrolysis liquid
Fig.3 Production of pyrolysis liquid. Table 1. Cost comparisons for wood pellets and pyrolysis fiom raw material to district heating in a 3 MW boiler. Costs in USD/MWh. All costs not included.
Cost item Production costs (not complete) Costs for transportation,
Wood fuel (pellets) Pyrolysis liquid Tall oil pitch 22
25
4
2
2
7.5
3.5
3.5
33.5
30.5
30.5
e
25')
200 lan
Costs in the boiler (retrofitting, reconstruction,. etc). Total ') Price (not cost)
For comparison, fuel oil costs were some 40 USDMWh in Sweden in 1998/98 for small consumers and some 30 - 40 USD for slightly bigger consumers, 0.1-1 MW. For the small boilers concerned - located in the middle of housing areas - only high quality oil is practically feasible at costs of this magnitude. At present hardly any of pellet producers pays the cost of 12.5 USDNWh for wood fuel, but rather 8 - 10 USDMWh if purchased externally. Consequently, the costs of established pellet production is lower than is estimated in Table 1.
-
872
Prices as of August 1999 in Sweden
Fig. 4 Production cost breakdown, comparison of renewable fuel alternatives. The difference between the costs of using wood pellets and the tentative costs of using pyrolysis liquid is some lo%, which in general is well withm the accuracy that can be expected. Thefore, no significant difference exists between these alternatives. It must be emphasized that the equivalence in costs is based on the assumption that the pyrolysis unit as well as the pellet production is equipped with a steam dryer enabling sales of low-pressure steam. In practice, this is today the situation for several pellet manufacturers, whereas none of the pyrolysis processes has been evaluated with tlus condition. Hence, if the present pellet production were compared to proposed pyrolysis liquid production, the pellets would be advantageous by some 5 - 10 USDMWh (pyrolysis liquid some 5 USD more expensive and the present pellets 5 10 USD cheaper). There are some reasons why wood pellets could be less costly: - The investment costs used in the study are based on actual figures for the production of pellets - although with large deviations - whereas the pyrolysis plant investment is a theoretical estimate, and these usually tend to underestimate the real costs. - Further, an overview of the key process steps indicates that the pyrolysis unit is more complex than the pelletisers. The other parts of the complete processes are alike. In the investment cost estimate this is valued by about 100% more for the pyrolysis - which, as commented before, seems a slightly small factor. We have no reason to question the estimate and cannot produce any more reliable data but at least the engineering part should imply a bit larger difference. - Any upgrading of pyrolysis liquid - to improve is applicability - will cost more money.
873
CONCLUSIONS
In the future, several factors may change. Most likely, further development of pyrolysis liquid will improve the process as well as the properties of the product. Secondly, the costs of firing both these hels in boilers will probably decrease. This may not alter the relation between the fuels. Thirdly, cheaper raw materials may appear. This will not affect the relation much, since the raw material share of the production costs is about the same. It might be easier to handle low-value raw materials in the pyrolysis as this includes a chemical breakdown of the material, Pelletising is to some extent sensitive to the structure of the wood. On the other hand, experiences have shown large variations in pyrolysis liquid depending on raw material. The overall conclusion is that pyrolysis liquid may be produced and fired at the same cost level as pellets from wood. At present though, development work on pyrolysis liquid has to secure an acceptable and even quality of the upgraded wood fuel as far as that of pellets. REFERENCES
Elliott, D. C. (1988). IEA co-operative Project D1, Biomass Liquefaction Test Facility Project, Volume 4: Analysis and upgrading of biomass liquefaction products. DOE/NBM--1062 Vol. 4. Springfield, Virginia: National Technical Information Service. 2 McKinley, J. (1988). Biomass liquefaction: Centralized analysis, final report. DSS File No. 232-4-6192. Ottawa: Energy, Mines and Resources Ministry. Rick, F. & Vix, U. (1990). Product standards for pyrolysis products for use as 3 fuel in industrial fuing plants. In: Bridgwater, A. & Grassi, G. (eds.). Biomass pyrolysis liquids upgrading and utilisation. Barking: Elsevier Applied Science. Pp. 177-218. 4 Fagems, L. (1995). Chemical and physical characterisation of biomass-based pyrolysis oils. Literature review. Espoo: VTT. 113 p. (VTT Research Notes 1706). Oasmaa, A. et al. (1997). Physical characterisation of biomass-based pyrolysis 5 liquids. Application of standard fuel oil analyses. Gasification and pyrolysis of biomass. An international conference, Stuttgart, 9-1 1 April, 1997. 6 Huffmann, D. (1997). RTP bio-crude: A combustiodemissions review. In: Bridgwater, A. V. & Boocock, D. (eds.). Developments in thermochemical biomass conversion. London: lBackie Academic & Professional. Pp. 489 - 494. 7 A confidential report by Birka Energi, 1996. 8 Oasmaa, A. et al. this conference Solantausta, Y., Bridgwater, A. & Beckman, D. (1996). Electricity production 9 by advanced biomass power systems. Espoo: Technical Research Centre of Finland. 120 p.+ app. 61 p. (VTT Research Notes 1729). 10 Solantausta, Y., Koljonen, T., Podesser, E., Beckman, D., Overend, R. (1999). Feasibility studies on selected bioenergy concepts producing electricity, heat, and liquid fuel. Espoo: VTT. 46 p. (VTT Research Notes 1961). 1
874
Development of catalytic wood fired boiler: Integration, deactivation and regeneration of netbased catalysts
f"l. Berg', T. Hargitai2,J. Brandin2and N. Berge'
TPS Termiska Processer AB, Studsvik, S-61I 82 Nykoping, Sweden Katator AB, Ideon Research Park, S-223 70 Lund, Sweden
ABSTRACT: Catalytic oxidation of unburned components in the flue gas from domestic wood fired boilers using net-based catalysts instead of the traditionally used monolithic catalysts is addressed in t h s paper. With the net-based catalyst and the described integration of the catalyst in the boiler a CO-conversion of 80% to 90% was achieved depending on the number of nets used. The achieved conversion as a function of the number of nets is compared with a previously developed computer model describing the catalytic oxidation and mass transfer limitations over the net-based catalyst and a good agreement is acheved between measured and calculated results. It is also experimentally shown that the deactivation of the net-based catalyst is a severe problem but that full regeneration of the catalyst by washing in acidic solution is possible. Finally a method to slow down the catalyst deactivation by using protective layers of e.g. gypsum is presented and experimentallyevaluated. These results and results from baking the net-based catalysts in wood ash and model substances supports the theory that potassium (present mainly in the form of K2C03)and possibly also other salts migrates from the ash to the catalyst where the pores of the catalyst are blocked.
INTRODUCTION Residential combustion of biohels, characterised by batch wise firing and unstable combustion conditions, has been identified as a major source for the emissions of VOC (Volatile Organic Carbons) and PAH (PolyAromatic Hydrocarbons). This, being an environmental as well as health problem, has resulted in a growing demand for domestic appliances with lower emissions. One interesting alternative, already commercially available for wood fired room heaters, is the integration of oxidation catalysts. The potential for catalytic abatement to decrease the emissions from wood fired boilers has previously been shown [l] and more than 80% reduction of the CO emissions could be reached over the catalyst [2]. This was achieved by integrating a monolithic catalyst in a modified boiler. A numerical model has also been developed by whch the conversion over monolithic catalysts can be calculated for different catalyst configurations and inlet conditions and a good agreement between predicted and 875
measured results was achieved [3]. The effect of the monolibc catalyst on the different hydrocarbons has also been shown in a previous paper [4]. T h s work has now reached the stage of field-testing and a number of commercial wood boilers, equipped with monolithc catalysts, will be evaluated in field trials over the corning firing seasons. In this paper the use of net-based catalysts instead of the traditional monolithic catalysts are evaluated. The net-based catalysts, with their pros and cons compared to monolithic catalysts, have been evaluated for this specific application in laboratory scale experiments [5] but have previously not been used for longer periods in wood fired boilers.
OBJECTIVE
The overall objective of the project was to develop catalytic wood fired boilers with low emissions of unburned components. In this specific work the objective was to evaluate the possibility to use net-based catalysts instead of the more commonly applied monolithic catalysts. The conversion over the catalyst when integrated in the wood fired boiler is measured and compared with calculated values and the deactivation, regeneration and possibilities to prolong the lifetime of the catalyst are shown. EXPERIMENTAL THE WOOD FIRED BOILER AND TEST CONDITIONS USED A commercial wood fired boiler with a nominal output of 35 kWm, (Heimax, supplied by BENTONE Al3) designed with down draught combustion, has been used throughout the experimental work described in ths paper. The boiler is of forced draught type and a single fan is used for the supply of both primary and secondary air. The primary air is preheated once the combustion process has reached a sufficiently high temperature since the air is used to cool the lid of the fuel storage. The primary air is supplied in the top of the fuel storage, which has dimensions 400*580*650 rnm.In the bottom of the fuel storage there is a ceramic grate with two parallel slots where flames and combustible gases leave the primary combustion zone. In these slots the secondary air is supplied and mixed with the gases. The boiler has been slightly modified from the original design, mainly by the integration of a baMe plate for the distribution of the primary air and a second baffle plate installed underneath the grate. These modifications and the positive effect they had on the emissions are further described in a separate paper [2]. All results presented in this paper are from the boiler after these modifications had been performed.
FUEL QUALITY AND FUEL FEEDING
Throughout the tests, well-seasoned birch wood was used as fuel. The moisture content was around 11%. The fuel was cut into 470 mm long logs and split so that the maximum diameter was 100 rnm. In each experiment the boiler was started from cold conditions, i.e. water temperature <30°C.When the water temperature in the boiler reached 80°C the thermal valve to the accumulator tank opened and the water temperature in the boiler was 876
thereafter kept at thls temperature. This enabled operation at nominal output throughout the whole combustion cycle, thereby avoiding periods of part load operation with subsequent reduction in air supply. The boiler was always ignited with 4 kg of finely split wood and during the first period the by-pass damper was left open to establish sufficient draught out of the boiler and in the chimney. After approximately five minutes the by-pass damper was closed and the flue gases were forced through the fuel bed and into the secondary combustion chamber, i.e. the operation of the boiler was shifted to down draught combustion. Ten minutes after the damper was closed, a bed of glowing char was formed and the first main batch of wood, 2 1 kg, was put on top of the glowing char. During fuelling the bypass damper must be opened but when the hel-feeding lid is closed the damper is also closed. When most of the wood is burnt out and the C 0 2 concentration reaches 4% the boiler is refuelled with 13 kg of wood placed on top of the glowing char bed remaining from the first batch of wood. This test procedure is commonly accepted in Sweden.
CATALYST INTEGRA TION The catalyst must be positioned where its operating temperature is reached quickly but long term exposure to high temperatures, >900°C,must be avoided. The position where suitable conditions can be met is at the inlet of the convection section. A schematic drawing of the boiler is shown in Figure 1 after the described modifications and the integration of the catalyst.
Figure 1
Drawing of the boiler with catalyst
The back end of the boiler was also modified to make room for a catalyst with a larger cross section area. After the modification the cross-sectional area used for the catalyst was 275*177 mm. As part of the convection section was used for this modification there was an increase in the flue gas temperature in the chimney and thus the boiler efficiency dropped. In a commercial boiler h s would have to be 877
compensated for but this boiler was regarded as laboratory equipment used for catalyst testing and not as a commercial product. FLUE GAS FLOW The burning rate (kg fuel/%)is not constant over the combustion cycle. However, since a fan is used for the air supply to the boiler the actual flue gas flow does not vary too much with time. Based on calculations from the total fuel consumption and oxygen content in the flue gas and on measurements using a Prandtl-tube in the chimney, the average flue gas flow was estimated to be 100 m3,/h. GAS ANALYSIS The concentrations of 02,C02, CO, NO and NO and THC were measured in two positions in the boiler, 0.1 m upstream and 1.5 m downstream of the catalyst. Oxygen was measured with a paramagnetic instrument (M&C PMA25), CO and C02 with IR (Leybold-Heraeus Binos 1.2 and Rosemount NGA 2000), NO and NO2 with WAR (Rosemount NGA 2000) and THC using a FID (Flame Ionisation Detector, Rosemount Termo-FID). The FID was calibrated using methane. The sampling point upstream of the catalyst was directly underneath the d e t of the catalyst, i.e. in the position where the flow of the flue gas shifts upwards. The gas is withdrawn through the back wall of the boiler and the reactions rapidly quenched by cooling. The sampling position downstream of the catalyst is-in the lower part of the chimney, i.e. after the convection section where the flue gas temperature is low. The gas from each sampling position passed through a heated particle filter (200°C) before being dried in a cooled trap (20°C). The gas was filtered at this temperature and distributed to the different analytical instruments. CA TALYST PREPARA TION The catalyst was prepared in accordance to a procedure developed at Katator [6].It was a 6 Mesh, 0.9 mm thread diameter, Kanthal net with a porous layer of thermally sprayed alumina 0.1 mm thick, thus resulting in openings of 3*3 mm. As active components a mixture (molar ratio 4:l) of Pd and Pt salt was used for the impregnation. Typically a fresh net catalyst has a specific surface area of 3.5 m2/g. The specific surface area might appear low, but this is because most of the weight originates from the metal gauge, the surface area per net is around 10 000 m2/m2. The catalyst was fitted into a cassette, 180x280 mm and 100 mm high, up to 24 pieces of net (0.61 m2) was fitted in this cassette. REGENERATION PROCEDURE The catalyst activity is reduced during long term operation due to capture of alkali, mainly potassium. To regenerate a deactivated net catalyst it is sufficient to wash the net in a water solution of citric acid, equal to pH 1.The nets are soaked in the solution for 5 minutes, left to drain and dried at 50°C in a hot air oven.
878
PREPAM TION OF CATALYST WITH PROLONGED LIFETIME
To improve the resistance of the catalyst towards poisoning and deactivation the fresh catalyst was immersed into a slurry of CaS04 x % H20, 30 weight-% in H 2 0 . The nets were blown clean with compressed air, and then left to harden. This procedure was repeated until 130 g/m2 of gypsum were deposited on the nets.
RESULTS CONVERSION OVER NET-BASED CATALYST AS A FUNCTION OF NUMBER OF NETS.
The conversion of CO was measured with fresh catalyst, over 6, 12 and 24 pieces of nets fitted into the cassette, in a temperature range of 600-7OO0C, with an average flow of 100 Nm3/h. At those temperatures, the catalyst is in the outer mass transport region, the activity depends only on the rate of transport of reactants from the gas bulk to the surface of catalyst. The conversion is therefore only slightly dependent on variations in temperature. In Figure 2 the conversion of CO is plotted versus the number of nets, together with simulated data, and G*,,= 100 Nm3/h, according to the model for netbased catalysts developed by Fredrik Silversand [6]. 100
ao 0 60 0
c
.-0
$>
40
c
s
20
0
2 00
3 00
4 00
5 00
6 00
7 00
temperature "C Figure 2 Measured and simulated conversion of CO versus the number of nets (1 80x280 mm) Gflow=100 Nm3/h. DEACTIVATION OF CATALYST
The catalyst starts to deactivate immediately when it is exposed to the flue gases. The activity has an exponential decline versus the numbers of combustion cycles, Figure 3. The conversion falls from 80 % down to 40 %, after 17 cycles, were it seems to level
879
out. After 17 combustion cycles the catalyst was removed and regenerated with citric acid solution, then it was re-installed and the measurement continued (Figure 3). 100
I
I
I
I
I
80 -
90
.--\
$ 70-
-
0 0
60-
2
50-
9 C
40
-
'c
.-0
-
-
N
-
N
-
30-
4-
10
0 .
I
-
Regenerated I
I
1
1
Numbers of Combustion Cycles
Figure 3 The conversion of CO over 12 nets, fresh and regenerated net-catalyst, versus the number of combustion cycles.
The deactivation is caused by loss of active surface, which can be restored by washing in citric acid solution. The results of a pore distribution measurement, measured by the desorption isotherm of nitrogen in a Micromeretics Tristar physisorption equipment, is given in Figure 4, as the incremental pore surface area for deactivated and regenerated catalyst (same sample). After the regeneration, the specific surface area and the original activity (Figure 3) are restored. 2.2
+I ; I
2.0-
-=-
-+-
I
Deactivated in Boiler Regenerated
1.4
Q)
6
a
0.8
0.4
0.2 0.0 10
100
1000
Pore diameter A
Figure 4 Change in the incremental specific surface between deactivated and regenerated net-catalyst (same sample).
880
GYPSUM AS PROTECTIVE COATING
12 gypsum-coated nets were prepared for tests in the boiler (Figure 5 ) . By the use of gypsum as a protective coating, the rate of deactivation slows down.
l90
o
o
m
J
80
-
-
70-
'c
-
50-
.-
0
9 C
0
-
40-
30-
10 -
0
-
Uncoated Fresh Catalyst Gypsum coated Catalyst I
I
I
Number of Combustion Cycles Figure 5 Comparison of deactivation rate between uncoated and gypsum coated net catalysts (12 nets). Gflow=100 Nm3/h, temperature range 600-700°C
DISCUSSION Catalytic combustion of CO is a fast reaction. At normal operation the catalyst is in the outer mass transport controlled region, this means that the reaction rate of the catalyst is higher than the rate of transport of reactants from the gas bulk to the outer surface of the catalyst. The apparent activity is then affected by parameters like geometry of catalyst, flow rate and properties of the gas, like viscosity and density. A change in temperature effects the properties of the gas and the flow rate, it therefore effects the conversion slightly, however it is not the change in intrinsic reaction rate that can be seen. This phenomenon is the same for all types of geometric shapes of catalysts, monoliths, pellets or nets. In a monolith the mass transfer, between gas bulk and the outer surface of catalyst, is not particular good since the flow, at least in the boundary layers, tends to become laminar, but the pressure drop is low. In a packed bed with pellets the mass transfer is very good in general, but the pressure drop is high. A stack of nets, however, combine good mass transfer and low pressure drop, it is in between a monolith and a packed bed. The measured conversion and the simulated conversion as hnction of the number of nets are in fair agreement. Two of the data sets are close to the simulated curves and one is a bit off, but there are of course spreading in the result. This confirms the validity of the model developed by Fredrik Silversand [6]. This model takes into account outer
88 1
mass transfer by calculation of mass transfer constants from geometry data of the nets, inner mass transfer by difhsion and kinetics of the reaction. When the catalyst is exposed to the flue gas in a biolkel-fired boiler, it starts to deactivate immediately. The conversion (figure 3) declines exponentially from 80% down to about 40 %, where it seems to levels out, after about 17 combustion cycles. The reason for the deactivation is a loss of specific surface area. Washing in a water solution of citric acid can restore this area and the activity (figure 3 and 4). The catalyst cannot be restored by thermal treatment, at least not by heating to 700-800°C.This means that the poison is non-organic and easily soluble in slightly acidic solution compound, probably a salt. BAKING EXPERIMENTS
In earlier work the results of poising the same type of catalysts by baking has been reported [7].Since this report only is available in Swedish the results and conclusions fiom that work is summarised below. Baking means that the catalyst is placed embedded in the substance suspected for causing the deactivation, wood ash for instance, at high temperature (500°C). At certain time intervals, the catalyst is removed and the activity is measured as function of baking time. Baking as a method for deactivation can be criticised since the catalyst is not exposed to the gas components in the flue gas and the composition of the ash can change by heating in air under long time. However, the deactivation show a similar behaviour, both catalysts deactivated in real flue gases and catalysts baked in wood ash in the laboratory, loses specific surface area. Washing in citric acid solution can also regenerate both catalysts. The time scale for the deactivation, in real life and in the laboratory, differs though. Wood ash, salt extracted from the wood ash and the residual washed ash, K2C03 (pa), KCI (pa) and K2S04 (pa) was experimentally investigated (Figure 6). 100
90
80
0
2 0
70 60
1
1
Temperature "C Figure 6 Conversion of CO versus temperature for net catalysts deactivated by baking at 500°C for 24 h.
882
Initially it was concluded that balung with wood ash deactivated the catalyst severely. The contact between the catalyst and the ash is important, if the catalyst is placed some millimetres above the ash during the heating the catalyst becomes unaffected. This means that the deactivation during the balung is not caused by any gas phase component released by the ash. Secondly, the salt extracted from the ash deactivates the catalyst, but not the remaining washed ash. The main component in the ash salt is KZCO3, with smaller amount of Na, Mg, Si and SO:-. Baking with pure salts (pro-analysis) showed sever deactivation with K2C03, some deactivation with KC1 and none with &SO4. The difference between the ash salt and the K2CO3 is that the ash salt is contaminated with other salts. This influences the melting point, it is lower than the melting point of pure K2CO3 (mp 891°C). The grains of powdered ash salt sinter together to a cake (the grains can still be seen) in 24 h when heated to 500°C. If some of the ash salt is placed on a microscope slide of glass and heated to 500"C,a spot can be seen on the glass when the salt is brushed away after 24 h. The salt wets or reacts with the glass, although the salt is not melted. Neither of the pure salts exhlbits this behaviour, but when 1 weight% of sodium carbonate is mixed in the potassium carbonate, a considerable wetting of the glass occurs. The conclusion is that there is a considerable mobility of the ions in the salt, they migrate from the salt bulk to the surface of the catalyst, and thus the salt is far from melted. In the presence of other salts a eutectic mixture is formed which significantly reduce the initial melting point of the mixture and increases this effect. There must however be other effects than melting as KCl with a melting point of 770"C, deactivate the catalyst somewhat, but not as mush as K2CO3 (mp 891OC). T h s might depend on the interaction between the salt and surface of catalyst, the surface tension for instance, or some conversion that K2CO3 undergoes in higher extent than KC1 during the heating. The presented hypothesis is that the deactivation of the catalyst, during balung, is caused by migration of salt from the ash to the surface of the catalyst, where it causes bloclung of the pores, and subsequently loss of activity. To avoid or at least slow down the process of deactivation, it might be possible to collect the mobile potassium ions in a collector substance. To catch the ions, the melting point of the formed substance should be high, a suitable candidate is K2SO4 with melting point of 1069°C for pure salt, and the potassium sulphate did not show any deactivation of the catalyst during bakmg either. A thermodynamic calculation of the stability of K2CO3 and CaS04 at the actual temperatures:
shows that the equilibrium is completely shifted towards the right; towards K2S04 and CaC03. T h ~ would s give a thermodynamic fundament for the use of gypsum as collector substance for potassium, however it says nothmg about the rate of collection. It may be to slow to be useful, or another, for the system more preferable reaction, may occur instead. Baking in wood ash, with gypsum coated net-catalysts showed good results (Figure 7).
883
"j
-
,
-O-Gypsumcoateddeac(ivated
180 h
10
2 00
2 50
3 00
350
400
4 50
500
Temperature "C Figure 7 Comparison of the deactivation for uncoated and gypsum coated net catalyst by baking in 24 hours at 500°C. The gypsum coating had significant effect on the deactivation during baking. To test the hypothesis that gypsum had some special effect, also coatings with alumina and silica were tested. Those coatings did not slow down or protect the catalyst fiom deactivation. GYPSUM COATED NETS IN THE BOILER
The results fiom the tests in the boiler of the gypsum-coated nets are shown in figure 5 . Compared with the uncoated catalyst, the gypsum coated catalyst shows as slower deactivation rate. The conversion also seams to level out at a higher level. This means that the gypsum coating protects the catalyst in real flue gases as well as in the baking of catalysts in the laboratory. There are, however, problems in connection with using gypsum as coating. The reason for gypsum hardens and stay together is the formation of dihydrate (CaS04 x 2 H20).When the gypsum is heated, it loses water, and the firmness is lost. DEACTIVATION OF SCR-CATALYSTS IN BIOMASS FIRED BOILERS
Deactivation is not only a problem for oxidation catalysts in combustion of bio-fuel, also SCR catalysts used in biomass fired boilers are deactivated. Andersson et al. [8] investigated the deactivation of SCR catalysts in a couple of different large scale biofuelled boilers. The SCR catalyst works at lower temperature (not above 400°C). According to the authors, no loss of surface area occurs and the deactivation is explained by adsorption of gaseous potassium on acidic site of the catalyst. The catalyst can only be partly regenerated by washing in acidic solution. The reason for this difference between precious metal based oxidation catalyst and vanadium pentaoxide
884
based SCR catalyst is that alkali, potassium for instance, is a true catalyst poison for the SCR catalyst. Alkali destroys the active vanadium pentaoxide phase by a chemical reaction. The precious metal catalyst is not chemical sensitive towards alkali, in our case it is a physical blocking of the surface. This is also the reason why the oxidation catalyst can be fully restored by the washing procedure.
CONCLUSIONS In this presentation the potential for net-based catalysts for the oxidation of unburned components in the flue gas from small-scale wood combustion has been reported. It is clear that a high initial conversion of CO over the catalyst can be achieved but that a rapid deactivation results in decreased conversion already after a couple of weeks of operation. It has also been shown that the model for simulation of the conversion over netbased catalyst, previously developed by Silversand [6], is in fair agreement with experimental data. The deactivation by bakmg of the net-based catalyst is caused by salt contents of the ash. The salt contains malnly K2C03but also minor amounts of other salts that lowers the melting point. The ions of the salt have considerable mobility thus the salt is far from melted and they migrate from the ash to the surface of the catalyst where the pores of the catalyst are blocked. This causes the loss of activity and the measured catalyst deactivation. Washing in an acidic solution that solves the salt and restores the surface, for instance citric acid, can regenerate the catalyst.. By coating with a collector substance, for instance gypsum, the migration of salt can be avoided or at least retarded. Tests in the boiler show that the gypsum-coated catalyst exhibits hlgher resistance to deactivation than uncoated catalyst. This result supports that the migration of salt from the ash to the catalyst, either alone or in combination with other deactivation mechanisms, causes the deactivation in real flue gases from combustion of bio-fuel.
ACKNOWLEDGEMENT
Th~swork has been performed in collaboration between TPS Termiska Processer AB and Katator AB within the national program “Sdskalig forb,ranning av biobrhle” under contract no P9744-2 and contract no P9675-2. The support from STEM, the Swedish National Energy Adrmnstration, is gratefully acknowledged. REFERENCES 1. Camo J., Berg M. and JarPs S . (1996) Catalytic abatement of emissions from smallscale combustion of wood. A comparison of the catalytic effect in model and real flue gases. Fuel, 75(8) pp. 959-965. 2. Berg M. and Berge N. (1999) Development of domestic wood fired boilers with catalytic abatement of emissions. Proceedings of the 2”dOlle Lindstrom Symposium on Renewable Energy - Bioenergy, pp. 33-39.
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3. Wanker R., Berg M., Raupenstrauch H. and Staudinger G. Numerical Simulation of Monolithic Catalysts with a Heterogeneous Model and Comparison with Experimental Results from a Wood-fired Domestic Boiler. Accepted for publication in Chemical Engineering Journal. 4. Berg M., Rudling L. and Berge N. (2000) Abatement of emissions from domestic wood fired boilers by catalytic flue gas cleaning. Proceedings of the ECSBT2, pp. 199-207. 5. Hargitai T., Silversand F. and Janner J. (1997) Katalytiska metoder for begransning av skadliga utslapp frh f o r b r e g av biobransle, etapp I. Energimyndigheten rapport 471-1997, (In Swedish). 6. Silversand F. Catalytic Combustion In Environmental Protection and Energy Production, Doctoral Dissertation 1996, University of Lund, Department of Chemical Engineering 11. 7. Hargitai T., Silversand F. and Brandin J. Katalytiska metoder for begransning av skadliga utsliipp frh forbranning av biobrwle, etapp 11. Energimpdigheten rapport June 2000, (In Swedish, under preparation). 8. Anderson C., Odenbrand I. and Anderson L. H. (1998) SCR vid Biobransleeldning. Varmeforsk rapport 646, (In Swedish).
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Combustion of chlorine-containing biomass: V205-W03-Ti02 monoliths for Cl-VOCs abatement in the flue exit gas. Part I J.Corella, J. M. Toledo, M. GutiCrrez Department of Chemical Engineering, University 'Complutense' of Madrid, 28040 Madrid, Spain
ABSTRACT Eleven catalysts for DeNOx and dioxins abatement from eight differet manufacturers have been tested for total oxidation of chlorinated VOCs generated in waste incinerators. These catalysts are based on vanadia (Vz05)-tungstena(W03) supported on titania (TiOz).They are tested in monolithic shape of around 2 x 2 x 20 cm, with gas hourly space-velocities of 2,000 to 10,OOOh" (n.c.) and 100 to 1,000 ppm Cl-VOC at the reactor inlet. C1-VOCs used were chloroethane, methylene chloride (CH2C12),trichloroethylene and mono- and di-chlorobenzenes. A lot of conversion-temperature curves are presented in this paper for very different experimental conditions. The V-W-Ti catalysts prove to be, by average and under the same exp. conditions, ten times more actives than the competitive chromia, Pt or Pd based oxidation catalysts. INTRODUCTION Combustion of biomass containing important (around 1 wt%) amounts of chlorine, such as municipal solid waste or refuse derived fuel and some industrial residues, sludges and waste is hmdered by the generation and emissions of harmful chlorinated dibenzo-dioxins, fkans and of several other chlorinated polyaromatics (CI-VOCs), besides of pollutants like VOCs, PAHs, trace metals, NOx, etc... A promising method or process for dioxins and C1-VOCs abatement is their total oxidation to C 0 2 and HC1 by using oxidation catalysts downstream from the combustor or incinerator [l-51. To this end, good potential catalysts for total oxidation of C1-VOCs need to be identified. Such catalysts should have a hgh activity and selectivity (towards C02 + HCI) at low temperatures, no-foxmation of dangerous or toxic by-products, low cost and high life time. There are several types of catalysts which could be useful for this application. Among them, results obtained with chromia and with Pt and Pd-based catalysts have been already reported [refs. 6 and 7, respectively]. T h ~ swork is devoted only to the effectiveness of Vanadia (Vz05)- Tungstena (W03) catalysts supported on Titania
887
(TiOz). They are usually designed and used for deNOx applications [8-121 but they also have proved some usefulness for dioxins abatement [2,3, 131. Although there are a lot of papers and technical brochures on using V-W-Ti catalysts for deNOx, there is very little detailed scientific information on catalytic deDioxins and/or Cl-VOCs abatement.The main objective of this work is thus to study the potential use of these V-W-Ti catalysts in total oxidation of C1-VOCs. Simultaneous deNOx activity is absolutely out of the scope of this paper. The first step will be to study these catalysts with some targeted Cl-VOCs at small (lab) scale (part I). Under the conditions in which these catalysts prove to be useful, they will be furhter tested, in a second step (part I1 of this work) in a small pilot plant based on a fluidized bed incinerator.
FACILITIES AND REACTORS USED Two very different facilities have been used. The work reported in thls paper (part I) has been carried out in the lab. scale facility shown in Figure 1. The targeted C1VOCs used have been ethyl chlorure (C2H5C1, EC), methylene chloride (CH2C12, DCM), trichloroethylene (C2HC13,TCE) and mono- and di-chlorobenzenes. Cl-VOCs concentrations at bed inlet (C,) ranged between 100 and 1,000 p m. Gas hourly space velocities (GHSV) were varied between 2,000 and 10,000 h- (n.c.). The catalytic activity could be better compared on the basis of test runs at the same area velocity (volumetric feed flowhotal geometric surface area of the catalysts), but the standard GHSV was preferred for an easy further comparison with Cr, Pt and Pd-based catalysts (which usually are spheres or pellets). Reactors used in the lab-scale facility shown in Figure 1 were glass made and had several sizes to place the different monoliths. Commercial 15 x 15 x 30 cm monoliths were cut to smaller ones, of around 2.0 x 2.0 x 20 cm. Number of cells used in these "small" monoliths were 3 x 3 , 4 x 4 or 5 x 5, depending on cell density. To avoid bypass in the catalytic reactor, the free space between the monolith and the reactor wall was always carefully filled with carburundum (Sic). For some tests, monoliths were crushed, sieved to several sizes below 1 mm, and then used as particles. Temperature was measured with two thermocouples located just at the d e t and exit of the monolith. Temperature differences between these two points were I 2 "C. Gas sampling and analysis methods in this small facility were already reported [6,7]. Part I1 of this research has been made (with the catalysts and exp. conditions selected from tlus part I) in the small pilot plant shown in Figure 2. It is based on a fluidized bed incinerator of 15 cm i.d. and 5.3 m high. More details about this pilot plant can be found in ref. 20. The metallic and glass-made catalytic reactors in this facility are placed in a by-pass downstream from the incinerator. Catalysts operate in this case under a realistic and complex flue gas composition. Gas sampling in this facility was made periodically before and after the catalytic reactor following the USTO-14 reference method and using tubes with an adsorbent bed (0.1 g of Tenax TA 60/80 mesh). Gas analysis (TD/GC/MS) was carried out in an analytical system formed by: 1st) Thermal desorption unit (Perkin-Elmer, ATD-400) coupled with a gas chromatograph and a Mass Selective Detector. Desorption conditions were: T (desorpt. 1): 250" C, 10 miq T (desorpt. 2, trap): 250" C, 10 min; desorb. flow: 72 d m i n , outlet flow: 40 d m i n . 2nd) Gas chromatograph (Hewlett-Packard, 5890) with a HP-5 column, 30 m, 0.25 mm; gas: He, 2 psi; Temp: 50" C, 2 min, 10" C/min, 250" C, 3 min. 3rd) MSD (Hewlett-Packard, 5971-A): scan 34-300 AMUS
P
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CATALYSTS TESTED Manufacturers and names or codes of the commercial V-W-Ti catalysts (monoliths) used in this work were: Manufacturer Name or code of the catalvst 04-85 & 04-86 BASF AG SIEMENS AG A, B, c FRAUENTHAL (CERAM's group) MN12 KWH (Katalysatorenwerke Huls GmbH) ZERONOX
Reference [81 ~141 ~9,151 [161
Four European research groups also provided these authors with samples of their V-W-Ti catalysts (monoliths too): 0 0 0
0
Politechnico di Milano, IT [ 171 Spanish Council of Research. Institute of Catalysis (ICP), [ 181 University of Twente (Van Ommen et al), NL [ 191. Techn. Univ. of Wroclaw, PL. Table 1. Characterization of some V-W-Ti catalysts used in t h s work.
N2adsortion
Hg porosimetry Catalyst
A (SIEMENS) B ''
c
"
ZERONOX
Total cum. vol
Specific surface area
(cc/ g) 0.33 0.34 0.34 0.27
(m2/ g) 47.3 52.4 40.8 56.0
Pore Bulk Diameter Density Average
(A) 162 154 217 97
(g/cc) 1.85 1.79 1.63 2.11
Pore volume
BET surface Area
Pore Diameter Average
(cc/g) 0.25 0.27 0.27 0.21
(m2/ g)
(A)
58.3 61.0 61.9 46.5
164 163 162 181
b) Main geometrical characteristics Cell Density Wall shape (cells/inch2) thickness (inch) A (SIEMENS) B ''
c
square ' "
"
ZERONOX MN12 (FRAUENTHAL)
" I'
36 36 36 48 225
889
0.022
Pitch Geometric (inch) surface area inch2/inch3 0.16
"
<'
"
"
0.047 0.011
0.04 0.059
21 ' 18 54
Table 2.- Chemical composition (according to catalyst catalogues and/or providers) of the catalysts used. Manufacturer BASF AG
KWH SIEMENS FRAUENTHAL (CERAM) Politecn. di Milano CSIC-ICP
Code Composition (4-85) 29-1 107-01-960 3% V2O5,7% W03, 90% Ti02 (4-85) 36-0404-01-810 0.5% V20s, 9.5% W03, 90% Ti02 0 4-86 6.8% V20s,4% W03,/ Ti02 ZERONOX 6% V20s,4% W03, 0.2 Cr203/ Ti02 A 4% V20s, 7% W03/ Ti02 B 2.4% V20s,6.8% W03/ Ti02 6% V20s, 8% W03/ Ti02 C MN12 3% V2Os PM h-456 h-457 V20JTi02
0.48% V20s, 9% W03/ Ti02 2% VzOs, 1% W03/ Ti02 4% V2O5, 1% W03/ Ti02 4.5 % V20s/Ti02
Univ. of Twente (NL) Techn. Univ. of 8% V205/92% Ti02-Si02 V205/SM 1 Wroclaw (PL) Main physical properties of the catalysts, measured in this work, are shown in Table 1. Their chemicdl composition, according to their manufactures, is indicated in Table 2. More details about these catalysts can be found in the manufacturers' catalogues [8, 9, 13, 14, 15, 161. RESULTS AND DISCUSSION INTRODUCTORY TESTS Error ofthe tests. Although the experimental error in the small, laboratory, glassmade facility was already known by previous research [6, 71, it was checked again with these V-W-Ti catalysts. So, three tests were made under the same exp. conditions for several catalysts. It was checked how experimental error was low (always f 10" C to obtain a given conversion in different test runs with the same catalyst). Errors theory [refs 21 and 221 was kept in mind in the analysis of results andlor figures, thus. Effect of the face gas velocity (mass transfer limitations). It is well known how these monolithic catalysts work in laminar flow because of the small size of their channels. There are transversal gradients of concentration in them, thus. This effect has been quantify in this research by two ways: 1st) Some tests were made with the same gas hourly space-velocity [lO,OOO h" (450" C)] varying the face gas velocity (equivalent to the superficial gas velocity in fured beds) from 28 to 83 cm/s. The monoliths were then cut to different lengths (fiom 10 to 30 cm). Results on oxidation of EC (1,000 ppm at inlet) are shown in Figure 3 for two BASF catalysts, and indicate some external diffusion control at low face gas velocities. To keep it to a minimum extent, most of fbrther tests were made at the maximum (in this facility) face velocity: 83 c d s . 2nd) Mass transfer limitations in these monoliths have also been studied by the well established procedures in chemical engineering, following the detailed
890
method and formulas given by Rodenhausen [ 111 for the same deNOx monoliths. As just one example of these calculations, in the MN12 catalyst working at 300" C with a feeding of 500 ppm DCM in air and a face velocity of 83 c d s , the calculated [ 111 mass transfer coeficient has a value of of 0.10 d s and the gradient of concentrations (between the averaged one in the channel and the interface or wall of the monolith) is of 10 ppm DCM at monolith inlet and of 3 ppm at monolith exit. Working at 83 c d s , mass transfer is not the rate controlling step, thus. 02-content in flue gas. The first part of this research has been carried out with C1-VOCs in air (21 vol % O2 content) but the catalysts here tested would be used in flue gas downstream from waste incinerators, whch usually have 02-contents (regulated by laws) of 7-1 1 ~01%.To know the influence of the 02-content in the flue gas, some oxidation tests were canied out with the same catalyst(s) but with the C1VOC diluted in air and in a simulated flue gas containing 10.7 vol % O2 only. Results in Figure 4 for total oxidation of EC with two BASF catalysts show the lmportance of the 02-content on the conversion of the C1-VOC. T h s fact has to be taken into account when extrapolating or applying results obtained in this first part of the overall study to the incineration pilot plant. Effect o f t h e internal dzfision. To know to which extent internal diffision into the wall of the monolith is important in this study, some monoliths were crushedground, sieved, and then used as particulates. Since the amount of catalyst used in the reactor is relatively high, it was supposed that the composition of all sue fractions is the same. Due to the softness of these materials, only three fractions of particulates were prepared (0.20-0.32, 0.2-0.8 and 0.8-1 mm). With these 3 fractions the same oxidation test was made. Results with DCM on ground ZERONOX are shown in Figure 5 . Even taken into account the error in these measurements, at high temperature (above 300" C) the smallest particle diameter (0.20-0.32 mm) gives the highest conversions, indicating some internal diffusion control at hgh temperature. ACTIVITY TESTS Oxidation of DCM:
X-T or activity curves for oxidation of DCM over three different (so-called A, B, C) catalysts from SIEMENS are shown in Figure 6. X-T curves obtained for different SV and Co values and for all catalysts tested in DCM oxidation are shown in Figures 7, 8, 9, 10 and 11. Experimental conditions are indicated in the figure captions. From these results it is deduced that not all W-V-Ti catalysts have the same activity. The MN12 from FRAUENTHAL and the ones from Politechnico di Milano, and from T e c h . Univ. of Wroclaw were the most active ones. Oxidation of TCE:
Activities of the three (A, B, C) SIEMENS catalysts for oxidation of TCE are shown in Figure 12. Activities of all catalysts tested for oxidation of TCE are shown in Figure 13. Experimental conditions used are indicated in the respective figure captions.
89 1
Results indicate that the 3 catalysts from SIEMENS have a different activity and that the most active ones (under the exp. conditions here used, of course) are MN12 from FRAUENTHAL and “C” and “A” from SIEMENS. When X-T curves for oxidation of TCE and of DCM on the A and C SIEMENS catalysts (shown in Figures 7 and 13) are compared, conversion of TCE (at a given T) was always higher than that of DCM. It means that using SIEMENS catalysts, DCM is more refractory or difficult to be oxidized than TCE. Effect of a second hydrocarbon in theflue gas (preliminary results). It is well known that in the stack or exit gas of waste incineration plants there is not only one VOC or C1-VOC but a mixture of them. Some preliminary tests were made introducing a second hydrocarbon mixed with TCE. Heptane and toluene were selected as key aliphatic and aromatic hydrocarbons. The activity of the catalyst was exp. determined with and without these hydrocarbons as well as with and without steam in the flue gas. Results for mixtures of TCE- (500 ppm) Heptane are shown in Figure 14 and for mixtures TCE-(300 ppm) Toluene in Figure 15. It is seen that: i) Some (up to 1 ~01%) steam in the flue gas enhances the catalytic activity (for C1-VOC oxidation), ii) heptane enhances the catalytic activity while the aromatic hydrocarbon seems to decrease it. Of course, if these V-W-Ti catalysts be selected for a further research, a lot of other tests should be made with mixtures of VOCs and Cl-VOCs. Eflect of the Vanadia-contentin the catalyst: Temperatures to obtain 50% and 90% conversion of the Cl-VOC, T5o and TW,were calculated from X-T curves. Such T50 and T9,, values are plotted against the V2O5content (in the range from 2.5 to 6.0 wt %) in Figure 16. For DCM oxidation the V205-contentdoes not seem to have an effect on T, values. For TCE oxidation T, decreases somewhat on increasing the vanadia content. The fact that the MN12 catalyst has the lowest T, values, Figure 16, can be attributed to its higher cell density: 225 cells/inch2instead of 36 for the SIEMENS catalysts or 48 for ZERONOX (Table 1). By the same reason the catalyst from CSIC-ICP with a very low cell density had the highest T, values (lowest activity) among all catalysts tested. Catalysts comparison using a simpleJirst order kinetic model It is not easy to compare the activity of the V-W-Ti catalysts here tested with the lot of chromia, Pt and Pd based catalysts previously used because they have different shapes (monoliths and spheres) and because very different particle sizes are involved (having thus very different effectiveness factors). For comparison purposes, all X-T curves were adjusted to a simple first order kinetic model (with rate based on overall volume of catalyst, both for monoliths and for fixed beds). From the kinetic constants so obtained (see details of the method in ref 7), the preexponential factors (k,,) of the Arrhenius law and the apparent energies of activation (EWp)were calculated for all catalysts. One example is shown in Figure 17.By the well known compensation effect between k,, and Eapp,the ko values so obtained were recalculated for a given E, value of 44 kJ/mol. Such new k,, value was used [7] as an activity index of the catalyst. The k,, values obtained for the V-W-Ti catalysts here tested are in a broad interval. Even for a given catalyst, the k,, value is quite different depending on it is referred to
892
the monolith or to particles (ground monoliths). For V-W-Ti monoliths ko ranged from (10 to 40) lo3 s-'. For particles the highest ko value (for Eapp=44 kJ/mol) for the most active catalyst [the V205/Ti02catalyst from Univ. of Twente] was (230-260) lo3 S-'.
CONCLUSION Activity for CI-VOC oxidation at lab scale of eleven catalysts in monolithic shape from eight different manufacturers has been here reported. When the ko (for Eapp=44 kJ/mol) values are compared with the similar ones for the Pt, Pd, and chromia based catalysts, indicated in Tables 4 and 5 of ref 7, it is concluded that ground (particles) V-W-Ti catalysts are, by average, more active than the competitive (used for the same application) Pt, Pd and Cr203 catalysts. Nevertheless, as monoliths they are not so active, and the activity of commercial V-W-Ti monoliths is similar to the commercial Pt or Pd spheres. Besides, these results about catalytic activity have to be used joint with the deactivation or life time results (for the same catalysts) shown in forthcoming part I1 of this paper. Their deactivation above 200-250" C for relative high contents of C1-VOCs (besides their price) will be the key or limiting aspect in their use.
NOMENCLATURE Concentration of CI-VOC at the reactor inlet, ppm Apparent activation energy, kJ.mol-' preexponential factor of the Amhenius law, s-' ko: Length of the monolith, cm L: SV (or GHSV): Gas hourly space velocity [gas flow rate at normal conditions or at 450 "C, as indicated in the text/overall volume of the catalyst (bed or monolith)], h-' Temperature at the exit of monolith or the catalytic bed, "C T: Temperature needed to get a given (X) conversion, "C Tx: Face gas velocity at 450" C, cm/s uo: Conversion of the halocarbon, dimensionless X: C0: Eapp:
ACKNOWLEDGEMENT. This work has been carried out under the EC, DGXII, contracts no EV5V-CT94-0530 and ICOP-DEMO-2094-1996. The authors thank to the European Commission, Environmental and INCO Programmes, its financial support. Authors thank very much to Dr. Irene Bergsteiger (Porcellanfabrik Frauenthal GmbH), Dr. Hess from BASF AG, Dr. S. Fischer from SIEMENS-KWU, Dr. Ollenik from KWH, Dr. Avila from CSIC-ICP, Prof. Pi0 Forzatti from Politechnico di Milano, Dr. Van Ommen from Twente Univ. (NL) and to Dr. Janusz Trawczynsh from Techn. Univ. of Wroclaw (PL) the supply of samples of catalysts for their testing at UCM, and to Ms. Ana Padilla for canylng out some tests runs.
893
REFERENCES 1.
2.
3.
4. 5.
6. 7. 8.
9. 10. 11. 12. 13. 14. 15. 16. 17.
18.
19.
Narvhez I., Corella J., Aznar M. P. (1993) "New Processes for Flue Gas Cleaning in Solid Waste Incinerators". Ingenieria Quimica (Madrid), 221-229. Furrer J., Dropsch H., Stoh J. (1998) "Catalyst Development for the Destruction of Volatile Organic Compounds in the Flue Gas of Municipal Waste Incinerators". In Proceed. of IT3 Conference held in Salt Lake City, UT, USA, 337-340. (Ed. by Univ. of California at Irvine). Stoll M., Furrer J., Seifert H. (2000) "Catalytic Destruction of Chlorinated Aromatic Compounds with a V-Oxide Catalyst" Proceed. of IT5 Conference on Incineration held in Portland, OR, USA. Shiraishi Y., Kawabata H., Chichibu S., Furuta S . (1995) "Total Flue Gas Treatment System of Municipal Solid Waste Incineration Plant". Kobelco Tech. Review, no. 18,46-49. Greene H. L., Cheung M., Danals R. S., Vimawala S. V. (1991) "Catalyst for Destruction of Hazardous Chlorinated Wastes and Process for Preparing the Catalysts". U.S. Patent no 5,075,273. Dec. Padilla A. M., Corella J., Toledo J. M. (1999) "Total Oxidation of Some Chlorinated Hydrocarbons with Commercial Chromia Based Catalysts". Applied Catal. B: Environm., 22, 107-121. Corella J., Toledo J. M., Padilla A. M. (2000) "On the Selection of the Catalyst among the Commercial Platinum-based Ones for Total Oxidation of some Chlorinated Hydrocarbons. Applied Catal. 3: Envirom., 27,243-256. BASF AG. "The BASF Catalyst for Dioxin Descomposition". Catalogue, Ludwishafen, Germany. Begsteiger I. (1999) "Improved emission control due to a new generation of high-void-fraction SCR-DeNOx catalysts". Catalysis Today, 27,343. Koser H. (ed.) (1992) Proceed. of the SCR-DeNOx-meeting in Essen, Germany; Vulkan-Verlag, Essen, Ge. Rodenhausen R. (1999) "Case Study: Choosing Selective Catalytic Reduction as a Preferred Technology for the Destruction of NOx". Envirom. Progress, 18 (4), 260-266. Koebel M., Elsener M. (1998) "Selective Catalyhc Reduction of NO over Commercial DeNOx Catalyst: Comparison of the Measured and Calculated Performance". Ind. Chem. Eng. Res., 37,327-335. Odermatt P., Ipek B. "BASF Catalysts for the Removal of NOx and Dioxin in Various applications". BASF AG brochure, Ludwigshafen, Ge. SIEMENS AG Power Generation (KW),"DIOx Catalysts decompose dioxins in the flue gas of waste incineration plants". Catalogue, Erlangen, Germany. Frauenthal Keramic AG, "DeNOx Catalysts". Catalogue, Frauenthal, Austria. KWH GmbH, "ZeronoxRCatalysts". Catalogue, Marl, Germany. Lietti L., Forzatti P., Bregani F. (1996) "Steady-State and Trasient Reactivity Study of Ti02-Supported V205-W03 DeNOx Catalysts: Relevance of the Vanadium-Tungsten Interaction on the Catalytic Activity". Ind. Chem. Eng. Res., 35,3884-3892. Bahamonde A., Beretta A., Avila P., Tronconi E. (1996) "An experimental and theoretical investigation of the catalytic behaviour of monolithic Ti-V-WSepiolite catalyst in the reduction of NOx with NHT. Ind. Eng. Chem. Res., 35 (8), 2516-2521. v& Hengstum A.J., van Ommen J.G. (1983). Applied Catal., 8,369.
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20. Corella, J., Toledo, J.M. (2000) “Incineration of doped sludges in fluidized bed”. J. Hazard. Materials B80, 81-105. 21. Neuilly M. et CETAMA (1998) “Modelisation et estimation des erreurs de mesure”, (2nd ed.), Lavoisier (Paris). 22. Senent F. “Errors’ theory” (booklet for students), Ed. by Universities of Valladolid (1961-65) and of Valencia (from 1966 to date), Spain.
895
Biomass burner designed to reduce nanoparticle emissions C.K. Gaegauf'; U. Wieser' ; S . Unterberger*; K.R.G. Hein' 'Centre of Appropriate Technology,CH-4438 Langenbruck, Switzerland; 'Institute of Process Engineering and Power Plant Technology(IVD),University of Stuttgart, D-70569 Stuttgart, Germany
ABSTRACT: A process development unit (PDU) to bum biomass fuels was designed in order to improve the thermal conversion process and to minimise the formation of nanoparticles. Comparison is made with two commercially available wood chip burners. The PDU consist of a primary combustion chamber, a high temperature cyclone and a secondary combustion chamber, were a vortex is created to establish strong mixing of the gas components. The design allows process staging with respect to temperature and combustion air. The particle emission characteristics were monitored with a scanning mobility particle sizer (SMPS) for nanoparticle size distribution and number concentration. The results demonstrate that the PDU covers most of the required features for new biomass burners: complete combustion and low emissions. Under identical combustion parameters the new burner had significantly lower particle emissions than the investigated commercial burners. The total suspended particle (TSP) concentration in the PDU flue gases varied between 36 and 124 mg/m; depending on excess air settings. The flow pattern of the vortex in the secondary combustion chamber was measured with a Laser Doppler Velocimeter (LDV). The profile measurements were performed at various feed rates and air settings as well as different hardware configurations. The LDV measured the axial and tangential velocity vectors of the flow in the hot flame simultaneously. OBJECTIVES There is a growing interest for biomass fuels to cover a larger part of the required energy supply. The increasing use.of renewable and C02-neutral energy sources needs to go hand in hand with improved thermochemical conversion processes with minimum harm to human beings and the environment. Solid fuel burning combustion systems produce smoke emission, which are of concern to authorities and the public. The organic components of these emissions are of most concern because they include particulates covered with compounds known to be carcinogens and respiratory irritants. Because of the small size of the particulates (typically of the order of 1 micron and less) they easily pass through the nose and throat into the lungs [ 13.
896
The World Health Organisation (WHO) specifies limits on air quality in the form of limits on the Total Suspended Particulates (TPS) allowed in the atmosphere. In turn, authorities use these specifications to place limits on airborne particulate matter allowed from emissions sources in areas, which suffers from air pollution. Any improvement on thermochemical conversion of biomass ought to consider the reduction of particulate matter (PM). One has to take into account that particulates harmful to the environment in the size less than 1 micron behave like gaseous effluents. They can not be withdrawn by means of cyclones, etc. The goal of any improvement is to decrease the amount of particulates already in the thermochemical conversion.
SCIENTIFIC AND TECHNICAL DESCRIPTION ANALYSIS OF NANOPARTICULATE EMISSIONS A scanning mobility particle sizer (SMPS) was applied to analyze nanoparticles [2][4]. Size distributions and total number concentrations (TNC) of particles in the range from 0.01-0.7 pm are determined by analysis of particle mobility. An impactor with a cut-off size of 1 pm is used to withdraw the coarse particle fraction. Exhaust gas is taken with a probe, which is also fed with particle free air. The resulting dilution factor is adjusted by the flow rate of the diluting air and the total flow. To prevent condensation of water onto the particle surface, the dilution factor is chosen high enough, to achieve a dew point below ambient temperature. The analytical set-up is shown in Figure 1.
Gas 4-
Particle size distributaon
8
LXend 1 Flueduct 2 Rotatmg disc dduhon unt 3 Exhaustgaspump 4 Parhclefilter 5 Flow & ddutaon controler 6 Impactor, cut-off at dp 0 5 pm
Fig. I :
7 Radioachve neutdzer 8 Differentd mobility analyzer DMA 9 Controlaunt 10 Condensed particle counter CPC 11 Vacuum pump for sample flow 12 Size distribuhon analysis 13 PAH sensor
Instrumental set-up for exhaust gas analysis and particulate measurement
897
Polydisperse aerosol particles in the sample gas passes through a radioactive bipolar charger, establishing a bipolar equilibrium on the particles. The particles then enter the differential mobility analyzer (DMA, TSI 3071) and are separated according to their electric mobility. A subsequent condensation particle counter (CPC, TSI 3025) evaluates the number concentration of the monodisperse aerosol particles. The number size distribution is measured by varying the DMA voltage over the measuring range and by recording the accompanying particle concentrations with the CPC. This scanning method is controlled by computer software. The characteristic data of particles are given by the mode diameter (MD) as the most frequent size of a particle population, the geometric mean diameter (GMD) of the particles and the total number concentration (TNC) as the total amount of particles over the whole measured range. The TNC is based on the flue gas volume at a standard oxygen content of 13%. The particle distribution graphs are done as commonly used in aerosol measurement: the channel width, which represents particle diameter (dp) range, is plotted on a logarithmic scale against the total number concentration (TNC), that is calculated from the measured number of particles (dN) divided by the logarithm of the channel width (dlog(dp)), where dp is the mobility diameter. STATE OF THE ART BOILER TECHNOLOGY
The emissions of two different commercially available types of wood chip combustion systems were analysed: one boiler with a power output of 70 kW with a stoker type burner, the other boiler with a power output of 600 kW and a moving grate burner. Both systems had a computer controlled feed rate and combustion air supply. The particle size distribution of the 70 kW boiler is shown in Fig. 2. All size distribution data were averaged over a time period of 45 minutes at a feed rate of 20 kg/h. The mode diameters of the particles vary from 62.6 nm at oxygen concentration of [02]= 1 1.6 % in the flue gas up to 87 nm at [02]=2.7 %, depending on excess air supply. NC = dN/dhg(dp) [em-]
100
10
loo0
Partide mobility diameter dp [mu]
11.6 % 0 2 7.9 % 0 2 - 6.2 % 0 2 -2.7 C-__--__ . ..-
Fig. 2:
K02
1
Particle number concentration of a 70 kW wood chips burner as function of excess air given as surplus oxygen (02) concentrations in the flue gas.
898
The total number concentrations (TNC)vary between 4.64E+9 per cm3 for an oxygen concentration of [04=2.7 % and 1.16E+10 per cm3 for [02]=1 1.6 %. With higher excess air ratios there is a slight decrease in size but an increasing total number concentration. The increase of particle concentration was more distinct towards enhanced surplus air. The combustion chamber of the moving grate boiler system was run over a wide range of bum rates, starting from 150 kW up to 600 kW. The analysis of particle emissions such as particle diameter or total number concentrations (TNC) showed only little variations with the process parameters (Fig. 3). Similar to the 70 kW burner system, all test runs resulted in mode diameters (most frequent size diameter) of 100 nm. The geometric mean diameter (GMD) has been analysed in the range of 90.45 nm up to 95.0 nm. The total number concentration (TNC)was between 2.81E+7 and 5.12E+7, depending on excess air supply. NC = dN/dlog(dp) [em-31 1.OEM8
8.0Ei-07
6.OEi-07
4.OEi-07
2.OEM7
O.OEi-00 100 Mobility dismeter [nm]
10
I Fig. 3:
A
260 kW, L 1.3
450 kW, L 1.3 -310
kW, L 1.9 -560
1000
kW, L 1.9
I
Particle analysis of a 600 kW wood chip burner at different burn rates of the combustion chamber and two excess air settings (L) at 1.3 and I .9.
DESIGN AND CONSTRUCTION OF THE PROCESS DEVELOPMENT UNIT (PDU) The goal of the project was the development of a low emission biomass burner for boilers in the range of 50 - 500 kW thermal output [ 5 ] . The main feature of the PDU burner design is the secondary combustion chamber creating a vortex to increase gas phase turbulence in order to maximise complete combustion. The main emphasis of the burner design is to reduce the release of particulate matter. The combination of vortex and air staging techniques is supposed to reduce also the NO,-emission level. The combustion system is designed to bum preferably wood chips. The process development unit (PDU) was designed and built in co-operation with the boiler manufacturer Schmid AG, Eschlikon, Switzerland (Figure 4).
899
8
4 Heat Exchanger
Cross Section Vortex-Burner
I I
7
Fig. 4:
Layout of basic design of the process development unit (PDU) 1 primary combustion chamber, 2 hot cyclone, 3 ash disposal, 4 primary vortex, 5 secondary vortex in the secondary combustion chamber, 6 primary combustion air, 7 secondary combustion air, 8 flue gas
The PDU was designed to meet following requirements: - Complete combustion with low emissions of products of incomplete combustion (PIC) such as volatile organic compounds (VOC) and carbon monoxide (CO) - Low particulate emissions and low nitrogen emissions (NO,) To meet these requirements, the PDU combines a primary thermal conversion chamber with hot cyclone and a secondary combustion chamber. Theprimary thermal conversion chamber comprises 1) a stoker 2) a pre-drying zone to remove fuel water, 3) a primary combustion zone to produce the required energy for the pyrolysis/gasificationprocess, and 4) a pyrolysis and gasification zone to convert solid fuel into gaseous ones. The high temperature cyclone allows a fuel rich atmosphere (mainly carbon monoxide, CO) but no oxygen for denitrification of nitrogen compounds (NO,) The secondary combustion chamber creates a vortex by secondary combustion air to establish strong mixing of gas components for facilitating chemical reaction, temperature level to achieve complete combustion with low carbon particulate content and consequently low formation of PCDF and PCDD.
900
Fig. 5:
PDU site in the Laboratories of the Centre of Appropriate Technology: Left: Laser Doppler Velocimeter (LDV) for profile velocity measuring system in the secondary combustion chamber. Right: Seeding system to entrain tracer particles in combustion air system.
The design of the burner allows a staging of the process by means of temperatures and combustion air. To achieve the required goals the following boundary conditions needed to be met in the thermal conversion processes: Primary thermal conversion chamber: reduction atmosphere at 1160°C at air ratios of 0.7. Hot cyclone: denitrification of NO, at temperatures > 1100°C and potential of removal of ash fractions with low content of heavy metals [3]. Secondary combustion chamber: combustion temperature > 800°C for complete combustion with low particulate emissions.
RESULTS PROCESS PARAMETERS OF THE PROCESS DEVELOPMENT UNIT (PDU) The process development unit (PDU) was fed with 2 species of wood chips: a softwood (mainly spruce, Picea abies Karst.) and a hardwood (beech, Fagus sylvatica L.). The hardwood was supplied by IVD as reference fuel, that has been tested in other combustion equipment for reference. The PDU was operated at an hourly feed rate of 18 to 40 kg fuel. The other process parameters are listed in Table 1. The retention time of the hot gases in the cyclone were varied between 0.7 seconds (40 kg/h feed) up to 1.5 seconds (1 8 k g h feed) and !?om 0.4 up to 0.7 seconds in the secondary combustion chamber.
90 1
70 % to 80 % of the energy in the fuel was converted in the primary combustion chamber and the cyclone. The specific heat load was in the range of 235 up to 507 kW/md in the cyclone and between 50 and 102 kW/md in the secondary combustion chamber.
Table I .-
Process parameters of process development unit (PDU) at different fuel loads and fuel species
Fuel input
I
Process parameters
20
78
I .5
79 %
255
0.7
21 Yo
50
30
118
I .O
79 %
377
0.5
21 Yo
78
40
157
0.7
79 %
507
0.4
21 %
101
18
76
I .5
72 %
235
0.7
28 %
57
27
114
1.o
71 %
332
0.5
29 %
102
32
135
0.8
80 %
454
0.4
20 %
75
FLOW PATTERNS IN SECONDARY COMBUSTION CHAMBER OF PDU The flow pattern in the secondary combustion chamber was measured with a Laser Doppler Velocimeter (LDV). The burner was equipped with a heat resistant glass cover to give the laser diagnostic access to the combustion process. To make the flow patterns visible, magnesia powder particles were added to the primary and secondary combustion air. The profile measurements were performed at various feed rates and air settings as well as different hardware configurations. The LDV measured the axial and tangential velocity vectors of the flow in the hot flame simultaneously but not the radial ones. Starting from a zero level at the position of the combustion air nozzles, a total of 14 levels were measured at equivalent distances of 30 mm above. At each of these levels 10 points of the combustion chamber diameter were measured every 30 mm. The velocity vectors indicate a back flow of flue gases up to 240 mm above zero level. A low axial velocity was observed at the level of 90 mm in the centre zone (250 mm) and increased up draught velocity vectors in the position 0 - 150 mm (Figure 6). The back flow at position 130 - 220 mm and the up draught are caused by the torous shape of the vortex. This could be confirmed by observation of the combustion process through the glass cover.
902
240 m m above ground level 4
-I -2
I
J 180 m m above ground level
4
1: 3
0
Fig. 6:
Axial flow patterns of the secondary combustion chamber at various levels above the secondary air nozzles (ground level). Vortex patterns are visible at 90 mm level (back flow towards centre, position 250 mm, and up flow at 0 to 150 mm position). At 180 mm level there is a back flow of flue gas at 120 - 220 mm position.
The axial velocity vectors increase at the level of 180 mm in the centre as well as at the periphery. There is still a back flow in the radial position of 120 to 220 mm. This indicates a back flow of flue gas with recirculation effects in the combustion process. The recirculation of flue gas is visible on levels up to 240 mm. The entire vortex is rotating around the middle axis with a velocity between 1 and 5 d s (Figure 7).
903
240 mm above g r o u n d level
5
1-
4
x
3
t: 1 - I0
0
-2
180 mm above & m d level
90 mm lbove #roundlevel 5 ,
1
4
z
3
t:
lo -I
0
-2
Radial position from periphery [m]
Fig. 7:
Radial flow patterns of the secondary combustion chamber at various levels above the secondary air nozzles. A cyclone flow is visible on all levels. The velocity decreases towards the periphery.
EMISSION CHARACTERISTICS OF THE PROCESS DEVELOPMENT UNIT (PDU Particulate emissions For measurement of particles and gaseous emissions the same equipment and methods were applied for the PDU as for the previously tested commercial wood chip boilers. The patterns of particle size distribution of emissions from the PDU showed similar characteristics as those from commercial boilers. The mode diameters (MD) are at sizes between 80 and 100 nm (Figure 8).
904
NC = dtildloeldo) Icm-31
Fig. 8:
Typical particle size distribution of the process development unit (PDU) with softwood
Particles from softwood are slightly coarser than those of hardwood (reference fuel). Hardwood particle number concentration is somewhat higher than softwood (Figure 9). NC = dN/dbg(dp) [em
.. ..
,
10
IW
.,. .. ,. ,. .. .. ..., .,. ... ,.. ... ... .
,
,
I
#
IWD
McbiIiidiumn[ nml
Fig. 9:
Typical particle size distribution of the process development unit (PDU) with hardwood (Beech, reference fuel)
The particle number concentration of the PDU is of an order of magnitude smaller than those of the commercial burners (Table 2). Volume and surface of flue gas particles were calculated from the number concentration, assuming spherical particle shapes. Particle surface attracts polycyclic aromatic hydrocarbons (PAH). Any increase of surface is a potential of health risks and a decrease in particle surface thus reduces the transport of carcinogens compounds into the lungs. The particle surface from the PDU is thus reduced to a factor of 20 compared to the commercial boilers. The total suspended particle (TSP) in the fuel gas of the PDU depended on the excess air settings. A higher air throughput in the primary combustion chamber led to more entrained ash particles in the combustion gases. This is also true for the 70 kW commercial boiler, which was very sensible to a high airflow in the stoker.
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Table 2:
Comparison of total number concentration, surface and total suspended particles of commercial boiler and process development unit (PDU)
Plant
Process development unit (PDU) ~~
Fuel
~~~
Softwood
Commercial boiler
~~~
Hardwood
Waste wood
Softwood
Waste wood
reference fuel Total particle number concentration (TNC)[C~-’]
2.0 EM8
2.61 EM8
-
35.1 E+08
-
Total particle surface (TPS) [cm2cm.’]
6.7 E-02
5.34 E-02
-
131 E-02
-
106 236
78-600
1 1 1 -251
-
17 - 174
Total suspended particulates TSP [mglm’@13% 0 2 1
36 124
-
-
[mgl MJI
2 5 - 86
-
74 - 164
54 417
Gaseous emissions
The combustion in the vortex-burner was homogenous. The products of incomplete combustion (PIC), here as carbon monoxide, were very low, typically between 5 and 500 mg/m;. The nitrogen oxides in the flue gas varied considerable depending on the fitel used. The nitrogen oxides (NO,) in the flue gas for one softwood chip was around 90 to 1 10 mg/md , for the other fie1 in the range of 150 to 200 mg/md ,based on 13% oxygen. CONCLUSION The basic layout of the process development unit covers most of the required features of innovative wood combustion systems such as: - Combustion with low emissions of products of incomplete combustion (PIC) such as volatile organic compounds (VOC) and carbon monoxide (CO). - Low particulate emissions and low nitrogen emissions (NOx). - Ash fractions with low content of heavy metals. It could be shown that under identical conditions the new burner had significantly lower particte emissions than the investigated commercially available burners. The number of nanoparticle emissions could be reduced from a typical total particle number concentration of 35.0E+08 per cm3 in the commercial wood chip burner down to 2.00E+08 per cm3 in the process development unit. The knowledge and the results achieved are a good basis for new wood chip boiler design. The combustion stability over the entire heat load range needs some further improvement and the design of the vortex chamber with the swirl formed by fuel gases and combustion air needs some more engineering refining.
906
ACKNOWLEDGEMENT The research is part of a joint European project of the European Commission in the framework of the Non Nuclear Energy Programme Joule 111. The Swiss Federal Office for Education and Science (BBW) funded the Swiss part of the project. REFERENCES Non-Biological Particles and Health (1995). Committee on the Medical Effects of Air Pollutants. HMSO Publication, 1995, London H u g h C.( 1996) New Applications of Aerosol Photoemission: Characterization of Wood CombustionParticles and TimeResolved Thermal Desorption Studies, Diss. ETH Nr. 11975, Swiss Federal Institute of Technology, Zurich Brunner, T., Obemberger, I. (1996). New Technologiesfor NOx-Reductionand Ash Utilizationfor Biomass Combustion Plants, Abstracts of the 9th European Bioenergy Conference, Copenhagen, 1996 H u g h C., Gaegauf C., Kuenzel S., Burtscher H. (1997). Characterization of WoodCombustionParticles. Morphology, Mobility and Photoelecfiic Activiy. Environmental Science and Technology, 3 1, p. 3439-3447. Gaegauf C.K., Wieser U. (1998) Biomass Burner with Low Emissions of Particulates, Proceedings of the International Conference Wiirzburg, Germany, C.A.R.M.E.N. D-97222 Rimpar, Germany
907
Emission of UHC and CO From a Biomass Furnace N. Griselin and X.S. Bai Division of Fluid Mechanics, Lund University, S-22100 Lund, Sweden
ABSTRACT: Biomass combustion in a 40MW furnace has been investigated experimentally and numerically, aimed at understanding the processes of combustion and pollutant emissions under different conditions. Experiments were carried out previously using fine thermocouple for temperature and sample suction probe with flame ionization and electrochemical cells for analyzing gas compositions.3-D numerical simulation of combustion and pollutant emissions are currently performed using Favre Averaged Navier-Stokes equations, species transport equations, energy conservation equation and k-E turbulent closure, as well as Lagrangian-Eulerian two-way coupling for particle tracking and combustion modeling. Numerical simulationsand measurement of unburned hydrocarbons (UHC), CO, 0 2 and temperature on the top of the fixed bed are used to model the amount of tar and char formed in pyrolysis and combustion of biomass fuel (wet wood chips) in the bed. Numerical calculations are compared to the measured data. It is shown that different over-fire secondary air supply leads to different CO emissions at the outlet. The emissions of CO can be reduced through controlling the secondary air supply. Char formed in the bed is low in terms of its influence on the heat release, however it has significant influence on the CO distribution in the upper part of the furnace and at the outlet INTRODUCTION Biomass combustion represents an important source of energy production because of the environmental and political benefits of burning biomass fuel. A feature of biomass fuel is its tremendous diversity of structure, reactivity and properties, and its regional and seasonal variation [l]. There is a need to investigate the fundamental biomass combustion processes. Previously, there have been numerous researches on coal and biomass pyrolysis and combustion, as reviewed in [l-51. Biomass fuels are generally low in carbon and high in volatile matter, oxygen, and moisture, explaining their low calorific values. Together with their low bulk densities, it implies that larger volumes are required to supply the same heat input as for coal. It is found that biomass fuels release a proportion of their volatile matter early and at a rapid rate followed by a slower release of the remaining content. As a result there is more tar (liquid) formed and less char formed in biomass pyrolysis. This makes up difficulties for burning biomass in industrial boilers at optimal conditions and controlling the unwanted species emissions. Overall furnace combustion efficiency depends strongly on the different subprocesses that interact one another. These are the chemical reactions involved in the 908
oxidation of the fuel and in the pollutants formatioddestructions processes; fluid flow (turbulence); as well as heat and mass transfer processes. Driven by the need for more efficient means of power generation and pollutant reduction, computational simulation of furnace operation has assumed an increasingly important role. Computational tools make it possible to predict trends in furnace performance characteristics, with reasonable accuracy. Testing of such furnaces is extremely expensive and prohibitive thus it is desirable to develop models and simulation tools to analyze and predict the formation and destruction processes of pollutants, such as fly ash and unburned particles, unburned hydrocarbons, CO, and NOx. Theoretical modeling of turbulent combustion and char burnout can be carried out by numerically solving the Favre averaged Navier-Stokes equations with k-E turbulence models [l], or using large eddy simulation (LES) technique when further details in smaller time and length scales are wanted. Char particles emitted from the bed can be tracked using Lagrangian or Eulerian approach [6].Despite the progresses in different research areas, more effort is still needed in utilizing different sub-models for prediction of overall biomass combustion and pollutant emissions in full-scale furnace. This work is aimed at exploration of pollutant emissions (e.g. CO, UHC) in a 40 MW fixed-bed furnace burning wet wood chips, using both experimental measurements and theoretical simulations. The following issues are emphasized: (1) the influence of turbulent flow motion, namely the over-fire secondary air supply on CO formations and (2) the influence of char burnout on CO formations. unburned light hydrocarbons (UHC) and Experimental measurements of CO, 02, temperature are carried out under different operating conditions and over-fire secondary air staging. It is shown that the emissions of CO can be reduced through controlling the secondary air supply. Char formed in the bed is low in terms of its influence on the heat release, however it has significant influence on the CO distribution in the upper part of the furnace and at the outlet.
FURNACE AND EXPERIMENTS The furnace used in this work is sketched in Fig. 1. It is designed specifically to burn high-moisture wood chips. Fuel is fed onto the top of the bed and moves downward as it is consumed. At the top of the bed (section l), fresh fuel receives heat from the convective flow of recycled combustion products, which dries the top layer of fuel. Further heating of fuel in the next layer down (section 2) drives out the volatile matter consisting primarily of hydrocarbon vapours and tar, carbon monoxide, and hydrogen. There is little oxygen in the gas at this point. The next layer down (section 3) is the char layer. The char first exposed to the air burns out and the ash is mostly retained on the grate (sections 3 and 4),thereby insulating the grate from the highest temperature. In the full-scale fixed-bed systems studied here, typical air distribution profiles employ 23% of the air as under-fire air. The remaining air, over-fire air,is split between three levels of jets. The first level of jets is located close to the bottom of the bed (zone 6), for combustion of the bed-pyrolysis products. The second level of jets (zones 7 and 8), at the extremity of the furnace arch, creates an air curtain through which the mixture of combustion gases and particles must pass. The third level of jets (zone 9), in the upperpart of the furnace, has tangentially over-fire design and hence creates a turbulent, recirculation and fully three-dimensional flow that thoroughly mixes the air with combustion gases coming off the bed, to form final products, COz, HzO.
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Figure 1:Sketch of the fixed-bed biomass furnace showing the different bed modeling sections, under-fire and overfire air jet levels, and measurement ports M-2, M-3 and M-9. It is shown that different over-fire secondary air supply leads to different pollutant emissions at the outlet. Two different operating conditions have been considered to demonstrate this. (a) In Case 1,24% of the over-fire air was injected at zone 6, 19% at zones 7 and 8, and 32% at zone 9; (b) In Case 2, more over-fire air was injected at zone 6, (30%), while less air at zones 7 and 8, (15%).The under-tire air and the third level or air jet (zone 9) was kept constant. Measurements were carried out at a few points at three furnace heights, denoted as M2, M-3 and M-9 (Fig. 1). Species concentrationand temperature were measured inside the furnace through the three ports and at the outlet. At each of the three ports, four points were measured along the zdirection (depth direction, normal to the plane shown in Fig. 1). The temperature was measured using thermocouple type K (Ni-Cr/Ni-Al), with functional range up to 1650 K. Light UHC (gas phase only) were measured by sampling suction probe (SSP)and analyzed with flame ionization. 02 and CO were also measured using SSP and analyzed with electrochemical cells. High CO concentration could not be directly measured using the current apparatus. CO was diluted when the volume fraction of CO was larger than 1%. The uncertainty in the measurement were mainly due to the
910
heat losses from thermocouple to surroundings by conduction and radiation, the catalytic reactions on the surface of the thermocouple and the dilution of high CO concentrations. More details about the experimental work were presented previously [7]. The experiments at M-2 and M-3are mainly for exploration of biomass pyrolysis and combustion in the bed. This information is used to calibrate the bed modeling. Measurement at M-9 is used to calibrate the gas phase combustion and char burned modeling.
MODELING Measurement provides only limited information about the combustion and pollutant emission process. To explore further details about the flow turbulence, the interaction of different processes and influencing factors, numerical calculations are performed. Different aspects need to be considered in the modeling of fixed-bed furnace combustion, including fixed-bed modeling, gas-phase and particle-phase simulation, as well as interactionbetween the different phases. Two bed models are tested: one with char and tar (model 1) and the other one with only tar included (model 2). In these models the biomass pyrolysis and combustion in the bed is constructed using global mass and energy conservationlaw. The products from the combustion and pyrolysis of wet wood chips are presumed to be tar, C&, CO, H2. COZ, H20 and char. In both models, the content of tar is assumed to be Cl,,Hs (67%), C& In model (17%),C7Hsand CsHlo(each 8%) [7]. Tar is therefore denoted by c9.2f&S.1429. 1, char is included in the calculation as flying particles blown off the bed, thus less fuel is available in the bed. It is postulated that near the furnace’s side-walls, the major process in the bed is the devolatilization and char formation. Near the center of the furnace char and tar are mostly oxidized to form light UHC and CO, meanwhile consuming more oxygen. The measurement at M-2 and M-3 are used to validate the bed models. The gas and char formed in the bed is assumed to follow the global reactions listed below.
Reactions I-IV are global steps for gaseous fuel oxidation. Reactions V-VII, are for char oxidation. The single film model is used here, where the particle is consumed via reactions with oxygen (or carbon dioxide) and no reaction occurs in the boundary layer. CO and C02 are the two products formed at the particle surface. The first four reactions are treated based on the eddy-dissipation concept [8], whch assumes that chemical reactions in the gaseous phase occur rapidly and the mean consumption rate of fuel is limited by the mixing rate of fuel and oxidant. The char reactions are treated using kinetic Arrhenius expression. The equations for the gas and particle phases are solved separately, with coupling between phases. The time-averaged gas equations are solved in an Eulerian framework.
91 1
The gas phase fluid motion is modeled using Favre averaged Navier-Stokes equations with standard k-E turbulence closure, together with Favre averaged energy conservation equation, species transport equations. Discrete transfer method is used for calculating the radiation source term in the energy conservation equation [9]. Finite difference method is used to solve the general form of the equation yielding solutions for the gas-phase temperatures, velocities, the turbulence intensities and species concentrations. Sub-models for the particle motion, heat transfer, and combustion are used in the framework of a Lagrangian approach. During its flight the particle interacts with the gas phase flow field producing source terms which are implemented using the particle source in cell method (PSIC), in the momentum, enthalpy and mass equations [lo, 111. The particle-particle interactions are neglected since the particle phase is dilute in the furnace studied here. Fluctuating components of gas velocity are selected from a Gaussian disfribution with variance derived from the local value of turbulent kinetic energy. Integration of the particle equation of motion yields the particle position at any given instant in time. The particle mass change due to char burnout is controlled by the rate of oxidizer diffusion to the particle and the rate of chemical reaction at the surface given by
where A , is the external particle surface area, Yi is the oxidizer species ( 0 2 or COz). is the oxidizing species partial surface density far from the particle, n, is the stoichiometric ratio, Mc and MYi are the molecular weight of carbon and oxidizer, respectively. The overall rate k, is given by the combination of chemical and diffusion rate terms kt = kch’kph/(kch-k kpd,
pri(m)
where b,the chemical rate constant, is determined using the following Arrhenius expression kch = k0,c (TgWMYi) exp(-EJR Tp). Here h,cis the pre-exponential factor, Ec is the activation energy, R is the universal gas constant, T, and Tpare the temperature of gas and particle, respectively. The mass transfer coefficientk@ is given by kph = Nu D J d P Here D m is the binary diffusion coefficient, dp is the particle diameter, and Nu is the Nusselt number. Major data used in this work have been taken from both experimental data [7] and beech wood data [ 121. Particle temperature is obtained from the particle energy equation, which includes the effects of particle mass loss, particle combustion and heat loss to the surroundings. The Lagrangian particle equations are solved for a representative number of particle trajectories. Typically, about 600,000 trajectories have been traced, with nearly lo00 starting locations. Particle size of 50 ,um radius have been chosen and injected at each location. The (PSIC) procedure, briefly described previously, is repeated until convergence is achieved on the particle source terms. The numerical computations are carried out on a 42 x 102 x 38 Cartesian finite difference cells. The system of transport equations discretized using finite difference
912
schemes is solved by pointwise Distributive Gauss-Seidel (DGS) relaxation, with MultiGrid method to accelerate the convergence [ 131.
RESULTS Experimental and computational results are shown in Fig2 to Fig.5, and Table 1. Fig.2 to Fig.4, provide description of the species distribution for case 1, respectively CO, 02 and UHC, both in the x-y plane at the middle of the furnace (vertical cut), and in the x-z plane, in the height of the measured points (horizontal cut). An extra cut in the top of the furnace (called M-10) is also presented to show the effect of the flying char combustion in the upper part of the furnace. Fig.5 presents the temperature profile in the x-y plane, at the middle of the furnace. Table 1, finally, summarizes the calculation results of CO emissions at the outlet, together with experimental results for both case 1 and case 2.
Figure 2: CO measured and calculated for case 1, at M-2, M-9, and in the upper part of the furnace (M-10). (+) exp. data; (- - -) model 1 (char burnout); (-) model 2 (no char burnout).
As seen from the experimental data at M-2 in the lower part of the furnace, CO and UHC profiles are non-uniform. CO volume fraction changes from less than 1% near the furnace wall (z=O) up to 3.5% at the center of the furnace. UHC has similar trend, though the UHC volume fraction is an order of magnitude lower than CO. From the experimental data at M-9, one can see that UHC and CO have considerably oxidized. Measurements in Table 1 shows that about 600 ppm of CO is emitted at the outlet. The 0 2 profile is also non-uniform at M-2, with a trend opposite to the UHC and CO. About 5% to 15% O2 is 913
measured, showing slower oxidation of fuels near the walls. At M-9, 0 2 profile remains non-uniform, indicating that either the chemical reactions between fuel and oxygen or the mixing between air from zones 7 and 8, and products formed in the lower part of the furnace is still in progress. The measurements of 02,CO, UHC and temperature at M-3 also provide information about the pyrolysis and combustion at the bed section 3 and 4 (results have similar trends and are therefore not presented here).
Figure 3: O2measured and calculated for case 1, at M-2, M-9, and in the upper part of the furnace (M-10). (+) exp. data; (- - -) model 1 (char burnout); (-) model 2 (no char burnout).
As shown in Fig.5, the computed temperature is relatively low (< 1500 K), which is in agreement with the current measured data and data that have been reported in other studies of biomass combustion in fixed-bed furnace [141. From the measured data in Table 1, different arrangement of the secondary air supply is found to be influential. By decreasing the intensity of the air curtain (case 2), more CO (about 800 ppm) are found at the outlet. There are few questions concerning the experimental results: (1) How is CO formed in the furnace? (2) How does the secondary air supply affect the CO emissions? As seen in Fig. 2 to Fig. 4, char combustion, as flying particles, in a fixed bed biomass furnace, has a little effect on the gas field in the lower part of the furnace. This can be explained by the fact that the mass of char emitted from the bed is only 5% of the carbon content in the wood,and a large amount of other fuels (tar, CO, UHC) are not yet completely oxidized. In the upper part of the furnace, tar and other gas phase species are mostly consumed, leaving char burnout the most contributing source for the CO distributions (note the 914
Figure 4: UHC measured and calculated for case 1, at M-2, M-9, and in the upper part of the furnace (M-10). (6)exp. data; (- - -) model 1 (char burnout); (-) model 2 (no char burnout).
Case 1
Species
CO (ppm wet)
Model 1
580
Model 2
Case
2
Exp. Data
Model 1
Exp. Data
592
78 1
183
212
Table 1: Emission of CO at the outlet: model 1 (char emitted from the bed is 5% of the total carbon mass of the wet-wood chip); model 2 (no char burnout included).
difference of char burnout on CO at M-10). As char is burning along the path of the particle to the outlet, the amount of CO released by the particle in the last part of the furnace does not have sufficient time to be converted to final products, yielding high level of CO emissions. This can be shown in the calculation in case 1 of CO emission at the 915
outlet between model 1 (char burnout included) and model 2 (no char burnout included). In the former, 580 ppm CO is predicted at the outlet, while in the latter, 272 ppm is found. The prediction of case 2, i.e. the effect of modifying the over-fire excess air is presented in Table 1. Although the air supply in the two cases provide roughly the same amount of air to the furnace, the CO emissions are fairly different, as found from the experiment. Air supply conditions in case 2 yields less effective combustion with higher emissions of CO. Modifying the mass flow of the over-fue air jets can sipficantly change the residence time of species in the furnace. In case 2, the flow is accelerated by the fust level of air jets above the bed, while the intensity of the air curtain from the second level of air jets has been decreased by 20%.
Level m p
1135.43 1061.74 988.057 914.37 840.682 766.905 693.307 6 19.62
I
472945 388.557
Figure 5: Temperature (K) distribution in the x-y plane, at the middle of the furnace.
916
CONCLUSIONS The combustion process of wet wood chips and formation of pollutants in a biomass furnace have been investigated. Distributions of species CO, UHC, O2 where calculated numerically and compared to experimental data. It is shown that char, as flying particle, though in small amount has a significant influence on the CO emissions at the outlet. Numerical simulation indicates that half of the CO emission at the outlet is due to the combustion of flying char particles at the upper part of the furnace. Over-fire air staging has a significant influence on the residence time of particles and gas species in the furnace, and thus the conversion of fuel and intermediate species to final products.
ACKNOWLEDGEMENT: The authors gratefully acknowledge the support of this research by the Swedish National Energy Administration, and would like to thank MalmoViinne AB for the experimental support. REFERENCES 1. Abbas, T., Costen, P.G. and Lockwood, F.C., Twenty-Sixth Symposium (International)on Combustion,The Combustion Institute, Pittsburgh, PA, 1996, pp. 304113058. 2. Serio, M.A., Charpenay, S., Bassilaskis, R., and Solomon, P.R., J. Biomass Bioenergy 7:107-124 (1994). 3. Chen, Y., Charpenay, S., Jensen, A., Wojtowicz, M.A. and Serio, M.A., TwentySeventh Symposium (International) on Combustion, The Combustion Institute, Pittsburgh, PA, 1998,pp. 1327-1334. 4. Hobbs, M.L., Radulovic, P.T., and Smoot, L.D., Prog. Energy Comb. Sci, 19505586 (1993). 5. Di Blasi, C., Prog. Energy Comb. Sci., 19:71-104 (1993). 6. Peirano, E., Lecher, B., Prog. Energy Comb. Sci.. 24:259-296 (1998). 7. Lindsjo, H., Bai, X.S. and Fuchs, L., Environmental Comb. Tech, in press (2000). 8. Magnussen, B.F. and Hjertager, B.H., Sixteenth Symposium (International) on Combustion,The Combustion Institute,Pittsburgh,PA, 1976, pp. 7 19-729. 9. Lockwood, F.C., and Shah, N.G., Eighteenth Symposium (International) on Combustion,The Combustion Institute,Pittsburgh, PA, 1980, pp. 1405-1414. 10. Mann, A.P., and Kent, J.H., Combust. Flume, 99:147-156 (1994). 11. Wornat, M.J., Hurt, R.H., Yang, N.Y.C., and Headley. T.J., Combust. Flame, 100:131-143 (1995). 12. Winter, F., Prah, M.E., and Hofbauer, H., Combust. Flame, 108:302-314(1997). 13. Bai, X.S., and Fuchs, L., Computers and Fluids, 23:507-521 (1994). 14. Nordin, A., On the Chemistry of Combustion and Gas$cation of Biomass Fuels, Peat and Waste, Ph.D. Dissertation, Dept. of Chemistry, Umea University, Sweden.
917
No, Reduction of Biomass Combustion by Optimized Combustion Chamber Design and Combustion Control R. Padinger Joanneum Research Forschungsgesellschaft m.b.H., Elisabethstrasse 5, A-8010 Graz, Austria
ABSTRACT: NO, emissions of wood combustion significantly depend on the stoichiometric conditions in the heterogeneous reaction stage. Staging of the combustion process into a highly understoichiometric “primary combustion” (heterogeneous reaction in the glow bed) and an appropriate stoichiometric “secondary combustion” significantly reduces NO, emissions. These options can be realized by optimized combustion chamber design and combustion air control. In a 250 kW commercial wood chip m a c e such a combustion air control system has been installed. NO, emissions could be reduced by some 50 %.
INTRODUCTION Energy from biomass, especially wood, worldwide is regarded as one of the most important future renewable energy sources. However, the use of wood as a fuel requires low emission combustion systems. Research efforts in the past 15 years have produced furnaces with acceptable emissions of carbon monoxide and hydrocarbons 11-Y. Initial investigations on reduction of nitrogen oxides (NO,) emissions have been camed out during that time /4 - 61. A joint effort of several research groups was carried out in the JOULE I11 project “Reduction of Nitrogen Oxide Emissions from Wood Chip Grate Furnaces”, which has been presented in the 1” World Conference on Biomass for Energy and Industry in Seville 171. As a part of these efforts, this paper presents results of investigations on NO, reduction of wood chip b c e s by combustion chamber design in combination with combustion control, as one of several possible NO, reducing techniques.
EQUIPMENT AND METHODS The investigations have been carried out with a 250 kW commercial wood chip furnace. The furnace was originally designed in order to investigate measures for reaching complete combustion with a minimum of carbon monoxide and hydrocarbon
918
emissions Ill. In order to develop a furnace with both, complete combustion and low NO, emissions, the investigations for NO, reduction have been carried out on this furnace, carefilly paying attention that no restriction of the complete combustion already reached will happen caused by some NO, reducing measures. The furnace shown i n j g 1 a n d j g 2 is designed for the combustion of wood chps with a maximum chip length of approximately 5 cm. It is composed of a combustion chamber topped by a heat-exchanger.
Fig. 1 Fan Conveyor Fed Furnace heatexchanger
111
\\W
\ secondary ar fan
-TB
F ! primary air Iin
g-
Fig. 2 Sectional drawing of the Fan Conveyor Fed Furnace 919
Wall
The combustion chamber is equipped with a special feeding system throwing the combustibles by means of a fan blower (brand name: Awinator) mounted on the front of the combustion chamber through an opening into the combustion chamber. By this method, the combustibles are more or less regularly dispersed onto a horizontal grate, and eventually form a glow bed. This feeding process has the following advantages: The glow bed shows a locally highly homogenous consistency. It is free from zones of non-optimal combustion conditions (drymg zones, low temperature carbonizing zones). Each combustible particle is dropped directly into a fully burning environment. Given the dispersion of particles, the combustion conditions are not negatively altered by fresh material at the point of impact. The glow bed is free from mechanical disturbances, otherwise caused 'by conveyors (screw conveyors or side rods) or by a sporadic gliding down of the glow in underfeed stokers and step grate incinerators. Emission peaks resulting from such disturbances can be avoided, In this furnace type, the combustible particles are too light to cause such disturbances in the glow bed. The continuous feeding of combustibles is the basis of a continuous (in terms of automatic control engineering: of a stable) partial load operation. Via the load control system a stable performance control of the furnace without ordoffoperation (combustion for performance, combustion to keep the glow going) can be reached. An increase in emissions caused during the non-stationary phases of an operation by odoff-operation is also avoided. Given the low mechanical strain on the grate, it is possible to use drilled chamotte plates as a grate. These plates are suited for higher temperatures than metal grates, whch ensures optimal combustion conditions in the glow bed. The s u m of the drilling cross-sections in the chamotte plate is small in comparison with the surface of the grate. Therefore, non-burned particles are less llkely to fall through. This ensures full combustion in the solid phase and a low ash output. At the back of the combustion chamber a vertical deflecting wall of chamotte keeps the particles from entering into the burn-out-zones of the combustion chamber. The flue gas flows over the upper rim of this deflecting wall to be deviated downwards and eventually upwards by another chamotte wall (dumping wall). The flue gas is then directed through a flow channel between the combustion chamber and the heat exchanger to the front of the combustion chamber and from there into the heat exchanger itself. Just as with other fimaces, fuel control is based on the boiler temperature. This means that the thermal output momentarily used is measured, and on the basis of this thermal output and a supposed calorific value combustion efficiency, the required quantity of combustibles is calculated by an approximate method and finally set. On the basis of the temperature curve of the boiler, this calculated quantity of fuel is permanently adjusted. The advantage of such a control system is that the feeding of fuel can be adapted to the time curve of the thermal output almost instantaneously. This ensures stable controlling whch is necessary for a continuous operation without odoff-operation. On the other hand, a fuel control reacting exclusively to the boiler temperature or the flow temperature, only allows a very slow adjustment of the fuel quantity to the load curve, as a change in load becomes effective only after a certain time. Due to this dead time, there is in general no stable control without disturbance value feed-forward
920
and incinerators can only be operated by clocking. However, this regularly leads to heavily non-stationary phases presenting non-optimal combustion conditions. The continuous combustible control helps to avoid such an increase in emissions of h a d l substances. The combustion air control is in connection with the composition of the flue gas. Normally, for reachmg low carbon monoxide and hydrocarbon emissions without paying attention on NO, emissions, the concentration of oxidizable flue gas components, measured by a special sensor, served as control input for the secondary air. The primary air input could be varied in a certain range without significant d u e n c e on the thermal output (which is principally given by the fuel inlet). Less primary air of course leads to slower velocity of heterogeneous reactions, and therefore to a longer residence time of the solid phase. Thus, for reaching invariant heat output, the glow bed volume has to be correspondingly increased. Moreover, reduction of primary air leads to highly sub-stoichometric reaction conditions in the solid phase. This however seems to be one of the most important preconditions for low nitrogen oxide emissions of the finace. The effect of primary air reduction respecting control of NO, emissions has been studied in detail under different boundary conditions (load, water content of the fuel, etc), and for different fuels. The results tally well with results of similar experiments of other research institutions/8,9/.
INVESTIGATIONS ON THE INFLUENCE OF THE PRIMARY AIR INPUT ON THE NOx EMISSIONS To investigate the influence of the primary air input on the NO, emissions, experiments under different combustion conditions have been carried out. The results generally show, that reducing of the primary air input leads to a significant reduction of the NO, emissions.
EXPERIMENTS WITH CHIPBOARD CHIPS AS FUEL The influence of the primary air input on the NO, emission of chpboard chip combustion can be seen in Fig. 3, which shows a number of measurement results with the primary air input as the variable parameter. All the other combustion parameters (especially heat output respecting fuel input and 02-concentration of the flue gas) were held approximately invariant. The heat power was 250 kW. The normal operation of the furnace at h s heat output is specified by a primary air input between approximately 15 and 18* 10” m3/s. Caused by the high nitrogen content of the chipboards, the NO, emissions normally are in the range between 250 and 400 mg/Nm3. The average value of all the measurements is nearly 300 mg/Nm3. (All emission data of this paper are related to an oxygen content of the flue gas of 13 %). The results of the normal operation of the h a c e can be seen as the big cluster of measurement values in the right upper comer of the diagram in Fig 3. By decreasing the primary air input a significant reduction of NO, emissions can be reached. The diagram in figure 3 shows, that at a primary air volume flow of approximately 10*10‘3m3/s, NO, emission values in the range of 150 Nm3/s have been measured. The primary air reduction is limited by the primary combustion chamber volume. A lower primary air input causes a lower reaction velocity in the solid phase
92 1
of the fuel, resp. in the glowbed. Thus, for reaching the same heat output, the residence time of the glowbed must be increased, which leads to a bigger glow bed volume and therefore to a bigger volume of the combustion chamber. Furthermore, the combustion air needed in total for complete combustion in the desired heat power range has to be put into the furnace by secondary or even tertiary air control in order to avoid CO or C,H, emissions.
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50.-~--r---r---p---r---r---r---
Llu(e3El
The observations led to the expectation, that in the investigated range, a nearly linear relationship between residence time of the solid phase and primary air volume flow is given. This means, that for reaching NO, reduction in the range of 50 %, approximately doubling of glow bed volume is necessary. This, of course, will lead to an enlarging of the furnace. However, the amount of enlarging needed for avoiding half of NO, emissions is to be seen as need which should not lead to unsolvable technical or financial problems. Therefore M e r experiments with a 250 kW Fan Conveyor Fed Furnace have been carried out with the aim of investigating the possibilities of NOx reduction by combustion air controlling. The investigations have been carried out with chipboard chips, chipboard sawdust, and spruce wood chips as fuel. The results show in general, that a reduction of primary air can lead to a significant decreasing of NO emissions. The decreasing is relatively high if fuel with high nitrogen content (chip boards, chipboard sawdust) is used. The decreasing is relatively lower, if fuel with low nitrogen content (spruce wood chips) is used (reduction in the range of 30 %). However the absolute level of NO emissions when burning low nitrogen containing fuel ranges between 40 and 80 rng/Nm3. Under certain conditions the dependency of NO concentration on the primary air input s e e m to be inverted if the NO concentration is expressed in terms of ppm (as the primary measurement result) instead of m@m3 rel. to 13 % O2 (as used for
922
emission evaluation and comparison). This could be important for realizing a control system based on a NO measurement: The primary measurement results (ppm) in this case must be normalized before used for control purpose. Fig. 4 shows the NO emissions with chipboard chps as fuel depending on primary air volume flow at different power ranges: (1) 180 kW, 8 % O2 (2) 180 kW, 10 % 0 2 , and (3) 100 kW, 14 % 0 2 . NO (mg/Nmarelated to 13% 02)
--r--r---
.
"
P
0
7
10
20
30
40
50
60 70 80 90 primary air volume flow (10' ma/r)
Fig. 4 NO emissions with chipboard chps as he1 versus primary air volume flow
The diagram shows a certain decreasing of NO emissions by reducing primary air volume flow. (The desired O2 concentration in the flue gas has been reached by secondary air controlling). One can see, that the lowest NO emissions could be reached at high heat output, low 0, concentration, and low primary air volume flow. EXPERIMENTS WITH CHIPBOARD SAWDUST AS FUEL Fig. 5 shows the NO emissions with chipboard sawdust as fuel depending on primary air volume flow at different heat outputs:
(1) 200 kW, 8 % 02 (2) 180 kW, 10 % 02,and (3) 120 kW, 14 % 02, The diagram shows a siflicant decreasing of NO emissions by reducing primary air volume flow. (On the left certain measurement points outside of the marked area are shown, but they are of no significance in this experiment). One can see the same results as already shown in Fig. 3 and 4: Lowest NO emissions could be reached at high power level, low 0 2 concentration, and low primary air volume flow.
923
EXPERIMENTS WITHSPRUCE WOOD AS FUEL
Fig. 6 shows the NO emissions in mg/Nm3 (related to 13 % 02) with spruce wood as fuel depending on primary air volume flow at full load with: (1) minimal secondary air input, 6 - 8 % 02,and (2) maximal secondary air input, 12 - 14 % 0 2 .
0
10
20
30
50
40
60 70 80 90 primary air volume flow ( 1 0 ~m*/s)
Fig.5 NO emissions with chipboard sawdust as fuel versus primary air volume flow for different 02-contents in the flue gas. The diagram shows a certain decreasing of NO emissions by reducing primary air volume flow. As already mentioned above, caused by the lower nitrogen content of the spruce wood chlps, the absolute NO level as well as the relative decrease are much lower than in the experiments with chip board chips and chipboard sawdust.
924
NO (mglNm' related to 13% 0 2 )
20.-
-
1
I
I
~
0
Omax. secondary air (12...14% 02)
min. secondary air (6...8% 02)
0, 0
10
20
30
40
50
60 70 80 90 primaryair volume flow ( 1 P m'k)
Fig. 6 NO emissions in mg/Nm3(related to 13 % 0,) with spruce wood as fie1 versus primary air volume flow Fig. 7 shows the results of Fig. 6 but the NO concentration is expressed in terms of ppm unrelated instead of mg/Nm3, related to 13 % 02.One can see that in thls case the dependency of NO emission on the primary air input seems to be inverted. This could be important for realizing a control system based on a NO measurement: The primary measurement results (ppm) in this case must be normalized before used for control purpose. NO (mglNm' related to 13% 02)
I 20'-
-
I
I
I
I
I
I
I --
~
@
min. secondary air (6...8% 0 2 )
Omax. secondary air (12...14% 02)
01 0
10
20
30
50
40
60 70 80 90 primary air volume flow (lo3 m%)
Fig. 7 Results of Fig. 6 but the NO concentration is expressed in terms of ppm instead of mg/Nm3,related to 13 % O2
925
OPTIMIZED COMBUSTION CONTROL SYSTEM
Following the above mentioned results, an air staging computational tool which controls the combustion air input in a way, that a minimum of primary air input is given in order to minimize the NO, emissions has been developed. The scheme of this control system is shown in Fig. 8.
Fig.8 Scheme of an optimized combustion control system FUEL CONTROL The fuel amount is in a separate control circle by the heat demand or the boiler- resp. reflux temperature only. PID algorithms ensure stable control behavior without odoff-operation A/. PRIMARY COMBUSTION AIR CONTROL The primary combustion air is also controlled in a separate control circle in a way, that maximum glowbed volume (glowbed height) is reached. Glow bed height can be measured more or less easily by an infrared barrier with a modulated signal. This, as described above, ensures, that in any operation conditions minimum NO, emissions are reached. SECONDARY COMBUSTION AIR CONTROL
The secondary combustion air is also controlled in a separate. control circle by the composition of the flue gas, especially the CO-concentration, resp. the concentration 926
of oxidizable gas components I l l . This parameter can be monitored easily by using ceramic sensors (normally used in fire alarm systems). They are cheap and very suitable for use in flue gas, up to 250 "C. Special control algorithms are used to search the operation mode of optimum combustion air mlet. Thus minimal emissions of oxidizable flue gas components (CO, C,H,) are reached at minimum oxygen level resp. highest efficiency. CONCLUSIONS On a commercial Fan Conveyor Fed furnace with a thermal output of approximately 250 kW investigations on primary measures for NO, reduction have been carried out. The results show, that decreasing of primary air leads to a significant reduction of NO, emissions up to 50 % of the previous value. The reduced primary air input of course causes lower reaction velocity in the solid phase of the fuel, resp. in the glow bed. Thus, for reaching the same heat output, the residence time of the glowbed must be increased, which leads to a bigger glow bed volume and therefore to a bigger volume of the combustion chamber. The results led to the conclusion, that to reach a 50 % NO, reduction, approximately doubling of the glow bed volume is necessary. Furthermore, secondary air control is proposed to be applied in order to avoid CO or C,H, emissions. The investigations have been carried out with chipboard chips, chipboard sawdust, and spruce wood chips as fuel. NO, reduction reaches 50 % for fuel with high nitrogen content (chip boards, chipboard sawdust). The reduction is lower, if fuel with low nitrogen content (spruce wood chips) is used (reduction in the range of 30 %). The absolute level of NO emissions when burning fuel with low nitrogen content ranges between 40 and 80 mg/Nm3. The absolute level of NO emissions when burning fuel with high nitrogen content however ranges between 150 and 500 mg/Nm3. Under certain conditions the dependency of NO concentration on the primary air input seems to be inverted if the NO concentration is expressed in terms of ppm unrelated instead of mg/Nm3,related to 13% 02. This could be important for realizing a control system based on a NO measurement: The primary measurement results (ppm) in this case must be normalized before used for control purpose.
REFERENCES 1. Padinger R. (1990) Control system for solidfirelfirrnaces, 5' European Conference - Biomass for Energy and Industry, Lisbon (P), October 9-13 1989, in Elsevier Applied Science Publishers, London New York, 1990, pp. 2570 - 2574 2. Good, J. (1992) Verbrennungsregelungbei automatischen Holzschnitzelfeuerungen PhD Thesis ETH 9771, Swiss Federal Institute of Technology, Zurich. 3. Nussbaumer, Th. (1998) Eficiency Improvement and Emission Reduction by
Advanced Combustion Control Technique (ACCT) with CO/Lambda Control and Setpoint Optimization, Biomass for Energy and Industry, lomEuropean Conference and Technology Exhibition, Wiirzburg (D), June 8 - 11 1998, pp. 1362 - 1365.
927
4. Keller, R. (1994) Primarmassnahmen zur NO, Minderung bei der Holzverbrennung mit dem Schwerpunkt der Luftstufung, PhD Thesis ETH Nr. 10514, Ziirich. 5. Nussbaumer, Th., Salrmann, R. (1996) Primary measures for NOx reduction in wood combustion by air staging and firel staging. 1'' European Conference on Small Burner Technology and Heating Equipment, Zuerich (CH), September 2526 1996, pp. 165-174. 6. Glarborg, P. Dam-Johansen, K., Miller, J.A. (1995) The reaction of Ammonia with Nitrogen Dioxide in a Flow Reactor: Implications for the NH2 + NO2 Reaction. Int. J. Chem. Kin.,Vol. 27, pp. 1207-1220. 7. Padinger R., Lecher B., h n d L.-E., Thunman H., Ghirelli F., Nussbaumer T., Good J., Hasler P., Salzmann R., Saastamoinen J., Oravainen H., Heiskanen V. P., HWllinen J., Taipale R., Bilbaoi R., Alzueta M. U., Millera A., Oliva M., Ibifiez J. C., Kilpinen P. (1999) Reduction of Nitrogen Oxide Emissionsfiom Wood Chip Grate Furnaces. JOULE-project JOR-3CT96-0059, presented in the lst World Converence and Exhibition on Biomassfor Energy and Industry, Seville (E), June 5 - 9 2000, (Proceedings in press). 8. Good, J., Nussbaumer, Th., Schaffner, H.P. (1998) NOX Reduction in Biomass Combustion by Combination of Air Staging and SNCR Technique. Biomass for Energy and Industry. 10* European Conference and Technology Exhibition, Wiirzburg (D) June 8 - 1 1 1998, pp. 1360 - 1361. 9. Saastamoinen, J.J., Hamiillinen, J.P., and Kilpinen, P. (1999) Release of nitrogen compoundsfi-om wood particles during pyrolysis. Environmental Combustion Technologies, pp.1-28.
928
Polycyclic Aromatic Hydrocarbons Associated to Particle Size Emitted from Biomass Fluidised Bed Combustion F. Saez’; A. Cabafias’; A. Gonzalez’; R. Escalada’, J.M. Martinez, J.J’. Rodriguez-Maroto’, J.L. Dorronsoro’, F. Gomez’ and D. Saenz’ I DepaTmento de Energias Renovables. DER-CIEMAT, Avd. Complutense, 22. 28040-Madrid. Spain Departamento de Combustibles Fbsiles. DCF-CIEMAT, Avd. Complutense, 22. 28040-Madrid. Spain
ABSTRACT The size distribution of polycyclic aromatic hydrocarbons (PAHs) on different particle size fractions in the fly ash emissions from biomass combustion have been measured by gas chromatography (GC) and mass spectrometer (MS). The experiments were carried out in a 1 MWth atmospheric fluidised bed combustion pilot plant, at different operation conditions. Emitted‘fly ash was collected with an Andersen cyclone cascade system from the stack. The hgest concentration of the dangerous PAHs were found on particles having aerodynamic diameter <2.0 pm. The objective of this study, is to expand our knowledge of PAH formation and characterization, emitted in the combustion of a determinate biomass (poplar tree)
INTRODUCTION All biomass combustion processes produce emissions whose composition contains CO, 02,C 0 2 ,NO, NO, and organic compounds including volatile organic compounds (VOCs) and polycyclic aromatic hydrocarbons. Biomass combustion is a great source of PAHs due to the high content in volatile compounds. Polycyclic aromatic hydrocarbons and their derivatives are part of a great group of toxic compounds and some known carcinogens. The possible health risk to humans through inhalation or ingestion of these PAHs makes necessary the determination of these compounds, to reduce its concentration in the emissions. Most of three-ring PAHs are present in the gas phase emissions but compounds with three, four, five and six ring PAHs are associated to fly ash biomass combustion emissions. Due to polycyclic aromatic hydrocarbons are adsorbed to particulate matter and some of them are carcinogenous pollutans, the effect on heath depended of PAHs concentration and distribution in the different size particle, with special mention on respirable fraction with a diameter size less of 7 pm.
929
Fluidised bed combustion has been chosen to study the pollution emission for these advantages: is adequate for large variety of biomass with different physicalchemical characteristics, high combustion efficiency and low pollution emission. The PAHs characterization knowledge by size particle distribution provides information to incorporate improvement to the gas cleaning system for the biomass combustion process.
EXPERIMENTAL METHODS MATERIAL
Poplar tree (short rotation crop) was used as fuel in the combustion experiments. The biomass was air dried. The characterization analysis was made with a homogenous mixture of leaves and stems milled to a 2 mm size. The material was a crushed using a forage harvester-crusher machine, and was then reduction by using a hammer mill. The following analytical tests were made for poplar characterization: proximate analysis (Fixed carbon, volatile matter, ash content), ultimate analysis (C, H, N, S,02,Cl), and heating value. BURNING CONDITIONS
The experiments were carried out in a 1 MWth atmospheric fluidised bed combustion pilot plant. The general technical characteristics of the plant are detailed as follows: (a) Feeding Systems. The pilot plant has three feeding systems. Two of them for feeding biomass and the third for feeding inert and reactive materials. Each feeding system consists of a bin and a helical screw feeder with speed variation. The biomass bin (3 m3) is equipped with mechanical and pneumatic facilities to prevent the formation of bridges. All bins screw feeders unload the materials into other screw feeder that introduce the materials into the combustor. (b) Atmospheric Bubbling Fluidised Bed Combustor The combustor is cylindrical with a 1.14 m internal diameter in the fluidised bed region, 1.O m2 surface area and a 4 m height. It consist of: 0 A distribution plate with more than 500 tuyeres. 0 Vertical heat exchangers assembled on the internal wall of the combustor. 0 Three horizontal heat exchangers placed on the freeboard, 3 m above the distribution plate. 0 The air supply s stem consists of two blowers, one of 1500 Nm3/h and the other of 500 Nm /h flow rate. 0 Two biomass sampling inlets. (c) Heat exchange to reduce the flue gas temperature from 700°C to 150°C. (d) Gas cleaning system. It consists of a multicyclone separator with an air preheater and a fan. (e) Gas analysis equipment from HartmannLkBraun. ( f ) Distributed control system (Contronic P of H&B).
Y
930
Several experiments were carried out to establish the adequate combustion operation: combustion temperature, air flow rate, air excess, fluidisation velocity, and feeding system for the type of biomass. Long duration and more stable operation conditions experiments were selected to establish in this work To optimize the combustion process, an adequate control of flow feeding, bed temperature and air excess has been stablished. The control of these parameters yields a high efficiency (>99%)and a clean combustion process. SAMPLINGSTATION A particle sampling and measuring system was set up to characterize the suspended particles in the biomass combustion (Fig. I). The system was capable of on-line measuring of the concentration number and particle size distribution in the micron and submicron range. In addition, mass concentration, particle morphology and chemical composition were obtained. The aerosol is first extracted from the duct and an
I
I
HOT BOX (I) she cWrlbtbn idmmhl comeotrslbn (0.7cDprlOmm).
(V) M u comentralbn unlckmlc.1 umlysb.
(9M o r p o b g y d r k m e a t s l coaposkba
(VI) Sze d#hlb.tlon d pmbcr eooccntritbn (DClmm).
(111, Iv) Aerodynmlr LlrhItloq mum Comeatrith idcbrrnk.1 u18lyaIh
Fig. I Schematic diagram of the sampling and measurement station. For the micron size range an optical on-line particle sizer was used, working on the principle of white light scattering on single particles. T h ~ sinstrument provided number concentration and size distribution measurements in the range 0.3-30 pm, with a maximum of allowed particle number concentration of lo5particles/cm3. Next, the sampling flow was guided to different measurement instruments. The concentration number and the size distribution of submicron particles (0.01-0.5 pm) was obtained with a TSI Scanning Mobility Particle Sizer (SMPS) consisting of a differential electrical mobility analyzer coupled with a condensation nuclei counter. Previously a preconditioning gas system was used. This basically consists of a cyclon,
93 1
a sampling chamber, a dryer, a preimpactor and an aerosol neutralizer. The cyclon removes the particles higher to 2 pm. The sampling chamber allow us to extract a second sample from the original without disturbing the isokinetic conditions. This sample is dned to avoid condensation in the DMA. The preimpactor precisely removes the particles that are larger than 1pm. Lastly, the aerosol neutralizer, containing 2 millicuries of Krypton 85, is designed to neutralize electrostatic charge on aerosol particles as they pass through the device. A device was designed to collect particle samples deposited on glass slides. The particles were deposited on the slides by impactation, sedimentation and diffusion, Subsequently, the samples were analyzed by scanning electron microscopy (SEM). Information on morphology and size distribution was obtained from image analysis, while the elemental composition of particles was determined by electron probe X-ray microanalysis (EPXMA). In addition, the aerodynamic particle size distribution was measured with two types of instruments: a stage cascade impactor with fiber glass paper substrates and a backup filter; and a five-cyclone train that collects particulate material. This set-up has several advantages: high collection capacity, bigger sample for accurate gravimetric and chemical analysis, and longer sampling times for better averaging. The five cyclone train also includes a backup filter that collects fine particles. In both of these instruments, the particles are classified into different aerodynamic sizes, (300.1 pm). Additionally, 47-mm fibber glass paper filters were used to measure the total mass concentration of aerosol. The filters, the substrates and the deposits retained in the cyclone collection cups were gravimetrically and chemically analyzed with gas chromatography and mass spectrometer. EXTRACTION
After the samples were collected the cyclone collection cups were carefully unloaded. Extreme care is required in this procedure to insure that all particulate deposits are recovered and placed in the proper sample container. The particles adhered to the internal surfaces of each cyclone were brushed and washed with methylene chloride to ensure the recovery of all particulate matter. Methylene chloride is used because it is considerably more volatile than the particles and it evaporate completely before weighing. The six collected fractions were extracted by soaking them in methylene chloride for 18 hours in a soxhlet apparatus after the addition of deuterated PAHs surrogate standard (acenaphthylene d8,pyrene dloand benzo[a]pyrene d12). The solvent was evaporated at 35°C using a turbovap (LV evaporator) under a flow of purified nitrogen. The residue obtained was redisolved in 500 pl of toluene. The glass material was previously cleaned with an organic solvent and placed in a muffle furnace at 500°C for 5 hours before use. The glassware was cleaned to remove organic traces and other pollutants. The organic solvent used in the extraction was supplied by Burdick & Jackson (pesticide grade). Internal standards were added to the samples before the analysis. The recovery of the surrogate standard was used to monitor and evaluate potential losses in extracted compounds.
932
ANALYSIS The analysis of polynuclear aromatic hydrocarbons pollutants was performed on a Fisons Instrument Gas Chromatograph GC-800 series, connected to a Fisons Instrument Mass Selective Detector (MS) MD-800. The analyses was performed with a 30m x 0.25 mm fused silica capillary column coated with a 0.25 pm thick film (DB5 J&W Scientific). Helium was used as carrier gas for analyte separation. An aliquot of 1 pl was injected in splitless mode. The temperature of the column was held at 70°C for 1 min, rate 7°C min-' to 310°C for 5 min. The electron impact ionization energy was 70 eV. The ion source temperature was 275°C. The mass selective detector was operated under ion monitoring mode (SIM). The quantitative analysis was carried out using a 16 PAHs (EPA 610) external standard and a solution of 5 deuterated PAHs as the internal standard, whch contained: naphtalene-dg, acenaphthene-d phenanthrene-dI o, chrysene-d and perylene-dLz. The PAHs analysis was made according to a modification of the Environmental Protection Agency (EPA) Method 610 and 625 and the NIOSH Manual of Analytical Method 55 15 (EPA, 1989;NIOSH, 1994). RESULTS AND DISCUSSION
FUEL COMPOSITION AND OPERA TION CONDITIONS
Poplar tree residue was pretreated to an adequate particle size of 5-15 cm in length and 1-20 mm in diameter on average, which is a suitable dimension to be used in the FB combustor. Table I shows the characterization of the biomass assayed.
Table 1 Composition o f the raw material. Sample Moisture (% d.b.) Proximate analysis (%d.6.) Fixed carbon Volatile Ash Ultimate analysis (% d.b) Carbon Hydrogen Nitrogen Sulphur Oxygen Chloride Heating value. (MJkg) Gross (0,O % d.b.) Gross (13,5 % d.b.) Net (0,O % d.b.) Net (133 % d.b.)
Poplar tree 13.5 18,43 79,9 198 48,2 5.9 0,57 0,05
43,49 <0,01 19,6 17,O 18,3 15,5
933
Its high volatile contents (79.9 % d.b.), low amounts ash (1.8 % d.b.), and heating value of 18.3 MJkg, make the poplar tree appropriate to be used in the fluidization bed combustion process. It noted the low content in sulphur (< 0,05% d.b.) and chloride (0.01 % d.b.) of
the biomass as an advantage for this process. Four experiments were carried out in the fluidised bed combustion pilot plant with a 16- 18 hour test time and whose operation conditions were modified in order to obtain an efficient combustion process. These two conditions yield lower pollutant emissions. Two operation variables were changed mainly in this study: feeding rate and air excess percentage. These two parameters, in turn,affect the others parameters. Table 2 shows the combustion operation conditions set in the experiments. In Table 3, the flue gas composition during the four tests is shown. The sampling flue gas temperature was 150"C, the probe, cyclone and filter was all held at the same temperature. The CO level emitted in Test 2 with a 40-60 air excess percentage is lower than in the other experiments. Table 2 Biomass burning operation conditions. Test 1
Test 2
Test 3
Test 4
235
220
185
220
40-70
40-60
85-130
20-40
780-820
830-850
800-850
770-780
1.13
1.o
1.2
1.13
Feed stock
K O Air excess %
Bed temperature "C Fluidisation velocity d
S
Table 3 Composition of combustion flue gas (dry and 6% 02). Test 1
Test 2
Test 3
Test 4
6-1 1
6-8
9-12
4-6
12-15
13-15
9-1 1
13-17
(mg/Nm3)
4000-6000
1000-2500
4000-6000
8000- 10000
NO (mg/Nm3)
150-225
150-225
350-500
200-300
>50
-40
25-20
125-175
0 2
("33 vol)
co2 (% vol.)
co
SO2
(rng/~m~)
Test 4, with a low air excess percentage (20-40%) was insufficient to burn the biomass with a 6000-10000 mg/Nm3 CO emission. On the other hand, Test 3 was perfomed under a high air excess condition (85-130 %) causing a large formation of CO (4000-8000 mg/Nm3).
934
Two of these tests have been considered for the comparative study to analyze the PAH size distribution: Test 2 that was ran under more stable operation conditions and Test 4 that showed increases in CO, PAH and particulate emission levels. The samples collected during the selected tests were gravimetricaly measured and the 16 PAH priorized by EPA were analyzed by Gas Chromatography-Mass Spectrometry. The major goal of this study was to evaluate the distribution of dangerous PAHs by particle size. CHAM CTERISATION OF THE EMISSIONS
Table 4 shows the total mass concentration (g/Nm3)in the two tests. Table 4 Particle mass concentration Total mass concentration ( g . m 3 Test 2 Test 4 0.12 0.35
In order to detennine the concentration number and size distribution two size intervals were considered. One corresponding to a particle diameter range from 0,Ol to 0,6 tun, measured with the SMPS; and another range from 0.4 to 20 pm measured with the optical instrument (PALAS). 6.0E+06
I
,
,
1
0.1
10
100
DPb m)
Fig. 2 Distribution of particle size : 0.01 pm ID, I0.6 pm. Figure 2 presents the particle emission size distribution for Tests 2 and 4 measured with SMPS. It is possible to verify the presence of one only mode in the submicron size range, although other nanometric modes could exist below 0.01 micron. The detected mode is located in the 0.2 to 0.3 micron range, and its geometric mean diameter (GMD) is approximately 0.2 pm. Notice that, although the particle
935
mass concentration emitted in the Test 4 is lower than the Test 2, the total concentration number is higher in the second test. Figure 3 shows the particle size distributions measured with the optical instrument.
These results correspond to the Test 2 and Test 4. Notice that these measurements correspond only at a fraction of the complete mode, related to the bigger particles of that mode. From t h s and the results of the SMPS (Fig. 2), it can be deduced that the distributions are unimodal. The mode is located under 1 micron. Cascade impactors and cyclones have been used in order to determine the aerodynamic size distribution. Impactors allow the classification of particles with an aerodynamic diameter D, between 0.1 pm ID, S 5 pm, whle cyclones work in the 2 to 20 pm range. ......
0,001
~
0.01
.
~
~. ..~ . . .. . . . . . . . ~ ~...... ~ ~ ~ ...~. ~ . ~ . ~~........
0,1
1
10
I00
DP (p)
Fig. 4 Aerodynamic particle size distribution : 0.1 pm ID, I 5 pm. Results obtained with cascade impactors.
936
Fig 4 shows the distribution obtained with the cascade impactors. It can be seen, as in the Fig. 2, that in the 0.1-1 pm size range the mass concentration is higher in Test 2 than in Test 4 although considering the total measurement interval, the size distribution in Test 4 is higher than in Test 2. Both aerodynamic distributions are one part of the main mode. More than 70 percent of the mass emission is formed by particles smaller than 0.1 pm. More than 90 % are smaller than 1 micron. Notice that, although the particle mass concentration emitted in the Test 2 is lower than the Test 4, the total concentration number is higher in Test 2.. Additionally the concentration number measured with the optical counter is higher in Test 2. An efficient combustion, as occurring in Test 2 favours the presence of fine particles as opposed to the poor combustion of Test 4 which produced a high concentration of coarse particles, thus increasing the mass concentration. PAH ANAL YSIS
Data on the PAH size distribution collected in the six stages with cut-points from 0.01 to 14 microns are represented in Table 5 and Table 6. Left y-axis represents the corrected distribution at the interval width and right y-axis the PAH's concentration at the interval width. The values are averages of three data points on the same sample. Data are obtained with respect to the volume of sampled air (ng/Nm3).The volume of air passing through each stage was the same. The distribution histograms for the four, five and six ring PAHs was represented in Fig. 5 and Fig. 6. 1
50'0 Fluormthene S 40.0
f-P
-
30.0
a
0 a 20.0
5
8
10,o 00
0.01
0.1
1 DP (rm)
10
40,O
30.0
-P F
=D a
g
0.0 0.01
4 s 3
100
f
Bea
g.
a
5
5 50
1
-
00 01
DP bun)'
1
10
100
10
100
30.0
:
3
0.01
0.1
40.0
200
0 g
10.0
loo
5
150
20.0
0
25 0
-P
10
Pvrene
DP (rm)
E
z
1;
10
I
I--~ommmn
20.0
10.0
1
10
0.1
1
DP (rm)
Fig. 5 Size distribution of four rings PAH in fly ash collected in cascade cyclone sampler. Test 2.
937
Most of the three-ring compounds were adsorbed by particles collected in stages 1 through 4, with a size range of 3.5 to 14.2 pm, with a high concentration present in stage 3. Four, five and six rings PAHs are mainly distributed into particle size smaller than 2.0 pm. The concentration of high molecular weight PAH increased while the particle aerodynamic diameter decreased. 50.0
8enropjfiuoranihen -cmm~'tw
20
7
20,o
IdOdbdDp)
............................
f
15
0.01
0.1
1 DP (rm)
10
f
15.0 ..........................
-,-
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0.1
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Indeno[l,2,3z d] pyrem
1-1,
@Jo
0.1
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6
-
5
3
b'
1w
DP (rm)
.....................................
0,Ol
l...I..I i
10
too
Zmbbn WdwIDPl
...............................
0.01
0.1
1
10
DP(w)
DP (14
....................................
roo
...................................
8
I i.
.................................
0.01
0.1
1
10
100
'0.01
DP (w)
0.1
1 DP (run)
10
100
Fig. 6 Size distribution of five and six rings PAH in fly ash collected in a cascade cyclone sampler. Test 2.
Comparing Tables 5 and 6 it can be seen that under more stable combustion operation conditions, the high molecular weight PAH concentration was decreased. Approximately 50 'YO three ring PAH decreased from Test 4 to Test 2. 65 ?LO Benz[a]anthracene (four ring) and 45 % benzo[a]pyrene diminished from Test 4 to Test 2. Taking into account, from 16 PAH compounds analysed in this study, the following substances are known to posses a high carcinogenic potential: benzo[a]pyrene, moderate carcinogenic activity is manifested by benzo[k]fluoranthene and benzo[g,h,i]perylene, the PAH value reduction obtained optimizing the operation conditions is noted.
93 8
Table 5 PAHs concentration (ng/Nm3) in particulate fractions of the cyclone sampler. Test 2
Naphthalene Acenaphtylene Acenaphthene Fluorene Phenanthrene Anthracene FIuoranthene Pyrene Benzralanthracene Chrysene Benzo[b]fluoranthene Benzo[k]fluoranthene Benzo[a]pyrene Indenor 1,2,3-cd]pyrene Dibenz[ahlanthracene Benzo[ghl]perylene
Etl(pm) Et2(pm) Et3(pm) Et4(pm) EtS(pm) Etf(pm) >14,2 14,2-8,6 8,6-6,4 6,4-3,5 3,5-2,0 <2,0 112.03 89.93 187.50 159.12 47.39 71.54 1.28 1.47 0.74 1.21 1.15 1.05 0.62 1.oo 1.96 1.43 0.46 0.50 2.3 1 2.87 2.18 2.47 2.43 2.79 19.12 11.41 21.38 24.18 23.29 19.82 1.13 1.10 1.02 1.03 1.05 0.75 8.84 5.44 3.89 6.37 6.08 6.39 6.5 1 2.91 2.22 3.69 3.43 4.79 4.26 0.72 0.89 0.74 2.79 0.87 12.72 2.91 3.46 2.98 2.77 8.49 17.70 3.40 3.65 3.05 9.32 1.69 5.91 1.36 1.92 1.23 4.45 0.92 5.45 1.21 1.46 3.13 4.57 0.65 17.86 5.76 2.29 6.10 1.98 34.22 2.82 0.99 0.78 1.13 4.66 0.96 24.58 8.85 3.25 10.43 12.17 59.33
Table 6 PAHs concentration (ng/Nm3) in particulate fractions of the cyclone sampler. Test 4
Naphthalene Acenaphtylene Acenaphthene Fluorene Phenanthrene Anthracene Fluoranthene Pyrene Benzralanthracene Chrysene Benzo[b]fluoranthene Benzo[k]fluoranthene Benzo[alpyrene Indenor 1,2,3-cd]pyrene Dibenz[ah]anthracene Benzo[ghi]perylene
Etl(pm) Et2(pm) Et3(pm) Et4(pm) Et5(pm) Etf(pm) >13,9 13,9-8,3 8,3-6,l 6,l-3,4 3,4-2,0 <2,0 4.01 82.24 41.11 25.46 378.73 288.45 0.52 1.70 3.42 2.73 3.37 2.40 0.03 0.92 0.80 1.05 1.19 0.97 1.24 2.40 4.42 2.36 3.74 3.13 22.21 24.69 61.47 36.30 47.85 31.65 1.27 1.21 3.96 2.88 2.99 2.10 6.78 5.52 15.61 9.86 12.60 13.05 4.14 3.87 10.91 8.10 10.19 11.13 0.46 0.62 3.91 2.85 3.48 5.03 1.51 1.86 9.06 5.97 7.85 12.50 2.20 2.77 8.92 6.11 8.47 16.14 1.39 1.97 5.58 4.10 5.25 8.63 1.29 1.82 8.09 6.01 6.96 13.48 9.45 4.72 60.96 15.95 20.12 40.00 0.25 0.42 2.41 1.99 2.63 4.26 6.59 7.84 32.15 24.43 33.18 64.97
CONCLUSIONS -The particle emissions in the studied combustion process, present a unimodal distribution with the mode situated between 0.2 and 0.3pm.
939
0
0
0
0
-Under stable combustion conditions the particle emission in the size range 0.1- 1 pm is increased with respect to the results obtained in less favourable combustion conditions. -A high combustion efficiency considerably reduces the PAHs emissions. -High concentration of four-,five-and six-ring PAHs is measured in small particle size distribution. -The major presence of high molecular polycyclic aromatic hydrocarbon in the lower size particle demands a more efficient cleaning system.
REFERENCES 1. Rodriguez-Maroto J.J.; Gomez, F.J.; Martin, M. and Bahillo, A. (1995) Aerosol Sci, 26, S685.
J.
2. Bemhard, R. (1981) Rapid Measurement of Particle Size Distributions by use of Light Scattering Methods, Paper presented at PARTEC Nuremberg May 6-9 (POLYTEC). 3. Knudson, E.O.; Whitby, K.T. (1975). J.Aeroso1 Sci, 6, 443. 4. Agarwai, J.K; Sem, G.I. (1980). J.Aeroso1Sci, 11, 343. 5 . Lodge, J.P.; Chan, T.L. (1986). Cascade impactors Sampling & data Analysis. American Industrial Hygiene Association: Akron, OH. 6. Gyula Kiss, Zita Varga-Puchony, Gabor Rohrbacher, Jozsef Hlavay. Atmospheric Research 46 (1998) 235-261 7. Jiigen Schenelle, Kathrin Wolf, Gerhard Frank, Bernhard Hietel, Itsvan Gebefigi and Antonious Kettrup. Analyst, September 1996. Vol121 (1301-1304). 8. J.M. Martinez, R. Escalada, J.M. Murillo and J.E. Carrasco. Biomass for Energy and Indusby, pp 1420-1423 (1998). 9. F. Saez; A. Cabaiias; A. Gonzdlez and R. Escalada. Proceeding of the 10th European Conference. Biomass fon Energy and Industry, 1998, Ed. C.A.R.M.E.N, 1417.
940
Fuel staging for NO, reduction in automatic wood furnaces R. Salzmann* and Th. Nussbaumer**+
*SwissFederal Institute of Technology, LTNT, 8092 Zurich, Switzerland 'Verenum, Langmauerstrasse 109, 8006 Zurich, Switzerland
ABSTRACT NO, emissions from biomass combustion originate mainly from fuel bound nitrogen. Fuel NO, can be reduced by air staging and fuel staging. Air staging for wood fuels was investigated earlier and has found its way into practice for understoker and grate furnaces. However its application is limited, since a significant NO, reduction is only achieved in a narrow range of the stoichiometry and at a relatively high temperature (i.e. 1100 O C - 1200 "C). Fuel staging has not yet been applied for non-pulverized solid fuels. The aim of the present work was to investigate the potential of fuel staging for non-pulverized biofuels. An novel understoker furnace of 75 kWth with two fuel beds in series was developed. Experiments were performed with wood chips and UF-chipboards. The influences of stoichiometric ratio, temperatures, residence time, and fuel properties on the conversion of fuel nitrogen to NO, and other Nspecies was investigated. Fuel staging was found to enable a significant NO, reduction at lower temperatures than air staging, i.e. above 900 OC. Under optimum conditions up to 78 YOreduction was achieved for UF-chipboard and up to 66 % for wood chips. Under typical combustion conditions, no side-products were found in the flue gas. Fuel staging is regarded as a promising technology for NO, reduction in automatic biomass furnaces enabling a broader range of application than air staging. The need of two feeding systems and advanced process control leads to additional costs, while no reducing agent is needed as it is the case for secondary measures. Fuel staging can also be combined with air staging, enabling part load operation under air staged conditions.
INTRODUCTION
Fuel bound nitrogen is the main source of NO, emissions from biomass combustion'. Due to relatively low temperatures in biomass furnaces, thermal and prompt NO, formation are not relevant. In Switzerland - as in many other countries - the concentrations of ozone and P M 10 (particluate matter < 10 pm) in the ambient air regularly exceed the imniission I i its. Therefore measures for the reduction of NO, and PM 10 are needed to achieve the goals of air pollution control. Since biomass furnace lead to far higher emissions of NO, and PM 10 than light oil boilers or natural gas boilers, the reduction of these two emission parameters is of high priority to improve the environmental advantages of
7
94 1
biomass as energy source. This is illustrated in 'Table I' which shows the environmental impact of wood, oil, and gas for room heating in Switzerland according to the Ecological Scarcity Method2. The results, which are based on a Life Cycle Assessment (LCA), indicate that 38.6 % of the environmental impact of a modern automatic wood furnace is due to NO,, and that the environmental impact of wood is smaller than of light oil boiler but higher than for natural gas.
Table I Environmental Impact Points (EIP) according to the Ecological Scarcity Method for heating systems with wood chips, light oil and natural gas2. For wood modern automatic furnaces and the use of non contaminated wood fuel is assumed, while for oil and gas combustion modern low-NOx boilers are assumed. Oil
Wood
SOX,NH,, CH4, NMVOC,
[EIPIGJ]
[Yo]
[EIPIGJ]
13030
38.6%
6 190
12600
36.5%
650
Gas
["h]
[EIPIGJ]
[%]
13.8%
3410
13.0%
130
0.5%
1.5%
670
2.0%
18 200
40.7%
13 300
50.9%
8 200
22.9%
19 760
44.0%
9 260
35.6%
34500
100%
44800
100%
26 100
100%
primary energy, residues, and others Total
Total compared to wood
100%
130%
76%
During the recent years, primary and secondary measures have been developed to reduce the formation of nitrogen oxides during the combustion process3. Staged combustion techniques have been investigated for large scale power plants, fired with gas, oil, or pulverized coal. Since the importance of renewable energy sources will grow in the future, clean combustion of solid biomass is becoming more and more important.
CONVERSION OF FUEL NITROGEN
'Fig. 2' shows a simplified global reaction path of fuel nitrogen. During the thermal degradation of the biomass the main products containing fuel nitrogen after this conversion process are HCN and NH.: Further reactions then lead to NO when they are oxidized by the reaction with radicals (0, H, OH) or to molecular nitrogen N2. However only part of the nitrogen is converted to NO,. The goal of staged combustion techniques is the promotion of the reaction path leading to N2 by the creation of optimized conditions during combustion. The reduction of fuel-N to molecular nitrogen in air staging is favored in the fuel rich primary combustion zone ( 'Fig. 3 Investigations on fixed bed wood furnaces have shown that a minimum of the Total Fixed Nitrogen (TFN = HCN+NH3+NO+ N02+2N20) emission from the primary combustion zone is reached for a stoichioC to 1200 "C and providing a metric ratio of 0.7 to 0.8 and a temperature of 1100 ' mean residence time of 0.5 ss. After the reduction zone the combustion is completed in the burnout zone by injection of the excess air. I).
942
Air staging has high demands concerning the fuel properties. Because of the relatively high temperatures needed in the reduction zone for air staging, ash sintering and ash slagging can lead to operational problems and limit the application of fixed bed combustion. Native wood usually has a high melting point (i.e. around 1300 "C). The melting point of other wood fuels as UF-chipboard or demolition wood and of other biomass can be much lower, depending on their composition. Herbaceous fuels for example have melting points at 800 "C to 900 "C because they contain relatively high amounts of inorganic substances (mainly K)6.
Zeldovichmechanism
Fig. 2 Simplified reaction path diagram of the NO, formation and destruction in the gas phase. Fuel stagin or reburning has first been tested on coal fired utilities using natural gas as reburn fuel . In fuel staging the first stage of the combustion process operates slightly fuel lean, whereby the NO, production is high ('Fig. 3 7. Then, additional fuel is added which creates fuel-rich conditions in the reburn zone. When hydrocarbons are used as reburn fuel, the hydrocarbon radicals entering the reburn zone can initialize the NO reduction mechanisms. The major reaction path found from research studying the reburning mechanism through elementary chemical reactions under fuel rich conditions is the formation of HCN':
9
C, CH, CHI, CHi => HCN ... HCN + 0, OH => N2 + ... Because the volatile content of wood is high (80 % - 85 %), wood is highly reactive and therefore well suited as reburn fuel. Also the nitrogen content may be beneficial since it leads to additional reducing NHi + N O => N2 + ...
(3)
Further improvement of the reburning in power plants lead to the advanced reburning technologies were fuel staging is combined with other NO, reduction technologies".
943
A
stage 1
stage 2
SR,< 1
SR>1
flue gas
B
Fig 3 Principles of air staging (A) and fuel staging (B). AIM The aim of the present work is to investigate whether a new concept for fuel staging based on fixed bed systems is suitable for NO, reduction of non-pulverized wood fuels”. Focus of the research is the elaboration of the optimum combustion conditions for an efficient NO, reduction. The measured potential is compared with air staging. All the experiments presented are measured on a pilot-scale research facility of about 75 kW thermal input. To investigate the effect of the fuel nitrogen two fuels with different nitrogen content are used for the experiments: wood chips (W) with low nitrogen content and UF-chipboard (UF) with high nitrogen content, ‘Table 3 The parameters investigated are the temperatures, the stoichiometric ratio and the residence time in the reburn zone and the ratio of primary fuel and reburn fuel. The research facility also allows air staging conditions as well as unstaged combustion. I.
EXPERIMENTS A new research facility was designed to investigate both air staging and fuel staging in a fixed bed furnace shown in ‘Fig. 4’. The main components of the facility are two understoker grates in series. The alignment chosen for the furnace tries to approximate a straight plug flow of the gases, preventing dead zones and backflow regions. The first stage is a conventional understoker furnace equipped with an air supply for primary and secondary air. Then, a small burnout chamber for the combustion gas follows. The design of the second grate is adapted to the higher gas flux due to the temperatures of 800 “C to 1200 O C . The maximum fuel input of both understoker systems is equal. The second understoker is protected by a water cooled shell. To minimize the heat loss of the gas to the cooled surface the understoker hosing is sheltered with bricks. Above the second grate, the reburn or reduction zone follows. At the entrance a static mixing element provides an enhanced mixing of the gases before passing the reburn zone. The reburn zone element has also a water cooled shell which keeps the reburn zone at low temperatures. The heat transfer can be influenced by the composition of the insulation layers. The reburn section has three levels of nozzles for the injection of the
944
burnout air. According to the injection levels, the mean residence times can be varied: depending on the fuel inputs and temperatures from about 0.4 s to 2 s. The reburn section also has several openings at different locations for gas sampling and temperature measurements. On the top, the burnout segment leads the gases further to the boiler unit of the facility. The lower part of the boiler unit is designed as combustion chamber. The upper part is the heat exchanger where the flue gas is cooled down to approx. 200 O C .
Fig. 4 Research facility: 1. understoker (stage I), 2. ash box, 3.,4. air inlet, 5. primary zone, 6 . burnout zone, 7. understoker (stage 2: reburn stage), &air inlet (level variable), 9. mixing element, 10. reduction zone, 11. burnout zone, 12. flyash box, 13. boiler. Besides the ash removal, the operation of the research facility is completely automatised and controlled by a PLC. The first fuel stage is equipped with an oxygen control to keep the oxygen content within a narrow range (SR,= 1.1 to 1.3). For fuel staging, it is essential that the stoichiometric ratio of the first stage is well controlled to enable understoichiometric conditions in the reduction zone with little reburn fuel 945
input. The amount of reburn fuel is dosed according to the desired stoichiometric ratio in the reburn zone. The injection of burnout air completes the combustion after the reduction zone. The amount of burnout air is controlled the way that complete burnout is achieved at a stoichiometric ratio between 1.5 and 2.0. The oxygen content in the flue gas is also continuously measured with a second gas sensor. A flue gas ventilator controls the pressure in the facility and guarantees safe operation. A more practical parameter than the stoichiometric ratio in the reburn zone is the reburn fuel rate RFR, which is the input of reburn fuel to the total fuel input. RFR = m2 LHV2 / (ml LHVl + m2 LHV2)
(4)
where m l , m2 are the mass flows and LHV,, LHV2 are the heating values of the fuels of stage one and two. When the stoichiometric ratio of the first stage is known, the stoichiometric ratio in the reburn zone can be calculated as a function of RFR. ‘Fig. 5 ’ shows the relation between RFR and the stiochiometric ratio in the rebum zone SR,.
1.4
- LHVJLHVZ = 1 .O
--
----.
0
0.1
0.2
0.3
LHVjILHV2 = I .8
0.4
0.5
RFR [-I
Fig. 5 Stoichiometric ratio in the reburn zone as a function of RFR. The measurements on the research facility were carried out at stationary or quasi stationary conditions. The measurements of air flows, gas temperatures, gas composition, and heat output were analysed continously and monitored online. The gas composition was analyzed in the flue gas after the boiler exit with industrial gas analyzers. For the analysis of the hot gas in the reduction zone a suction pyrometer combined with a probe for O2 detection was used. With this probe also short fluctuations could be monitored with extremery short delay. Besides, a hot gas sampling line with different analyzers for measuring the gas in the reburn zone was installed. ‘Table 2 ’ gives on overview over the gas analysis equipment. The experiments were carried out with wood chips and UF-chipboard (UF = Urea Formaldehyde), representing two biofuels with different nitrogen content. UF-chipboard has a very high nitrogen content compared to native wood due to the bonding agent, ‘Table3’.
946
Table 2 Gas analysis equipment for the measurements on the research facility. LOCATION Flue gas
SPECIE 0 2
NO, NO2 co, c02
Reburn zone
CXH,
c o , c02 CH4 H2 0 2
NO, NO2 HCN, NH3, N2O
METHOD paramegnetic CLD NDIR F ID NDIR NDIR HCD paramagnetic CLD FTIR
Table 3 Fuel properties of wood chips and UF-chipboard. WOOD CHIPS
Water Ash wf Volati les wf Char wf HHV LHV C wf H wf 0 wf wf N S wf NO2 max from N at 11% 0 2 Bed ash properties: Initial deformation temperatue Hemispherical temperature Fluid temperature
dl 0% g/lOOg g/IOOg dl oog MJ/kg MJlkn g/IOOg g/IOOg gl100g gl100g g/IOOg mgmrn3
17.9 1 .o 83.0 16.0 16.13 14.74 49.6 5.34 44.9 0.18 0.03 720
1280°C 1310°C 1360°C
UF-CHIPBOARD 4.7 1.3 79.9 18.8 18.9 17.54 49.2 6.0 41.1 3.58 0.12 12300
1230°C 1260°C 1290°C
wf: water free base
MODELING The aim of the modeling was to investigate the influence of stoichiometry and temperature in the reburn zone on the NO, reduction during the fuel staging experiments. Focus is given on the homogeneous chemistry during reburning and influence of the operation conditions on the nitrogen oxides and their precursors. The chemical reacting system is described as a combination of ideal reactor types (plug flow reactor PFR and perfeclty stirred reactor PSR). The reactors in question are the reburn and burnout section of the fuel staged combustion. The resulting furnace model used with all the possible variants for simulation is shown in ‘Fig. 6’. The input of the model are the gaseous combustion products of the first stage and the reburn fuel. The combustion
947
process after the reburn fuel input is separated and modeled as plug flow or mixed flow or both. For the conversion of the solid reburn fuel complete fuel conversion is assumed. The model for the conversion of the fuel bound nitrogen during pyrolysis into HCN and NH3 is based on experimental data of single particle pyrolysis at 800 OCi3, ‘Table 4 ’ . Since the correlation found between the HCN/NHj ratio and the fuel-O/fuel-N ratio is valid for various fuels and nitrogen containing model substances, it is also base for the implemented model. The conversion of the wood under fuel rich conditions is calculated assuming chemical eq~ilibrium’~. Under adiabatic conditions at 800 “C, not all of the reburn fuel is converted into gaseous products at pyrolysis state. An amount of char depending on the water content of the fuel and the process temperature remains. In a real furnace in fact, the charcoal burns out completely, also in the reburn stage since char burnout takes place after the pyrolysis. According to this fact, the amount of oxygen needed for the char burnout is computed and subtracted from the amount of O2 available in the gas coming from the first fuel stage.
Fig. 6 Scheme of furnace model for simulation. The main advance of the chosen model is the possibility of handling detailed reaction mechanisms for investigating the gas phase reactions. There are several reaction mechanisms available for natural gas (methane) combustion including nitrogen chemistry. The mechanism selected for the present model is the GRI-Mechanism V2.1 1 (49 species, 279 primary reactions)’’. For modeling the gas phase reactions of the furnace described model, the CHEMKIN I1 software package was usedi6. The generation of the input and output data of the different processes is accomplished with separate input routines. The aim of the model simulations is to predict the influence of the process parameters on the nitrogen species as a group, and not the calculation of absolute emission values. Therefore, a reference value named reduction rate R is introduced, allowing a dimensionless presentation of the results. The definition of R is R = 1 - (TFN / TFNo)
(5)
948
where TFNo is the input of all nitrogen species to the reburn zone. Assuming that all nitrogen species besides NO and NO2 emitted from the first fuel stage are negligible, this value corresponds to the measured NO, emission of the first fuel stage.
Table 4 Conversion of fuel nitrogen into HCN and NH3 assumed for simulation. WOOD CHIPS
O/N Conversion to HCN Conversion to NH3
223 0.03 0.05
UF-CHIPBOARD 11 0.007 0.005
RESULTS AND DISCUSSION The experiments with air staging carried out on the research facility confirm the NO, reduction potential shown in previous investigations. The experiments proof that the stoichiometric ratio in the reduction zone has a distinct influence on the NO, emission of the combustion process, providing that optimum conditions for temperature, mean residence time, and mixing rate are also accomplished. The experiments indicate the important role of the fuel rich zone created by air staging. 'Fig. 7 ' opposes air staging and conventional combustion and shows the NO, reduction potential of air staging, when the NO, emissions of air staging are correlated to the corresponding total stoichiometric ratio. The NO, reduction is about 72 % for UF-chipboard, meaning from 550 mg/Nm3 to 150 m o m 3 NO, at 11 % 0 2 . For wood chips, the reduction measured is about 65 %, corresponding to a NO, reduction from 260 mg/Nm3 to 90 mg/Nm3. The influence of the stoichiometry in the reduction zone is shown in 'Fig. 8'. A summary of the experimental results for fuel staging with different fuel combinations is presented in 'Fig. 9'. The measured NO, emissions are plotted as a function of the temperature in the reburn zone for different reburn fuel rates. The CO emissions of the evaluated measurements were always below 500 mg/Nm3 at 11 YO02,usually below 250 mg/Nm3. The measurements demonstrate that the NO, emission can be reduced significantly with reburning. Around 750 OC the NO, emissions already begin to decrease. Minimum emissions for UF-chipboard of 120 mg/Nm3 are reached, whereas for wood chip combustion less than 100 mg/Nm3 at a temperature between 800 "C and 900 "C are measured. The emission level for wood chip combustion is therefore comparable to the one measured with air staging. The NO, emissions reached for the combustion of UFchipboard are about 20 % to 25 YOlower than with air staging as 'Fig. fa' shows. The results indicates that the optimum stoichiometric ratio in the reduction zone is slightly higher for fuel staging than for air staging, meaning about 0.85 compared to 0.8. The optimum temperature range for NO, reduction begins above 800 O C for all fuel combinations tested. However NO, reduction begins to be effective already above 700 "C. A decrease of the reduction potential with increasing temperature cannot be seen within the measured data, although the data of fuel staging with wood chips at wood chip combustion (W/W) indicate a minimum at about 800 "C. If the indicated temperatures are compared it has to be taken into account, that the presented data were measured with thermocouples, while the real gas temperature, measured with the suction pyrometer are up to 100 OC higher.
949
1200 1
UF-chipboard
Aair staging 0 conventional
1.o
1.5
2.0
2.5
3.0
3.5
4.0
SR 1-1 Fig. 7 NO, emissons as a function of stoichiometric ratio for conventional and air staged combustion of UF-chipboard. 500 UF-chipboard
450 Q 0 400
0
wood chips
8
F: 350
Y
300
5E 250 Y
P v)
2oo
150
0
B
0”100
z
0
50
0 0.40
0.60
0.80
1 .oo
1.20
SRr [-I
Fig. 8 Influence of stoichiometric ratio in the reduction zone in air staged combustion of wood chips and UF-chipboard. The residence time of the gas in the reduction zone is calculated with the approximated gas flow and the given geometry. The maximum residence time is z (air nozzles at position 3), which corresponds to a mean residence time of about 2 s for most of the experiments. If the air is injected through the nozzles at position 2, the resulting mean residence time is shortened to about 5/9 z. The measurements indicate little influence of
950
the residence time, i.e. the NO, emission at 5/9 z is about 10% higher compared to the maximum residence time.
250 9
---+--MI,
RFR = 0.3
+WMI,
RFR = 0.4 UFMI, RFR = 0.2
+UFMI, RFR = 0.3 - - t - -UF/UF, RFR = 0.25 - - - -UF/UF, RFR = 0.3 1-A
o! 600
700
800
900
1000
1100
1200
T, ["CI
Fig. 9 Influence of temperature in the reburn zone on NO, emission for different RFR: UF = UF-chipboard; W = wood chips; CO < 500 mg/Nm3 at 11% O2 for all data, SRI = 1.2 to 1.3. The simulation results show the calculated influence of the reburn temperature and the reburn fuel rate on the reduction rate R. To simulate the influence of the mixing conditions, once a simple plug flow (PFR) is imposed and once a combination of a mixed flow (PSR) and a PFR with different mean residence times, i.e. the the mean residence time is splitted between the mixed reactor (PSR) and the plug flow reactor (PFR). The total mean residence time for the reburn zone and the burnout zone models is fixed for all cases to 2 s. The simulation with the GRI-mechanism indicates that the decomposition of HCN and NH3 follows the reactions: HCN + 0 = NCO + H HCN + OH = CN + H20 HCN + 0 = CN +CO HCN + OH = HNCO+H NH3 + OH = NH2 +H2O NH3 + H = NH2 +H2 NH3 + 0 =NH2 + OH
95 1
This reaction path is valid over the entire investigated temperature range. Towards low temperatures, almost all NO is reduced, but the model predicts an increasing formation of N 2 0 and NO*. 500 I ';3
1
Ofuel staging UFNV, 900°C
450
0
X air staging UF
400
X
r
5 350 27
X
E 300 -
5E 250 Y
N
200
-
0 150-
:1 0 0 -
v)
e
50-
o l 0.40
0.60
0.80
1.00
1.20
SR, [-I
Fig. I0 Influence of stoichiometric ratio in the reduction zone on the NO, emissions for air staging: UF-chipboard: T, = 1050- 1080"C, CO < 250 mg/Nm3 at 1 1% 02) and fuel staging: UF-chipboard main, wood chips reburn fuel: T, = 900"C, CO < 300 mg/Nm3 at 11% 02. With increasing temperature more and more HCN and NH3 is decomposed, HCN in higher rates than NH3. The N 2 0 formation is reduced drastically. The higher decrease of the reduction rate for RFR=0.25 is due to a diminished NO reduction. The main cause is the recombination of HNO with H, OH, or 0 to NO. However, with increasing temperature at low RFR the reaction rates of NO forming and reducing reactions fall. The temperature influences also the reaction velocities. Regarding the nitrogen species forming and consuming reactions, the characteristic reaction time at 800°C is about 10 times slower than at 1300 'C. For comparing the measured NO, emissions with simulation results, the modeling also has to include the burnout stage. This means that the gas of the fuel rich reburn zone has to react with the burnout air to the flue gas. For this purpose, the output data of the reburn simulation are used as input data for the burnout modeling. 'Fig. 11 ' shows an example of the calculated reduction rate after burnout assuming plug flow in the burnout zone (isothermal PFR at 1300 "C,2 s mean residence time). The two curves in the figure represent two different reburn models, once PFR only and once PSR plus PFR in series. The figure shows that the maximum reduction predicted varies between 82 % and 95 %, depending on the model for the reburn zone. However, the calculated reduction rates or higher than the measured. The burnout conditions have an important influence on the emission of the nitrogen species. The analysis of the simulations indicates that at high burnout temperature NH3 and HCN are destroyed leading only to a weak increase of the NO, emission. N20
952
which is formed at low temperatures in the reduction zone is also reduced to a negligible amount at this burnout temperature. The emission of NO2 is small compared to NO. 100% 90% 80%
g
70% 60%
Y
50%
7
40% 30% 20% 10% 0% 600
700
800
1000
900
Tr
1100
1200
1300
1400
I"c1
Fig. 14 Calculated reduction rate after burnout zone for RFR = 0.30 (SR,.=0.85) for
UF/UF (UF-chipboard main and reburn fuel). The observed change of the reduction rate at low temperatures may be influenced by the burnout modeling. A different temperature regime in the burnout zone for example may modify the reduction rates for low temperatures in the reburn zone. But calculations assuming isothermal plug flow at 1100 "C and 800 "C show that the burnout temperature has less influence on the over all reduction rate, providing that the preliminar reburn temperature is higher than 900 "C. Similar observations are reported from experiments on a laboratory plug flow reactor". The comparison of the measured data with the calculated NO, reduction indicates that the theoretical potential of fuel staging is even higher than the experiments on the research facility have demonstrated. However, the trends generally show a good agreement. The lower reduction rates at low temperatures confirms the role of the temperature not only in the reburn zone, but also in the burnout zone. Additionally, the simulations points out the importance of the mixing and flow conditions within the furnace.
953
CONCLUSIONS
The present investigation demonstrates that fuel staging can be successfully applied for the combustion of non-pulverized wood fuels in fixed bed combustion. For UF-chipboard as reburn fuel a NO, reduction of up to 78 % was achieved. Fuel staging was applied with a Reburnd Fuel Rate RFR between 0.2 and 0.4, meaning that 20 % to 40 % of the total energy input was provided by the secondary fuel. The temperature range for an efficient reduction of the nitrogen species by homogeneous reactions is significantly lower than for air staging. The experiments show that low NO, emissions (without formation of N 2 0 ) are already achievable at temperatures above 900 OC. Therefore, fuel staging may be a favorable method for fuels with a high nitrogen content. Fuel staging is also favorable for fuels with high ash content and low ash fusion temperature because it allows to keep the combustion temperature of the first fuel stage well below 1000 "C. The temperature control can be accomplished by heat extraction. As the temperature in the burnout zone after the reburn zone should be high enough to ensure complete burnout of the gas, the transferred heat can also be used to preheat the burnout air if necessary. All reburn fuels investigated showed a NO, reduction potential at comparable optimum temperature and stoichiometric ratio. For high reduction rates, a mean residence time in the reduction zone of about 1.5 s regarding the examined facility is provided. When a fixed bed is used for reburning with non-pulverized fuels, a high voidage of the fuel bed is beneficial to the reburning because a deep penetration of the combustion gas enhances and accelerates the fuel conversion rates and improves the mixing of the gas phase. For this reason, effects or conditions that make the passing of fuel bed for the gases more difficult, i.e. a dense fuel bed or slagging conditions, should be avoided. Concerning the design of the reburn zone, its shape should promote the gas mixing and prevent an inhomogeneous flow field. Fuel staging demands additional investments in the hardware and advanced process control of the furnace. On the other hand, costs for additional reducing agents and catalysts are saved in comparison to SCR and SNCR techniques and undesired side products such as HNCO, NH, and N 2 0 (which can be a major concern in Denox processes) have not been found. Additionally, fuel staging can be combined with air staging. If a sophisticated management of the furnace switching between air staging and fuel staging is applied, this may result in advantages of the part load behaviour compared to other NO, reduction techniques. The kinetic model for the gas phase reactions in the reburn and bournout zone enables the description of the influence of the reburn temperature and stoichiometry on the nitrogen species and hence is a suitable tool for the qualitative study of the influence of the main parameters. The simulation predicts a higher NO, reduction potential under ideal conditions than measured.
ACKNOWLEDGMENTS
The present project was carried out in cooperation with Tiba-Muller Lfd. (Balsthal, Switzerland) which is gratefully acknowledged. The investigation was funded by the Swiss Federal Office of Energy and the Swiss Federal Office for Ediicatioii and Sciences and it was part of the EU-Project JOR3CT96 ,Low-NO, Wood Chip Combustion' which was coordinated by Joanneum Research (Austria).
954
REFERENCES 1. Nussbaumer, T. (1 989) Schadstoffildung bei der Verbrennung von Holz. PhD
thesis ETH No 8838. Zurich 2. Kessler, F. and Frischknecht, R. (2000) Heizenergie aus Heizol, Erdgas oder Holz, Schriftenreihe Umwelt Nr. 3 15, Bundesamt fur Umwelt, Wald und Landschaft (Swiss Federal Office of Environment), Berne 2000 3. Nussbaumer, T. (1997) Primary and secondary measures for NO, reduction in Biomass combustion. In Developments in Thermochemical Biomass Conversion; Blackie Academic and Professional. 4. Aho, M.; Hamaleinen, J.; Tummavuori, J. (1993) Importance of solid fuel properties to nitrogen oxide formation through HCN and NH3 in small particle combustion. Combustion and Flame, 95, pp. 22-30. 5. Keller, R. (1994) Primarmassnahmen zur NO, Minderung. PhD thesis ETH No 10514. 6. Kaufmann, H. (1997) Chlorine compounds in emissions and residues from the combustion of herbaceous biomass. PhD thesis ETH No 12429. 7. Smoot, L.; Hill, S.; Xu H. (1998) NO, control through rebuming. Prog. Energy and Combustion Science, 24, pp. 385-408. 8. Miller, J. & Bowman, C. (1989) Mechanism dnd modeling of nitrogen chemistry in combustion. Prog. Energy and Combustion Science, 15, pp. 287-338. 9. Rudiger, H.; Greul, U.; Spliethoff, K. (1995) Pyrolysis gas of biomass as a NO,reductive in a coal fired test facility. In 3rd International Conference on Combustion Technologiesfor a Clean Environment, Lisbon (Portugal). lO.Kicherer, A.; Spliethoff, K.; Maier, H.; Hein, K. (1994) The effect of different rebuming fuels on NO, reduction. Fuel, 73, pp. 1443-1446. 1 1. Zamansky, V.; Maly, P.; Ho, L. (1 997) Family of advanced reburning technologies: pilot scale development. In Joint Power Generation Conference; ASME; pp. 107113. 12. Salzmann, R. (2000) Fuel staging for NO, reduction in automatic woodfurnaces. PhD thesis ETH No 13531. 13. Hamaleinen, J. (1995) Effect of fuel composition on the conversion of fuel-N to nitrogen oxides in the combustion of small single particles. PhD thesis, University of Jyv&kyla. 14. Salzmann, R.; Good, J.; Nussbaumer, T.; Leiser, 0. (1998) Temperature reduction by flue gas recirculation in biomass combustion with air staging: modeling and experimental results. In 10Ih European Conference & Technology Exhibition, Wurzburg (Germany). 15. http://euler.berkeley.edu/gri_mech/. 16. Kee, R.; Rupley, F; Miller, J. (1990) Chemkin 11: Fortran chemical kinetics package for the analysis of gas phase chemical kinetics. SAND89-8009, Sandia National Laboratories. 17. Kristensen, P.; Glarborg, P.; Dam-Johansen, K. (1996) Nitrogen chemistry during burnout in fuel staged combustion. Combustion and Flame, 107, pp. 2 1 1-222.
955
Estimate of the net CO, reduction by replacing coal and oil with biomass in Japan DOTE Y.*, OGI T.**, and YOKOYAMA S.** *Miyazaki University, Gakuen Kibanadai Nishi I -I, Miyazaki 8892192, JAPAN **NationalInstitutefor Resources and Environment, Onogawa 16-3, Tsukuba, Ibaraki 305-8569, JAPAN
ABSTRACT: We estimated the net C02 reduction, Rv, by replacing fossil fuels (coal and oil) with imported biomass by considering four C02 release processes: (a) plantation of biomass fiom field preparation to harvest, (b) collection of harvested biomass, (c) biomass transport to power generation plant, and (d) pretreatment of biomass fbr power generation. Combustion-steamturbine power generation and gasification-gasturbine power generation were considered as biomass energy conversion technologies. The gross COZreduction, &, was calculated as avoided C02 emissions fiom coal and oil combustion with biomass taken as carbon neutral. Under standard conditions, where electricity capacity was 25 MW,biomass yield was 10 t-biomassrhaly, and biomass transport distance was 20000 km, the maximum RN was 258 kgC/t-biomass (325 k g - C m ) for replacing coal with gasification of biomass, and the minimum RN 61 kg-C/t-biomass (100 k g - C m ) fbrreplacing oil with combustion of biomass. IfC02 release unit of the ship for transport of biomass was larger than 1 . 3 ~ 1 0kg-CWt-B -~ or electricity capacity was smaller than 4 MW, Rv was negative. To achieve the C02 reduction in Japan decided by COP 3 by using biomass, the am ofplantation for imported biomass was only 10-20%ofthe fbmt area that was lost between 1990 and 1994 worldwide. Total Co2 emission from the above processes was 50-80 % of&: the biomass transport process contributed to 80 % ofthe total C02 release. RN was independent ofthe carbon emission panmetem for collection of harvested biomass, biomass yield, standing period of biomass, and pretreatment of biomass fix power generation. INTRODUCTION
The C02 emission limitation and reduction for each country was agreed by COP 3 in 1998; for Japan, the COZemission of 94 % of the 1990 emission base year was decided. In the present, there has been no eiktive method to reduce C G emission in Japan although improving energy eaFciency and introducing eneqy saving technology and 956
new energy such as solar cell have been discussed. Biomass energy is expected to be one of the efective methods fir CQ reduction because biomass energy is &on neutral. In Japan, however, there has been no discussion of the use of biomass energy because it is dilmlt to establish large plantation a m s in Japan. The esediveness ofthe use ofimported biomass in Japan also has never been discussed. Although the kne-work methodology with to assess the effectiveness ofbioenergy systems in reducing net CO, emission was reported' and there is a little research on CO, emission f b m plantation and COZ reduction in power generation by replacing fossil &el with biomass%', the estimate of the net CO2 reduction by considering a total system induding plantation to power generation is s d y reported"9. Moreover, there is no discussion of the e i k t of long distance transport on net C 4 reduction. The purpose ofthis papa is to estimate the net COZ reduction by replacing coal and oil with biomass in Japan and the area of plantation for imported biomass, not to propose a calculation method 6 r reducing net COZ emission. Combustion-steam turbine power generation and gasification-gas turbine power generation were considered as biomass energy conversion technologies. METHODOLOGY The biomass power g e n d o n system considered in this paper comprised of five processes: (a) plantation ofbiomass fiom field preparation to hawest, (b) collection of harvested biomass, (c) biomass transport to a power generation plant, (d) pretreatment of biomass fbr power generation, and (e) biomass power generation. COZrelease per dry biomass h m the pmxsses (axd) was ndkned to as RP, RCB,RTB,and RPT, respectively, in kgC/t-B. The total ofthis COZ release was r c h d to as RT. COZ reduction per dry biomass in a biomass power generation process was r e h d to as gross COZ duction, RG in kgC/t-B. Net COZ reduction in the total system, RN, was calculated by the following equation; RN=&RT &B was calculated as &llows: Considering plantation area like a doughnut with an outer radius, R, and inner radius, RO of 0.5 km, harvested biomass was collected to the center ofthe plantation a m The am of plantation was d m d to as A, and the standing period ofbiomass as S. The plantation area was divided into S sectors, and one ofthe sectors was harvested annually. This sector was radially divided into A4 belts. The width ofthe belt was equivalent to the length ofthe side ofa square with am (Ay) equivalent to the loading capacity (LcB)of the tractor used 6 r colledion of harvested biomass. The m-th belt was d d d y divided into N m units; the area of the unit was equivalent to A.. The distance h m the center ofthe plantation am to the center of the unit was Feserred to as d, (Figw 1). The distance &r annual collection of biomass (0.was ) described as follows; D.= 2 (N,*d,) RCBwas described as follows; RCB=CCB*D~*LCB/(A *Y) where CCBis the COZ dease unit of the tractor (kgC/km/t-B), and Y the yield of biomass (t-Bhdy). Rra, WBS described as follows; Rn~=cm*&~ 957
whae C m is the CCh dease unit ofthe ship (kg-Clkm/t-B), and D m the distance of biomass transport (km).
&was calculated as follows;
I
I‘
M belts
’I
Figure I System for collection of harvested biomass.
& 6 c*( 17 81 17 c)*& whe~e6’~ is the COZ mission &or ihr coal or oil (kg-CIGJ) including upsemissions, V B electricity daency fir biomass (-), o c electricity e6aency Eor d or oil (-), and HEheating value of biomass (GJIt-B). D B W ~ calculated S follows; 17 B= a *CAB when CAis the electricity capacity. A and A. were calculated as follows; A= WJY A.=AtSl( WJLCB) where W, is the aund biomass weight needed 6 r power gagation of biomass and calculated as fbllows; Wa=C~*3600*24*365*104/(H~* 17 c)
RESULTS AND DISCUSSION OBTAllvEDPARAMETERS Pannnaers used as standard conditions 6 r calculation m listed in Table 1. Parenthesea in Table 1 express the range of pannneters obtained by reviewing the literature. For R p and R ~ T ,avaage values wen used. The value of Cn,was 6 r coastal transpott. For Dra, halfofthe a r r w n h c e of the earth was used, which is considered to be the maximum transport distance of imported biomass in Japan. For Y and H’,typical values wed. For CA,25 MW used b m it was the economic limitation’’. For S, 10 yeam was assumed. For OC, the value in Japan was used. For a and 0 , the values were obtained @om a mgmsion m e by plotting data tiom reviewing the literature (Figure 2).
NET CARBON DIOXIDE REDUCTION &and CCh dense tian each p m s , which werecalculated by using the stendad parameters, am listed in Table 2. The result that & was positive in any case shows that the use of imported biomass in Japan is e&tive in reducing C G emission in Japan. The maximum & was 258 kg4lt-B (325 kgG/MWh,) 6r q l a i n g coal 958
with gasification ofbiomass, and the minimum RN 61 kg-C/t-B (100 k g - C m ) 5br replacing oil with combustion ofbiomass. Table I Paranleters used for the estimate of net CQIreduction.
IElectricity e&bncy for coal or oil,
' """'I
I
' """'I
'
r) c
(-YE
.39
""I
0.179C:'26
0.1
1
10
100
lo00
0.1
1
10
100
1000
CAW)
b) Gasification
a) combustion
Figure 2 Relationship betweenelectricity &ciency and electricity capacity (CA).
959
Table 2 Estimated RN and COZrelease using standard parameters(kg-C/t-B). Oil Fuel coal Technology Combustion Gasification Combustion Gasification 303.0 392.8 Gross COZ Reduction, RG 385.4 499.8 241.5 241.5 241.5 Total CO, release, RT 241.5 32.0 32.0 32.0 32.0 Plantation, Rp 6.9 6.9 6.9 6.9 Pretreatment, RPT 0.6 0.5 0.6 Collection, RCB 0.5 202.0 202.0 202.0 202.0 Transport, RTB 143.9 258.3 61.4 151.4 Net CO, reduction, RN Coal replacement was more &ive in reduang COZ emission than oil replacement, which is due to the difkmce in BC between coal and oil. Gasificationwas more eikctive in reducing COZemission than combustion because rl B fix gasification was greater than that for combustion.
CONTMBUTION OF THE COz RELEASE PROCESS There was no di&ence in RT between coal and oil replacement or between biomass power generation technologies. RTwas 50-80 % of& and contributed significantlyto decreasing&. Because DT was much longer, COZrelease &omthe biomass transport process contributed to 80 % of&. This means that decreasingCTBis necessary to increme&.
EFFECT OF PARAMETERS For Rp, RPT,CCB,LCB,CTB,Y, and S, they had a wide rangq only one data was obtained, and/or they were arbitrarily assumed. Sensitivity analysis fir the parameterswas perfbnnedto show the etkct ofvarying the parameters (Table 3).
Replminn ooal with gasitication Factor Cn,
RP RPT Cm
h,S Y
0.25 1.586 1.093 1.020 1.002 l . m 0.998
0.5 1.391 1.062 1.013 1.001
1 1.OOO 1.OOO 1.OOO 1.OOO
0.999
1.OOO
l.m l.m
1.5 0.609 0.938 0.987 0.999 l . m 1.OOO
2 0.218 0.876 0.973 0.998
l.m 1.001
Replacing oil with wrnbustion
0.25 3.466 1.391 1.084 1.007 1.OOo 0.990
0.5 2.644 1.261 1.056 1.005 1.m 0.9%
1 1.ooO 1.OOo 1.OOO 1.OOO
l.m 1.OOO
1.5 -0.644 0.739 0.944 0.995 1.OOo 1.002
2 -2.289 0.478 0.888 0.990 l . m 1.003
The specific reduction in Table 3 means the ratio of RN calculated by changing one parametcx ranged h0.25 to 2 times the standard value to RN calculated by using standard parameters. RN was very dependent on CTBbecause of the long transport distance. Especially, fir replacing oil with combustion of biomass, RN became negative by using a 1.5 times m e r CTB, which means the use of imported biomass can not contribute to the reduang ofCG emission in Japan. Using CTBof 1.3~10” kg-CWt-8, RN became zero; thexefbr, biomass was carbon neutral. 960
RN was also dependent on RP;however, RN was independent ofRpr, CCe,La, Y , and S. EFFECT OF ELECTRICITY C4PACITY Figure 3 shows the e&t of CA on Rv. For replacing cod with gasification of biomass, RN was positive at CA of O.O05MW, whereas, fir replacing oil with combustion of biomass, RN was positive at CA of 4MW. This means that CA of greater than 4 MW is necessary for reducing COZ emission in Japan.
AREA OF PLANTATION FOR IMPORTED BI-OICIASS
Area of plantation for imported biomass to reduce C02 emission by replacing coal and oil was estimated as shown in Table 4. CO2 emission in Japan in 1995 was 332 Mt-C and that &ompower generation was 100 Mt-Cm. The contribution of coal and oil to COZemission ofthe total power generation was 25 % and 64 o/oZ', respectively, so that it is estimated that Cot emission h m cod-fired power generation was 25 M t C and that h m oil-fired power genedon 64 MtC. Because there was 320 MtG of COZ emission in Japan in 199OP, COZ reduction decided by COP 3 is estimated to be 32 Mt-C in 1995, Assuming that all coal-fired power generation (25 Mt-C) and 7 M t C equivalent oil-fired power generation are replaced, the area h r plantation is estimated to be 29 Mha fir combustion of biomass and 14 Mha fir gasification of biomass. The area is equivalent to 10-20% of the forest area that was lost between 1990 and 1994 worldwide (137 Therefbre, we conclude that 500 400
300 n
013
$
200
J?
100
U
z
LLZ
0
-100
0.001
0.01
1
0.1
10
100
c, (Mw) Figure 3 EfE.43 of electricity capacity (CA)on net COz reduction (RN). 961
1000
it would be possible to establish the area of plantation in foreign countries, although it is necessary to analyze the implications of actually supplying the biomass needed fir energy production in Japan.
Table 4 Area of plantation for imported biomass to reduce COr emission in Japan.
oil
61.O
1Sl.Ol
114.8
46.41
11.5
4.6
CONCLUSIONS
Net COZreduction in Japan by replacing coal and oil with biomass was estimated by considering a total biomass energy system. Using standard value of parameters, the maximum net COZreduction was estimated to be 258 kg-C/t-B (325 k g - C m ) hr replacing d with gasification ofbiomass, and the minimum RN 61 k g C / t B (100 k g - C m ) fbr replacing oil with combustion of biomass, which indicates that, even ifbiomass is imported, the use ofbiomass energy makes COr reduction possible in Japan. Using CIZIof 1.3~10”k g - C M t - B or CAof4Mw, however, RNbecame zero, which means that biomass was d o n neutral under these conditions. The area of plantation fir imported biomass was estimated to be 10-20 % of the finst area that was lost between 1990 and 1994 wotldwide. This indicates that the use of imported biomass to reduce C02 emission in Japan would be feasible.
REFERENCES 1. Schlamadmger B., Apps M.,Bohlin F.,Gustavsson L, Jungmeier G., Marland G., Pingoud K.,and Savolainen I. (1997) Towards a standard methodology hr
greenhousegas balances of bioenergy systems in comparison with fossil energy systems, In: L?evelopnts in tknnochemiml biomass conversion (Ed. by A.V. Bidwater and D . G . B . B o o ~ k ) pp. , 359-375, BLACKIE ACADEMIC & PROFESSIONAL. 2. Wright L. L. and Hughes E. E. (1993) U.S. Carbon offset potential using biomass energy systems, Water,Air, and Soil Pollution, 70, pp. 483-497. 3. Gustavsson L., Borjesson P., Johansson B., and Svenningsson P. (1995) Reducing C02 emissions by substituting biomass for fossil fbels, mew, 20, pp. 1079-1113. 4. Lewandowski I, Kicherer A,, and Vonier P. (1995) COrbalance fbr the cultivation
and combustion ofMiscanths, Biomass and Bimneqy, 8, pp. 8 1-90. 5. Borjesson P.I. I. (19%) Emission ofCo;! from biomass production and
transportation in agkulture and forestry, Energy Comers. Mgmt, 37, pp. 12351240. 6. Borjesson P.1. I. (1996) Energy analysis ofbiomass production and transportation, Biomass andBioeneqgu, 11, pp. 305-318. 7. Ford-Robertson J. B. (1996) Estimating the net carbon balance ofthe plantation fbrest industry in New zealand, Biomass and Bimneqy, 10, pp. 7-9. 8. Boman U. R. and Turnbull J. H.(1997) Integrated biomass energy systems and
962
emissions ofcarbon dioxide, Biomass and Bioemrgy, 13, pp. 333-343. 9.Dubuisson X.and Smtzoff I. (1998)Energy and C02 balances in dBzrent power generation routes using wood fuel fiom short rotation coppice, Biomass and Biaenergy, 15, pp. 379-390. 10.Ravindranath N. H.(1993) Biomass gasification: Environmentally sound technology for decentralized power generation, a case study flom India, Biomass and Bioenergy, 4, pp. 49-60. 1 1. Salo K. and Pate1 J. G. (1994)Integrated gasification combined cycle based on pressurized fluidized bed gasification, In: Advances in thennochemical biomass conversion (Ed. by A.V. Bridgwater), pp. 994-1005,BLACKIE ACADEMIC & PROFESSIONAL. 12.Babu S.P. (1995)Thermal gasification ofbiomass technology developments; end oftask report for 1992 to 1994,Biomass andBioenergy, 9,pp. 271-285. 13. Broek R. , Faaij A., and Wijk A. (1996)Biomass combustion for power generation, Biomass and Bioenergy, 11, pp. 271-281. 14.McGowin C.R., and Wiltsee G. A. (19%) Strategic analysis ofbiomass and waste fuels for eledric power generation, Biomass and Bioenergy, 10, pp. 167175. 15. Lamp P., Reichel A., and Funk R. (1997)The efficiency ofheat and power production fiom combustion and gasification, In: Developments in thennochemical biomass conversion (Ed. by A.V. Bridgwater and D.G.B.Boocock), pp. 1590-1599,BLACKIE ACADEMIC & PROFESSIONAL. 16.McIlveen-Wright D. R., Williams B. C., and McMullan J. T. (1997) Electricity generation fiom wood-fired power plants; the principal technologies reviewed, In: Developments in thennochemicalbiomass conversion (Ed. by A.V. Bridgwater and D.G.B.BoocoCk),pp. 1525-1538,BLACKIE ACADEMIC &. PROFESSIONAL. 17.Solantausta Y.,Bridgwater A. V. , and Beckman D. (1997)The performanceand economics of power fiom biomass, In: Developments in thennochemical biomass conversion (Ed. by A.V. Bridgwater and D.G.B.Boocock), pp. 1539-1555, BLACKIE ACADEMIC & PROFESSIONAL. 18.Tahara K., Kojima T., and Inaba A. (1997)Estimation ofpower plants by LCA, Kagaku Kougaku Ronbunn S’u, 23, pp. 88-94. 19.Toft A. J. and Bridgwater A. V. (1997)Opportunitiesfor $st pyrolysis in small-scale electricity generation, In: Developments in thennochemical biomass conversion (Ed. by A.V. Bridgwater and D.G.B.Boocock), pp. 1556-1566, BLACKIE ACADEMIC & PROFESSIONAL. 20.Ishitobi H.(1997)Current situation on international negotiation and other COP 3 efforts, Haikibutu Gakkaishi, 8, pp.413-420. 21.National Research Institute for Pollution and Resources (Ed) (1990), Chikyuu Onrlannka No Taisah Gijyutu, p. 93,O h u Shya. 22. Sangyou Gijyutu Kaigi (Ed) (1996)Enerugii To Kankyou, p.593,Sangyou Gijyutu Kaigi. 23.FA0 Data Base (2000)http://apps.fao.org/lim5OO/nph-
wrap.pl?LandUseBrDomain=LUI&servlet=l.
963
The Importance of Bioenergy and its Utilization Technologies Evaluated by a Global Energy and Land Use Model Hiromi Yamamoto’, Junichi Fujino2,Kenji Yamaji3 Central Research Institute of Electric Power Industry (CRIEPI), 16-1 Otemachi, Chiyoda-ku, Tokyo, 100-8126, Japan National Institute for Environmental Studies (NIES), 16-2 Onogawa, Tsukuba, Ibaraki, 305-0053, Japan The University of Tokyo, 7-3-1 Hongo, Bunkyo-ku, Tokyo, 1130033, Japan
‘ ’
ABSTRACT: In order to evaluate bioenergy resources and bioenergy utilization technologies systematically, we developed a global energy and land use model and its data. We set C02 emission constraints and conducted simulation analyses using the model. Through a set of simulations, we obtained the following results. (1) There will be supply potential of energy crops produced from surplus arable land in North America, Western Europe, Oceania, Centrally Planned Asia, Other Asia, Latin America, and Former USSR & Eastern Europe in 2050. “Ultimate supply potential” of biomass residues, which is defined as all discharged biomass minus material-recycled biomass, will be large in North America, Centrally Planned Asia, Other Asia, and Latin America where will be major consumers or exporters of biomass. (2) C02 emission constraints will be advantageous to bioenergy the net C02 emission of which is zero. The bioenergy economically used in 2050 in the world will be 50 EJ/yr in no-CO2-reduction (FREE) case and 117 EJ/yr in 3O%-CO2-reduction (C30P) case. Especially, COz emission constraints will enlarge uses of industrial roundwood harvesting residues, paper scrap, cereal harvesting residues, animal dung, human feces, and energy crops. (3) Various kinds of bioenergy utilization technologies will be used in a case with C02 emission constraint in 2050. Especially, technologies of gasification and industrial heat supply will be introduced largely under constraints of C02 emissions.
964
INTRODUCTION Bioenergy is expected to become one of major energy resources for sustainable development of mankind. However, bioenergy supply potential cannot be infinite since land area available for biomass production is limited and a certain amount of biomass must be reserved for food and material. However, bioenergy can be produced not only from bioenergy plantations, which occupy land, but also from biomass residues (such as straw, animal dung, and wood scrap) which do not occupy land directly. These biomass residues are discharged at various processes in biomass flow from harvest to consumption. In order t o evaluate the global bioenergy supply potential comprehensively, we developed a global land use and energy model (GLUE) considering land use competition and overall biomass flow including those of biomass residues. We reported the analyses using the model [1][2][3][4][5].However, we did not consider bioenergy-related costs concerning bioenergy resources and utilization technologies in those analyses. The purpose of this study is to analyze both the global bioenergy supply potential and the bioenergy-related costs a t once. For that purpose we developed a global energy and land use model. The model evaluates not only bioenergy supply potential but also hopeful bioenergy resources and utilization technologies considering cost-minimization in whole energy systems. In addition, the model shapes bioenergy use scenarios under various constraints of greenhouse gas emissions. In the following sections we explain the outline of the model, simulation results, and conclusions, respectively.
OUTLINE OF GLOBAL ENERGY AND LAND USE MODEL In this section we outline the structure of the model [6].
1. Modeling technique This model is described using Liner Programming (LP) technique in GAMS package (71. 2. Structure of the model
The model consists of two parts: an energy systems part and a land use part . The energy systems part is based on a global energy systems model named New Earth 21 (NE21)[8] and the land use part is base on a global land use and energy model (GLUE11)[6][5]. NE21 includes a detailed description of various energy utilization technologies and forecasts optimal energy systems using Non Linear Programming (NLP) technique [8]. In this study, we converted some non-linear constraints in NE21 to linear constraints in our model. Moreover, we added constraints concerning bioenergy utilization technologies. GLUE-11 evaluates global bioenergy supply potential considering land use competitions [6][5]. The objective function of the model is the summation of the energy system costs [6].
965
k
Y
P
w
E
k
Carbon monoxide
P
-s 0 r.
Sugarcane harvesting residues
elwood harv. residues Fuelwood haw. residues
Table 1: Regions - in the model. No. Regions 1 North America Western Europe Japan Oceania Centrally Planned Asia Other Asia Middle East & North Africa SubSahara Africa Latin America 10 Former USSR & Eastern Europe
The land use part covers a wide range of land uses and biomass flow including food chains, material recycling, and discharge of biomass residues (Figure 2 and Figure 3) [3][5]. 2. Structure of calculation
The structure of the land use part, which is the same as that of GLUE-11 [5], is as follows (Figure 4).
DEFINITION OF BIOENERGYS.UPPLYPOTENTIAL In this study, we define "ultimate bioenergy supply potential" for biomass residues, modern fuelwood, and energy crops. We make a numerical evaluation in the following parts of this study. In this study, we do not evaluate bioenergy potential from pasture, other land (such as desert, tundra, and residential area), and the water (the sea and fresh water) numerically. The reasons are as follows. (1)The great portion of biomass production (except feed use) in pasture area must be reserved for natural fertilizer in the own area 131. (2) The productivity of biomass on other land is small. (3) Fishery catch has already hit the ceiling 191. (4) It is considered that bioenergy production from the water (such as giant kelp) is difficult by the reason of the high costs [lo]. We define "ultimate bioenergy supply potential" of biomass residues as follows. (Ultimate bioenergy supply potential) = (all the discharged biomass residues)
. (1 - material-recycling ra-
tios)
We consider material-recycling ratios of timber scrap (including board scrap), sawmill residues, and paper scrap, which are specified in the wood biomass flow (Figure 2). "Ultimate energy potential" of energy crops is calculated by multiplying area of surplus arable land by productivity of energy crops.
967
Primary
Secondary
Intermediate
...................
Traditional helwoop
scrap
........................
Fuelwood
..........
Modem fuelwood
...........
Fuelwqod harvesting residues
(energy)
-.
(energy)
Figure 2: Wood biomass flow in the model.
Rimaw
Secondary
Intermediate
Figure 3: Food biomass flow in the model.
968
scrap
]" [Land
......."
/Wood Sector)
1
hput Data for Wood Demand * Populahon * Wood Demand Per Capita (Paper, Tunber, Trad.Fuelwod, and Mod.Fuelwood)
1
i
Fuelwood (Trad. and Mod.)
1
Woodv Biomass Residu.1-
-
( I Parameters of Forest Protection (Reforestation and Un-Sustainable Slash and Bum Farming)
'..,
""
of Determining Regional Land Uses Forest
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1
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i L
-7
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Potential
...,_
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Figure 4: Structure of the model.
969
Potential
^"
BIOMASS-RELATED DATA We set the data about biomass and land uses on the basis of the analyses in reference [5] and middle or reference projections in World Bank [ll],IPCC [12] [13], and other references [14] [15]. The details of the biomass supply and demand data were described in the reference [5]. On the other hand, we explain bioenergy-related data such as bioenergy resource costs and data about bioenergy utilization technologies in this section. The model includes only the cost of energy / bioenergy, and does not include costs of agriculture nor forestry. This is because we purpose to evaluate energy systems including bioenergy but not to evaluate economy of agriculture and forestry in this study.
OUTLINE OF BIOENERGY RESOURCE COST Kinds of biomass resources that can be used for energy can be used also for material or food. Therefore, we set bioenergy costs considering opportunity costs of biomass for material or food. If the production cost of a kind of biomass is less than the price of the biomass for material or food, the biomass cost is priced the opportunity cost for material or food. On the other hand, there are disposal costs of some kinds of biomass residues if the biomass has no value except for energy. We assume the principle to set bioenergy costs is as follows. When an opportunity cost occurs (Bioenergy resource cost) = (supply cost)
+ (opportunity cost)
When a disposal cost occurs (Bioenergy resource cost)
=
(supply cost) - (disposal cost)
In the other cases (Bioenergy resource cost) = (supply cost) (Note) The supply cost comprises costs of harvest, transport, and process of the bioenergy. As an example we explain prices of roundwood, timber, wood chips (for wood pulp), and saw dust. An average price of roundwood is about $84/m3 (about $7/GJ). Market prices of timber and wood chips for wood pulp are about 45,000 yen/m3 (about $31/GJ ) and about 8,000 yen/m3 (about $5/GJ), respectively. A market price of sawdust for fuel or manure is about 1,000 yen /m3 (about $0.7/GJ) (61. The wood prices for non-energy are higher the wood price for energy. The resource costs are assumed to increase with a stair-like function in proportion t o the resource utilization ratios (resource use per ultimate resource supply potential). We assume that bioenergy resource costs with disposal costs, such as black liquor, animal dung, and human manure, are zero. These costs may be negative actually.
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OUTLINE OF DATAOF BIOENERGY UTILIZATIONTECHNOLOGIES We consider data of bioenergy utilization technologies shown in Figure 1. The main data we set are shown in Table 2. These data are based on references I81[141[161.
Table 2: Main data of bioenergy utilization technologies. Capital cost
Power generation (steam) a) Power generation (IGCC) b) Power generation (waste) c) Liquefaction (syn-oil) a) Liquefaction (fermentation) a) Gasification d) Biogas and CHP e)
Capital cost unit
1,500 $/kW 3,000 $/kW 1,300 $/kW 150,000 $/TOEof 220,000 $/TOEof 160,000 $/TOEof 230,000 $/TOEof
bio./d bio./d bio./d bio./d
Ratio of annual expense to capital 0.17 0.17 0.17 0.25 0.25 0.25 0.25
Rate of operation
0.85 0.85 0.85 0.9 0.9 0.9 0.9
a) The data are based on [8]. b) The data are based on [lS]. We do not consider the future cost reduction of IGCC in this study. We plan to analyze the effect of the cost reduction of IGCC in the coming study. c) The data are based on [16]. We assume that Power generation (waste) is an add-on power plant to a municipal waste incineration plant. Therefore, the cost does not include incineration equipment and includes only add-on power generation equipment. d) The data are based on [B] and [14]. e) The data are based on [IS].
SIMULATION RESULTS First, we set three simulation cases where we vary COz emission constraints. Next, we conduct simulation analyses using the model and the data.
COZ EMISSION SCENARIOS The followings are the three simulation cases we set. 1. FREE
FREE is no COz constraint case. 2. CPSF
CP3F is COP3 forever case. In CPSF, the greenhouse gas constraints on the developed regions (including Former USSR) in 2010 in COP3 will continue t o be the same forever. There are no COz constraints on the developing regions. Tradable COZ permits are traded among the developed regions.
97 1
3. C30R COZ emissions in C30R in all the regions in the world will be by 30% less than those in CP3F in and after 2020. Tradable COz permits are traded among all the regions in the world.
BIOENERGY RESOURCES We explain simulation results of ultimate bioenergy supply potential and bioenergy consumption in 2050. First, there will be ultimate supply potential of energy crops produced from surplus arable land in North America, Western Europe, Oceania, Centrally Planned Asia, Other Asia, Latin America, and Former USSR & Eastern Europe in 2050 (Figure 5 ) . Next, ultimate supply potential of biomass residues will be large in North America, Centrally Planned Asia, Other Asia, and Latin America where will be major consumers or exporters of biomass. The values of the potential in those regions will be larger than 20 EJ/year in 2050, respectively (Figure 5). Next, we evaluate the amounts of bioenergy uses considering bioenergy-related costs and COz emission constraints in the whole energy systems. The bioenergy use will be large in regions where the supply potential will be large (Figure 6). The bioenergy use in the world will be 50, 77, and 119 EJ/yr in FREE, CP3F, and C30R, respectively (Table 3). The severer COz emission constraints will be, the larger bioenergy use will be. Last, Figure 7 shows various kinds of bioenergy resources in use in 2050. COz emission constraints will be advantageous to bioenergy the net COz emission of which is zero. Especially, COz emission constraints will enlarge uses of industrial roundwood harvesting residues, paper scrap, cereal harvesting residues, animal dung, human feces, and energy crops.
Table 3: Summary of bioenergy use (in 2050)(Unit: EJ/yr). Developed regions Developing regions World
FREE 19 31 50
CP3F 39 38 77
C30R 48 71 119
BIOENERGY UTILIZATION TECHNOLOGIES All the bioenergy utilization technologies except power generation (IGCC) will be used in C30R in 2050 (Figure 8). This is because the technology of power generation (IGCC) is under development and we assumed the capital cost of that, which is uncertain, is twice as many as that of the power generation (steam) (see Table 2) in this study. The uses of gasification and industrial heat supply will be sensitive to the constraints of COz emissions (Figure 8). The gasification technology is in a part of methanol production technologies; methanol will be an important energy
972
i.EnergycropsIoBiornass residues
Figure 5: Ultimate bioenergy supply potential (in 2050).
-3
25 20
\
t
15
P 10 c
6
5
0
m
m
.-0
0 ._ L
a
S
E
P r 0
S
r
Z
z"
4
m 'C 0
f?
a C .-c
J
c
I W 0 U
3
I -~
Figure 6: Bioenergy use (in regions in 2050).
973
Figure 7 Bioenergy use (in resources in 2050).
resource in C30R because COz content of methanol is smaller than fossil energy resources. In industrial heat demand, biomass, the net COZ emission of which is zero, will be competitive with coal in COZ constraint cases.
CONCLUSIONS In order to evaluate bioenergy resources and bioenergy utilization technologies systematically, we developed a global energy and land use model and its data. We set COZemission constraints and conducted simulation analyses using the model. Through a set of simulations, we obtained the following results. (1) There will be ultimate supply potential of energy crops produced from surplus arable land in North America, Western Europe, Oceania, Centrally Planned Asia, Other Asia, Latin America, and Former USSR & Eastern Europe in 2050. Ultimate supply potential of biomass residues will be large in North America, Centrally Planned Asia, Other Asia, and Latin America where will be major consumers or exporters of biomass. (2) C02 emission constraints will be advantageous to bioenergy the net C02 emission of which is zero. The bioenergy economically used in 2050 in the world will be 50 EJ/yr in no-COz-reduction (FREE) case and 117 EJ/yr in 30%-CO~-reduction(C30P) case. Especially, C02 emission constraints will enlarge uses of industrial roundwood harvesting residues, paper scrap, cereal harvesting residues, animal dung, human feces, and energy crops. (3) Various kinds of bioenergy utilization technologies will be used in a case with CO2 emission constraint in 2050. Especially, technologies of gasification and industrial heat supply will be introduced largely under constraints of C02 emissions.
974
50
45
-p
40 35
3 30
f 25
g 20
$ 15
OCP3F in 2050
U
n
I
10
5
0
Figure 8: Bioenergy use (in technologies in 2050).
REFERENCES [l] Yamamoto, H & Yamaji, K. (1997) in Developments in Therrnochemical Biomass Conversion, eds. Bridgwater, A. V & D.G.B.Boocock. (Blackie Academic & Professional, London) Vol. 2, pp. 1599-1613. [2] Yamamoto, H, Yamaji, K, & Fujino, J. (1998) International Journal of Global Energy Issues 11, 91-103. [3] Yamamoto, H, Yamaji, K, & Fujino, J. (1999) Applied Energy 63,101-113. [4] Yamamoto, H, Yamaji, K, & Fujino, J. (2000) Applied Energy 66,325-337. [5] Yamamoto, H, Fujino, J, & Yamaji, K. (1999) Estimation of regional bioenergy supply potential using a global land use and energy model, (Central Institute of Electric Power Industry (CFUEPI), Tokyo), Research Report Y98023. [6] Yamamoto, H, Fujino, J, & Yamaji, K. (1999) Bioenergy Systems Analysis with a Global Energy and Land Use Model. (United States Association for Energy Economics and International Assosiation for Energy Economics, Orlando, USA), pp. 245-254.
[7] Brooke, A et al. (1998) GAMS: a user’s guide, release 2.25. (The Scientific Press, South San Francisco). [8] Fujii, Y & Kaya, Y. (1993) Transactions of the Institute of Electrical Engineers of Japan 113-B,1213-1222.
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[9] World Resources Institute (WRI), ed. (1992) World Resources 1992-1993. (Oxford University Press).
[lo] Lee, K et al. (1991) Biomass state-of-the-art assessment -volume 1: Guide-, (Electric Power Research Institute (EPRI)), Technical Report GS-7471. [11] Bos, E et al. (1993) World Population Projections; 1992-93 Edition. (The John Hopkins University Press). [12] U S . Environmental Protection Agency (EPA). (1990) Report of the expert group on emissions scenarios of the response strategies working group of the intergovernmental panel on climate change - appendix, (EPA, Washington, D.C.), Technical report. [13] Pepper, W et al. (1992) Emission scenarios for IPCC an update, (IPCC Working Group l ) , Technical report. [14] Johansson, T et al., eds. (1993) Renewable Energy. (Island Press, Washington D.C). [15] Alcamo, J, ed. (1994) IMAGE 2.0; Integrated Modeling of Global Climate Change. (Kluwer Academic Publishers). [16] CADDET Renewable Energy. (1999) (On Internet (http://www.caddetre.org)).
976
An Overview of Fast Pyrolysis
',
AV Bridgwater S Czernik 2, J Piskorz I Bio-Energy Research Group, Aston Universip, Birmingham B4 7ET,
UK 3
NREL, I6I 7 Cole Boulevard, Golden, Colorado 80401, USA RTI Ltd, 110 Bafin Place, Unit 5, Waterloo, Ontario, N2V 127, Canada
ABSTRACT The process of fast pyrolysis is one of the most recent renewable energy processes to have been introduced and offers the advantages of a liquid product, biooil, that can be readily stored and transported and that can also be used for production of chemicals as well as being a fuel. Thermal biomass conversion has been investigated for many years as a source of renewable solid, gaseous and liquid fuels. Compared to combustion, which is widely practised commercially and gasification, which is being extensively demonstrated around the world, fast pyrolysis is at a relatively early stage of development. The technology has now achieved some commercial success for production of chemicals and is being actively developed for producing liquid fuels. Bio-oils have been successfully tested in engines, turbines and boilers, and have been upgraded to high quality hydrocarbon fuels although at an unacceptable energetic and financial cost. The paper critically reviews scientific and technical developments and applications to date paying particular attention to the research and developments reported in this book. It concludes with some suggestions for strategic developments. INTRODUCTION Renewable energy is of growing importance in satisfying environmental concerns over fossil fuel usage. Wood and other forms of biomass are some of the main renewable energy resources available and provide the only source of renewable liquid, gaseous and solid fuels. Wood and biomass can be used in a variety of ways to provide energy: 0
by direct combustion to provide heat. This technology is commercially available and presents minimum risk to investors. The product is heat, which must be used immediately for heat and/or power generation. Overall efficiencies to power tend to be rather low, although in many California power plants h g h efficiencies have been reported but there have been problems with ash build-up in many installations. The status is reviewed in these proceedings (1). by gasification to provide a fuel gas for combustion for heat, or in an engine or turbine for electricity generation. The fuel gas quality requirements, for turbines in
977
particular, are very high, although there is now extensive experience available from the Varnamo plant (2). The gas is very costly to store or transport SO it has to be used immediately. Hot gas efficiencies (total energy in raw product gas divided by energy in feed) can be as high as 9597% for close coupled turbine and boiler applications, and up to 85% for cold gas efficiencies. The status is reviewed in (3). There is renewed interest in synthesis of liquid hels from the product gas, but costs are still very high. by fast pyrolysis to provide a liquid fuel that can substitute for fuel oil in any static heating or electricity generation application. The liquid can also be used to produce a range of speciality and commodity chemicals. The key advantage is that a liquid can be readily stored and/or transported. This review explains how the technology is developing and how it depends on an improved understanding of the underlying science as reviewed four years ago (4). FAST PYROLYSIS Pyrolysis is by definition thermal decomposition occurring in the absence of oxygen. It is always also the first step in combustion and gasification processes where it is followed by total or partial oxidation of the primary products. Fast pyrolysis occurs in time of few seconds or less. Therefore, not only chemical reaction kinetics but also heat and mass transfer processes, as well as phase transition phenomena, play important roles. The critical issue is to bring the reacting biomass particle to the optimum process temperature and minimise its exposure to the intermediate (lower) temperatures that favour formation of charcoal. This objective can be achieved by using small particles, thus reducing the time necessary for heat up. This option is used in fluidised bed processes that are described later. Another possibility is to transfer heat very fast only to the particle surface that contacts the heat source. Because of the low thermal conductivity the deeper parts of the particles will be maintained at temperatures lower than necessary for char production. The products that form on the surface are immediately removed exposing that way consecutive biomass layers to the contact with the heat source. This second method is applied in ablative processes that are described later. Fast pyrolysis is not an equilibrium process. During fast pyrolysis dramatic changes occur in specific volume between the reactants (biopolymers) and the products (by a factor of x 500) causing the volatile products to leave the pyrolysis zone at considerable velocities. This results in the entrainment of solid particles and aerosols, which normally would not volatilise at the process temperature. All these phenomena have important implications on pyrolysis technologies and are discussed in this paper. MECHANISMS
Biomass is a complex polymeric material and its thermal decomposition is a multistage complicated process. Many pathways and mechanisms have been proposed to illustrate/explain the fundamental steps in pyrolysis (e.g. 5 , 6, 7, 8, 9, 10, 1I , 12, 13). Broido-Shafizadeh type kinetic models are perhaps the most widely used for cellulose pyrolysis but they can be also applied, at least qualitatively, to the whole biomass (see Figure 1). Correct estimation of the kinetic constants is essential in optimising process parameters for maximizing liquids production. Nevertheless, kmetic data does not contribute any predictive powers aiming to optimise yields of specific chemicals or 978
their classes. As fast pyrolysis is a non-steady state process, the derivation of isothermal kinetic data for modelling is extremely difficult, except in very small samples, as mass transfer processes, phase transition and heat transfer phenomena play important roles. The so-called “black box” engineering approach dominates our understanding in this area at present.
11Water, char, CO;! (dehydration, decarboxylation, carbonisation)
I11 Volatiles (depolymerisationand scission products, prompt gas)
IV Secondary tar, char, gas Figure 1 Typical cellulose decomposition model As shown in the model, pyrolysis of cellulose results in solid, liquid and gaseous products. However, the proportions of the product yields can change depending on the process conditions. The knowledge of thermodynamics and kinetics of the reaction pathways allows us to adjust the conditions to maximize the yield of the desired products. Dehydration of cellulose is exothermic while depolymerisation and secondary vapour cracking are endothermic and have higher activation energy than dehydration. Therefore, lower process temperature and longer vapour residence times will favour the production of charcoal. High temperature and longer residence time will increase the biomass conversion to gas and moderate temperature and short vapour residence time, necessary to minimize secondary cracking, are optimum for producing liquids. Table 1 provides data on the product distribution obtained from different modes of pyrolysis process.
Table 1 Typical product yields obtained by different modes of pyrolysis of wood Liquid Char FAST PYROLYSIS 75% 12% moderate temperature and short residence time, particularly vapour CARBONISATION 30% 35% low temperature and long residence time 5% 10% GASIFICATION high temperature and long residence times
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Gas 13%
35% 85%
PRINCIPLES Fast pyrolysis is a high temperature process in which biomass is rapidly heated in the absence of oxygen. As a result it decomposes to generate mostly vapours and aerosols and some charcoal. After cooling and condensation, a dark brown mobile liquid is formed which has a heating value about half that of conventional fuel oil. While it is related to the traditional pyrolysis processes for making charcoal, fast pyrolysis is an advanced process, with carefully controlled parameters to give high yields of liquid. The essential features of a fast pyrolysis process for producing liquids are: 0 very high heating and heat transfer rates at the reaction interface, whch usually requires a finely ground biomass feed, carefully controlled pyrolysis reaction temperature of around 500°C and vapour phase temperature of 4O0-45O0C, 0 short vapour residence times of typically less than 2 seconds, 0 rapid cooling of the pyrolysis vapours to give the bio-oil product. The main product, bio-oil, is obtained in yields of up to 75% wt on dry feed basis, together with by-product char and gas which are used within the process so there are no waste streams other than flue gas and ash. Figure 2 shows a conceptual schematic of fast pyrolysis process that includes the necessary steps of drying the feed to typically less than 10% water to minimise the water in the product liquid oil (although up to 15% can be acceptable), grinding the feed (to around 2 mm in the case of fluid bed reactors) to give sufficiently small particles to ensure rapid reaction, pyrolysis reaction, separation of solids (char), and collection of liquid product (bio-oil).
BIOMASS
*
Dry
)BIO-OIL
Grind
)Char
Figure 2. Conceptual Fast Pyrolysis Process
REACTORS The heart of a fast pyrolysis process is the reactor. Although it probably represents at most only about 20% of the total capital cost, almost all research and development has focussed on the reactor. The rest of the process consists of biomass reception, storage and handling, biomass drymg and comminution, and product collection, storage and, when relevant, upgrading. The key aspects of these peripheral steps are described later. The critical features of successful pyrolysis reactors have been defined above as very h g h heating rates, carefully controlled temperatures and rapid cooling or quenching of the gaseous product which includes vapours, aerosols and gases. A comprehensive survey of fast pyrolysis processes has recently been published that lists and describes all the pyrolysis processes for liquids production that have been built and tested in the last 10-15 years (14). Table 2 summarises the current significant operational fast pyrolysis plants for production of liquids.
980
Table 2 Operational pyrolysis units for liquids (see (14) for further details). Throughputs are on a dry wood basis. Fluid bed
Transported bed CFB Rotating cone Ablative Vacuum
250 kgih at Wellman 75 kgih at Dynamotive 20 kgih at RTI Many research units including in these proceedings (1 5) about 1500 kgh at Red Arrow (Ensyn) 650 kg/h at ENEL (Ensyn) 20 kg/h at VTT (Ensyn) 10 kgih at CRES (22) 150 kgh at BTG (28) 20 k g h at NREL 20 kgih at Aston 3500 kgih at Pyrovac (29)
Bubbling fluid beds Bubbling fluid beds have been selected for further development by several companies including Union Fenosa (1 6) who have built and operated a 200 kgih pilot unit in Spain, Dynamotive who have a 75 kgh unit in Canada based on a RTI design (17) with a 10 t/d plant under construction (18) and Wellman who have built a 250 kgih unit in the UK. Many research units have been built as they are relatively easy to construct and operate and give good results (14, and in these proceedings). Bubbling fluid beds have many attractive features, particularly for research and development, which include: Simple construction and operation, Good temperature control, Very efficient heat transfer to biomass particles due to high solids density, Easy scaling, Well understood technology, Good and consistent performance with h g h liquid yields: of typically 70-75%wt. from wood on a dry feed basis, Particular features that require consideration in design and operation include: Heating can be achieved in a variety of ways as shown in Figure 3, Residence time of solids and vapours is controlled by the fluidising gas flow rate and is higher for char than for vapours, Char acts as an effective vapour cracking catalyst at fast pyrolysis reaction temperatures so rapid and effective char separatiodelutriation is important, Small biomass particle sizes are needed to achieve hgh biomass heating rates, and particle sizes generally need to be less than 2 to 3 mm, Good char separation is important and this is usually acheved by ejection and entrainment followed by separation in one or more cyclones, Heat transfer to bed at large scale has to be considered carefully due to scale-up limitations (see also Figure 3).
981
A typical plant configuration is shown in Figure 4,in this case based on a bubbling fluid bed, but similar principles apply to most reactors.
Figure 3. Means of providing heat to a pyrolysis reactor
Figure 4. Bubbling fluid bed reactor
Circulating fluid beds and transported bed Circulating fluid beds and transported bed reactors have been developed to commercial status by Ensyn and their process is used commercially by Red Arrow in the USA for food flavourings in several plants of 1 to 1.5 t/h (19). Ensyn have also supplied a 650 kgh unit to ENEL in Italy (20) and a 20 k g h system to VTT in Finland (21). In some 982
of these plants it is believed that only sand is recycled while char is separated and collected, in others, operating in a dual reactor configuration, char is also recycled and burned in the second reactor to provide process energy by reheating and recycling the sand. CRES is operating a 10 k g h circulating fluid bed unit with the char combustor integrated into the base of the CFB riser as a bubbling fluid bed (22) as indicated in Figure 5. Liquid yields of 60-70% wt on a dry feed basis have been achieved. The principles of a circulating fluid bed system are shown in Figure 5. As discussed above, a second rector can be required for char combustion and sand reheating (Ensyn) or the char combustion can occur in the lower part of the pyrolyser (CRES) thus eliminating the need for the second vessel. Particular features of circulating fluid bed and transported bed reactors include: 0 0 0 0 0
0 0
Good temperature control can be achieved in reactor, Residence time for the char is almost the same as for vapours and gas, CFBs are suitable for very large throughputs, Well understood technology, Hydrodynamics more complex, Char is more attrited due to higher gas velocities; char separation is by cyclone, Closely integrated char combustion in a second reactor requires careful control, Heat transfer at large scale has to be proven.
Figure 5. Circulating fluid bed reactor
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Ablative pyrolysis Ablative pyrolysis is substantially different in concept compared to the other metkods of fast pyrolysis. In all these other methods, the rate of reaction is limited by the rate of heat transfer through a biomass particle, which is why small particles are required. The mode of reaction in ablative pyrolysis is analogous to melting butter in a frylng pan, when the rate of melting can be significantly enhanced by pressing down and moving the butter over the heated pan surface. In ablative pyrolysis heat is transferred fi-om the hot reactor wall to “melt” wood that is in contact with it under pressure. The pyrolysis front thus moves unidirectionally through the biomass particle. As the wood is mechanically moved away, the residual oil film both provides lubrication for successive biomass particles and also rapidly evaporates to give pyrolysis vapours for collection in the same way as other processes. The rate of reaction is strongly influenced by pressure, the relative velocity of wood on the heat exchange surface and the reactor surface temperature. The key features of ablative pyrolysis are therefore as follows: 0
0
a
High pressure of particle on hot reactor wall, achieved due to centrifugal force (NREL) or mechanically (Aston) High relative motion between particle and reactor wall, Reactor wall temperature less than 600°C. Particular features of ablative reactor systems include:
0 0
0
0
0 0 0
Use of large feed sizes, Inert gas is not required, so the processing equipment is smaller, (in case of mechanically applied pressure) The reaction system is more intensive, The process is limited by the rate of heat supply to the reactor rather than the rate of heat absorption by the pyrolysing biomass as in other reactors. Reaction rates are limited by heat transfer to the reactor, not to the biomass, The process is surface area controlled so scaling is more costly, The process is mechanically driven so the reactor is more complex.
Ablative pyrolysis is interesting as much larger particle sizes can be employed than in other systems and there is no requirement for inert gas. Both lead to a potentially lower cost system. Much of the pioneering work on ablative pyrolysis reactors was performed by CNRS at Nancy where extensive basic research has been carried out onto the relationships between pressure, motion and temperature (23). NREL developed an ablative vortex reactor (see Figure 6 and (14)),in which the biomass is accelerated to supersonic velocities to derive high tangential pressures inside a heated cylinder. Unreacted particles are recycled and the vapours and char fines leave the reactor axially for collection. Liquid yields of 60-65%wt. on dry feed basis are typically obtained. More recent developments have been carried out at Aston University with a prototype rotating blade reactor in which pressure and motion is derived mechanically thus obviating the need for a carrier gas (see Figure 7 and (14)). Liquid yields of 7075%wt. on dry feed are typically obtained. A second-generation reactor has recently been built and commissioned. Other configurations include the coiled tube at Castle Capital (14)(now owned by Enervision) and cyclonic reactors (24).
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Figure 6. NREL Vortex ablative reactor
Figure 7. Aston University rotating blade ablative reactor Entrained flow Entrained flow fast pyrolysis was developed at Georgia Tech Research Institute (35) and scaled up by Egemin (26). Further details are available in (14). Neither the GTRI nor Egemin process is now operational and there are no known plans for further development, probably because of the difficulties that have been encountered in acheving good heat transfer from a gaseous heat carrier to solid biomass. Liquid yields of 50-60%wt. on dry feed have been obtained. Rotating cone The rotating cone reactor, invented at the University of Twente (27) and being developed by BTG, is a recent development and effectively operates as a transported 985
bed reactor, but with transport effected by centrifugal forces rather than gas (28). A 250 kg/h unit is now operational, and plans for scale-up to 10 t/d have recently been announced. The basic principle is shown in Figure 8 and further details are available in (14) and (28). The key features of this technology are: 0
0
0
0
0
Centrifugation at 600 rpm drives hot sand and biomass up a rotating heated cone, Vapours are collected and processed conventionally, Char is burned in a secondary bubbling fluid bed combustor. The hot sand is recirculated to the pyrolyser, Carrier gas requirements in the pyrolysis reactor are much less than for fluid bed and transported bed systems (however, more gases are needed for char burn off and for sand transport) Complex integrated operation of three subsystems: rotating cone pyrolyser, bubbling bed char combustor, and riser for sand recycling. Liquid yields of 60-70% on dry feed are typically obtained. Particle trajectory
Figure 8. Principle of rotating cone pyrolysis reactor Vacuum pyrolysis
Vacuum pyrolysis is arguably not a true fast pyrolysis as the heat transfer rate to and through the solid biomass is much slower than in the previously described reactors, but the vapour residence time is comparable. The basic technology was developed at the University of Lava1 using a multiple hearth h a c e but is now based on a purpose designed horizontal moving bed. A 50 kg/h unit is available for research and the technology has been scaled up by Pyrovac to a 3.5 th unit, which is currently operating at Jonquiere in Canada (29). The technology was selected for some of the Non Fossil Fuel Obligation contracts in the UK where six plants totalling about 70MWe are at various stages of permitting and design (30). Key features of the process include: 0
It can process larger particles than most fast pyrolysis reactors, There is less char in the liquid product due to lower gas velocities, There is no requirement for a carrier gas, Liquid yields of 3550% on dry feed are typically obtained with higher char yields than fast pyrolysis systems. Conversely, the liquid yields are higher than in slow pyrolysis technologies because of fast removal of vapours from the reaction zone, The process is relatively complicated mechanically,
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CHAR REMOVAL Char acts as a vapour cracking catalyst so rapid and effective separation from the pyrolysis product vapours is essential. Cyclones are the usual method of char removal and two are usually provided - the first to remove the bulk of the material and the second to remove as much of the residual fines as possible. However, some fines always pass through the cyclones and collect in the liquid product where they accelerate aging and exacerbate the instabilityproblem, which is described below. There are several solutions to the problem of char entrainment. NREL has shown that hot vapour filtration, analogous to hot gas filtration in gasification processes, gives a high quality char free product (3 I), however the liquid yield is reduced by about 1020% due to the char accumulating on the filter surface that cracks the vapours. VTT has also been developing hot vapour filtration (32). Pressure filtration of the liquid is very difficult due to the complex interaction of the char and pyrolytic lignin, which appears to form a gel-like phase that rapidly blocks the filter. Modification of the liquid micro-structure by addition of solvents such as methanol or ethanol that solubilise the less soluble constituents will improve thls problem and also contribute to improvements in liquid stability as described below.
LIQUID COLLECTION The product of fast pyrolysis is vapours, aerosols and gases from decomposition of holocellulose and lignin with any carrier gases from fluidisation or transport. Aerosols consist of sub-micron liquid droplets and they present a severe problem in the successful recovery of the pyrolysis oils. These aerosols appear visually as smoke. The aerosols are probably formed directly from pyrolysing biomass, especially from submicron biomass particles that are rapidly depolymerised. The liquid product can then be entrained out of the reactor before it is vaporised. Another mechanism proposed for the formation of aerosols in the pyrolysis reactor involves the ejection of liquid droplets from internally pressurised cell capillaries of a pyrolysing particle (33). The vapours require rapid cooling to minimise secondary reactions and to condense the true vapours, while the aerosols require coalescence or agglomeration. Simple heat exchange can cause preferential deposition of lignin derived components leading to liquid fractionation and eventually blockage. Quenching in product oil or in an immiscible hydrocarbon solvent is widely practiced. Orthodox aerosol capture devices such as demisters and other commonly used impingement devices are not very effective and electrostatic precipitation is currently the preferred method. The vapour product from fluid bed and transported bed reactors has a low partial pressure of collectible products due to the large volumes of fluidising gas, and this is an important design consideration in liquid collection. The exit gas temperature from the liquid condensation train has a definite effect on the properties of the recovered oil product. As the exit gas temperature is lowered, progressively more of the volatile organic vapours and water will be condensed to become the pyrolysis oil product. These volatiles can be considered solvents as they have very low viscosities and thus have a very beneficial mpact on lowering the viscosity of the oil, even at low concentrations. The recovery of these volatile, oxygenated organic solvents is thought to decrease the tendency of the oil to phase separate. Use of hgher collection temperatures to minimise water condensation will result in lower yields and in higher viscosity of less stable bio-oil.
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PYROLYSIS LIQUID - BIO-OIL Pyrolysis liquid is referred to by many names including pyrolysis oil, bio-oil, biocrude-oil, bio-fuel-oil, wood liquids, wood oil, liquid smoke, wood distillates, pyroligneous tar, pyroligneous acid, and liquid wood. The crude pyrolysis liquid is dark b r o w and approximates to biomass in elemental composition. It is composed of a very complex mixture of oxygenated hydrocarbons with an appreciable proportion of water from both the original moisture and reaction product. Solid char and dissolved alkali metals from ash (34) may also be present. The complexity arises from the degradation of lignin, cellulose, hemicellulose and any other organics in the feed material, giving a broad spectrum of phenolic and many other classes of compounds that result from uncontrolled degradation as described below. The liquid from fast or flash pyrolysis has significantly different physical and chemical properties compared to the liquid from slow pyrolysis processes, which is more like a tar.
LIQUID PRODUCT CHARACTERISTICS Composition The liquid is formed by rapidly quenching and thus “freezing” the intermediate products of flash degradation of hemicellulose, cellulose and lignin as shown in Table 3 and Table 4. The liquid thus contains many reactive species, whch contribute to its unusual attributes. Bio-oil can be considered a micro-emulsion in which the continuous phase is an aqueous solution of holocellulose decomposition products, that stabilizes the discontinuous phase of pyrolytic lignin macro-molecules through mechanisms such as hydrogen bonding. Aging or instability is believed to result from a breakdown in thls emulsion. In some ways it is analogous to asphaltenes found in petroleum. Table 3 Degradation products of fast pyrolysis of biomass constituents (12) Hemicellulose produces acetic acid, furfural, furan Cellulose produces levoglucosan, 5-hydrox ymethylfurfal, hydroxyacetaldehyde, acetol, formaldehyde Lignin produces small amount of monomeric phenols (including phenols, cresols, guaiacols, syringols) but mostly oligomeric product of molecular mass ranging from few hundred to several thousand Da. Extractives produces molecules of waxy components such as fatty acids and rosin acids that are difficult to identify and are mostly immiscible with biaoil. They are relatively thermally stable and volatile. The content depends on the biomass feed and can reach 5-15wt% levels Compounds classes in bio-oil c1 formic acid, methanol, formaldehyde C2-C4 linear hydroxyl and 0x0 substituted aldehydes and ketones C5-C6 hydroxyl, hydroxymethyl and/or 0x0 substituted fiuans, furanones and pyranones Anhydrosugars e.g. levoglucosan, anhydro-oligosaccharides Substituted phenols Monomeric and dimeric methoxyl substituted phenols Pyrolytic lignin Waxes, resins, fatty acids, terpenoid derivatives
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Table 4 Representative chemical composition of fast pyrolysis liquid Maior Components
Mass %
Water Limin fragments: insoluble pyrolytic lignin Aldehvdes: formaldehyde, acetaldehyde, hydroxyacetaldehyde, glyoxal, methylglyoxal Carboxylic acids: formic, acetic, propionic, butyric, pentanoic, hexanoic, glycolic, (hydroxy acetic)) Carbohvdrates: cellobiosan, a- D- levoglucosan, oligosaccharides, 1.6 anhydroglucofkanose Phenols: phenol, cresols, guaiacols, syringols Furfurals Alcohols: methanol, ethanol Ketones acetol (l-hydroxy-2-propanone),cyclo pentanone
20-30 15-30 10-20 10-15 5-10 2-5 1-4 2-5 1-5
Fast pyrolysis liquid has a higher heating value of about 17 MJkg as produced with about 25% wt. water that cannot readily be separated. The liquid is often referred to as "oil" or "bio-oil" or "bio-crude" although it will not mix with any hydrocarbon liquids. It is composed of a complex mixture of oxygenated compounds that provide both the potential and challenge for utilisation. There are some important characteristics of this liquid that are summarised in Table 5 and discussed briefly below. Table 5 Typical properties and characteristics of wood derived crude bio-oil Phvsical property Moisture content PH Specific gravity Elemental analysis
Twical value 15-30% 2.5 1.20 5558% 5.5-7.0% 35-40% 0-0.2% 0-0.2% 16-19MJkg 40-100 CP 1% up to 50%
C H
0
N
Ash HHV as produced (depends on moisture) Viscosity (at 40°C and 25% water) Solids (char) Vacuum distillation residue Characteristics 0 Liquid fuel, 0 Ready substitution for conventional hels in many static applications such as boilers, engines, turbines, 0 Heating value of 17 MJkg at 25% wt. water, is about 40% that of fuel oil / diesel Does not mix with hydrocarbon fuels, 0 Not as stable as fossil fuels, 0 Quality needs definition for each application,
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Appearance Pyrolysis oil typically is a dark brown free flowing liquid. Depending upon the initial feedstock and the mode of fast pyrolysis, the colour can be almost black through dark red-brown to dark green, being influenced by the presence of micro-carbon in the liquid and by the chemical composition. Hot vapour filtration gives a more translucent red-brown appearance due to the absence of char. High nitrogen contents in the liquid can give it a dark green tinge. Odour The liquid has a distinctive odour - an acrid smoky smell, which can irritate the eyes if exposed for a prolonged period to the liquids. The cause of this smell is due to the low2 molecular weight aldehydes and acids. The liquid contains several hundred different chemicals in widely varying proportions, ranging from formaldehyde and acetic acid to complex high molecular weight phenols, anhydrosugars and other oligosaccharides. Miscibility The liquid contains varying quantities of water which forms a stable single phase mixture, ranging from about 15 wt% to an upper limit of about 30-50wt% water, depending on how it was produced and subsequently collected. Pyrolysis liquids can tolerate the addition of some water, but there is a limit to the amount of water, which can be added to the liquid before phase separation occurs, in other words the liquid cannot be dissolved in water. It is miscible with polar solvents such as methanol, acetone, etc. but totally immiscible with petroleum-derived fuels. Density The density of the liquid is very high at around 1.2 kg/litre compared to light fuel oil at around 0.85 kg/litre. This means that the liquid has about 42% of the energy content of fuel oil on a weight basis, but 61% on a volumetric basis, This has implications on the design and specification of equipment such as pumps. Viscosity The viscosity of the bio-oil as produced can vary from as low as 25 cSt to as high as 1000 cSt (measured at 4OoC ) or more depending on the feedstock, the water content of the oil, the amount of light ends that have been collected and the extent to which the oil has aged. Viscosity is important in many fuel applications (35). Distillation Pyrolysis liquids cannot be completely vaporised once they have been recovered from the vapour phase. If the liquid is heated to 100°C or more to try to remove water or distil off lighter fractions, it rapidly reacts and eventually produces a solid residue of around 50wt% of the original liquid and some distillate containing volatile organic compounds and water. The liquid is, therefore, chemically unstable, and the instability increases with heating, so it is preferable to store the liquid at room temperature. These
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changes do also occur at room temperature, but much more slowly and can be accommodated in a commercial application. Unusual Properties The complexity and nature of bio-oil causes some unusual behavior, specifically that the following properties tend to change with time: 0 Viscosity increases, 0 Volatility decreases, Phase separation and deposition of gums can occur. This is due to a complex interaction of physical and chemical processes such as: 0 Polymerizatiodcondensation 0 Esterification and etherification 0 Agglomeration of oligomeric molecules
UPGRADING OF PYROL YSIS LIQUID The properties that negatively affect bio-oil fuel quality are foremost low heating value, incompatibility with conventional fuels, solids content, hlgh viscosity, and chemical instability. The heating value can be significantly increased, but it requires extensive changes to the chemical structure of bio-oils, which is technically feasible but not economic. The other undesired characteristics can be improved using simpler, physical methods. Both options are reviewed below. Physical Methods Hot-gas filtration can reduce the ash content of the oil to less than 0.01% and the alkali content to less than 10 ppm - much lower than reported for biomass oils produced in systems using only cyclones. Diesel engine tests performed on crude and on hotfiltered oil showed a substantial increase in burning rate and a lower ignition delay for the latter, due to the lower average molecular weight for the filtered oil (36). Hot gas filtration has not yet been demonstrated over a long-term process operation. Pyrolysis oils are not miscible with hydrocarbon fuels but with the aid of surfactants they can be emulsified with diesel oil. A process for producing stable micro-emulsions with 5-30% of bio-oil in diesel has been developed at CANMET (37). The resultant emulsions showed promising ignition characteristics. A drawback of this approach is the cost of surfactants and the high energy required for emulsification. Other work on bio-emulsions is being carried out by the University of Florence, Italy, in an EC sponsored project, using proportions of bio-oil ranging from 5% to 75% (38). Polar solvents have been used for many years to homogenize and to reduce viscosity of biomass oils. The addition of solvents, especially methanol, also showed a significant effect on the oil stability. It was observed (39) that the rate of viscosity increase (“aging”) for the oil with 10 wt. % of methanol was almost 20 times less than for the oil without additives. Chemical Methods The chemical reactions that can occur between the bio-oil and methanol or ethanol are esterification and acetalization. Although not favoured thermodynamically, they can proceed to a significant extent if appropriate conditions are applied. For example, in
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the presence of an acid catalyst and molecular sieves (to adsorb water and to shft the reaction equilibria), bio-oil reacted with ethanol forming ethyl acetate, ethyl formate, and diethoxyacetal of hydroxyacetaldehyde at the expense of formic acid, acetic acid, and hydroxyacetaldehyde (40). Eventually, in addition to the decrease in viscosity and in the aging rate, other desirable changes such as reduced acidity, improved volatility, heating value, and miscibility with diesel fuels were also achieved. T h s field has been thoroughly reviewed (4 1). The chemicallcatalytic upgrading processes aim at the removal of oxygen, which is the main cause of instability and other unwanted characteristics of bio-oils. They are more complex and expensive than physical methods, but offer sigmficant improvements ranging from simple stabilization to hgh-quality fuel products (42). Full deoxygenation to high-grade products such as transportation can be accomplished by two main routes: hydrotreating and catalytic vapour cracking. Chemical methods for upgrading bio-oil by hydro-treating and zeolite craclung have been reviewed (43,44). Hydrotreating of bio-oil carried out at high temperature, high hydrogen pressure, and in the presence of catalysts results in elimination of oxygen as water and in hydrogenation-hydrocracking of large molecules. The catalysts (sulphided CoMo or NiMo supported on alumina) and the process conditions are similar to those used in the refining of petroleum cuts (45). Catalytic vapour cracking makes deoxygenation possible through simultaneous dehydration-decarboxylation over acidic zeolite catalysts. At 450°C and the atmospheric pressure oxygen is rejected as H20, COz, and CO producing mostly aromatics (46). The low H/C ratio in the bio-oils imposes a relatively low limit on the hydrocarbon yield and, in addition, the technical feasibility is not yet completely proven. The catalyst deactivation still raises many concerns for both routes. The processing costs are high and the products are not competitive with fossil fuels (47).
APPLICATIONS FOR BIO-OIL Bio-oil can substitute for fuel oil or diesel in many static applications including boilers, furnaces, engines and turbines for electricity generation. The possibilities are summarised in Figure 8. There is also a range of chemicals that can be extracted or derived including food flavourings, specialities, resins, agri-chemicals, fertilisers, and emissions control agents. Upgrading bio-oil to transportation fuels is feasible but currently not economic. Table 6 Demonstrated applications (with references to papers in these proceedings. See also Table 7) Electricity Heat Transport fuels Bulk chemicals Fine chemicals Emissions control
diesel, turbine, Stirling (48); CHP and boiler (49); upgrading, emulsions (38); e.g. resins (50,5 l), fertilizers, hydrogen (52) e.g. levoglucosenone (53) Calcium enriched bio-oil(54)
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I
PREPARED BIOMASSDried and controlled particle size
Stabilise or
upgrade I
ELECTRICITY and/or HEAT
I
Figure 8. Applications for Bio-oil.
ELECTRICITY PRODUCTION At least 500 hours operation has been acheved in the last few years on various engines from laboratory test units to 1.4 MWe modified dual fuel diesel engines. One such engine is a 250 kWe dual fuel engine on which nearly 400 hours have been logged in total, including several runs of over 9 hours, and with electricity being generated for 320 hours ( 5 5 ) . A 2.5 MWe gas turbine has been modified and successfully run on bio-oil(56), and work is being carried out with bio-oil in gas turbine combustors (57).
CHEMICALS A range of chemicals can also be produced from specialities such as levoglucosan to commodities such as resins and fertilisers as sumtnarised in Table 7. Food flavourings are commercially produced from wood pyrolysis products in many countries. All chemicals are attractive possibilities due to their much higher added value compared to fuels and energy products, and lead to the possibility of a bio-refinery concept in which the optimum combinations of fuels and chemicals are produced. Table 7 Chemicals from fast pyrolysis (with examples of references) Acetic acid (5 8) Food flavourings (19) Levoglucosan (60) Resins (50,5 1)
Adhesives (50,51) Hydrogen (52) Levoglucosenone (53) Slow release fertilisers (62)
Calcium enriched bio-oil(54) Hydroxyaceladehyde (59) Preservatives (6 1) Sugars (7)
SUMMARY AND CONCLUSIONS The liquid bio-oil product from fast pyrolysis has the considerable advantage of being storable and transportable as well as the potential to supply a number of valuable chemicals, but there are many challenges facing fast pyrolysis that relate to technology, product and applications. The problems facing the sector include the following:
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Cost of bio-oil, whch is 10% to 100% more than fossil hel, Availability: there are limited supplies for testing, There is a lack of standards for use and distribution of bio-oil and inconsistent quality lnhibits wider usage. Considerable work is required to characterise and standardise these liquids and develop a wider range of energy applications. Bio-oil is incompatible with conventional fuels, Users are unfamiliar with this material, Dedicated fuel handling systems are needed, Pyrolysis as a technology does not enjoy a good image. Therefore, more research is needed in the field of fast pyrolysis. The most important issues that need to be addressed seem to be: Scale-up, Cost reduction, Improving product quality including setting norms and standards for producers and users, Environment health and safety issues iq handling, transport and usage, Encouragement for developers to implement processes; and users to implement applications. Information dissemination. There is much potential for fiuther development and optimisation. Chemicals offer more interesting commercial opportunities and are likely to be a major focus of continuing research and development effort.
REFERENCES Baxter, L, “Biomass combustion: Status quo and challenges”, these proceedings. Stahl K, Neergaard M, Nieminen J, “Final report: Varnamo demonstration programme”, these proceedings. Maniatis K, “Progress in biomass gasification: An overview”, these proceedings. Diebold JP and Bridgwater AV, “Overview of Fast Pyrolysis of Biomass for the production of Liquid Fuels”, pp 5-26, Developments in Thermochemical Biomass Conversion, Bridgwater, AV and Boocock, DGB (Eds.) (Blackie Academic & Professional, London 1997) Bradbury AGW, Sakai Y, Shafuadeh F, “Kinetic Model for Pyrolysis of Cellulose”, J. Appl. Polym. Sci., Vol. 23, pp.3271, 1979. Boroson ML, Howard JB,Longwell JP, Peters WA, “Heterogeneous Cracking of Wood Pyrolysis Tars over Fresh Wood Char Surfaces”, Energy & Fuels, 3, 735740, 1989. Radlein D, Piskorz J, Scott DS, “ Fast Pyrolysis of Natural Polysaccharides as a Potential Industrial Process”, J. Anal. Appl. Pyrol., 19,41-63, 1991. Vladars-Usas A, “Thermal Decomposition of Cellulose”, M.A.Sc. Thesis, University of Waterloo, Ont. Canada, 1993. Diebold J, “A Unified Global Model for the Pyrolysis of Cellulose”, Biomass and Bioenergy Vol. 7, Nos. 1-6, pp.75-85, 1994.
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Piskorz J, Scott D, Radlein D, “Mechanisms of the Fast Pyrolysis of Biomass: Comments on Some Sources of Confusion”, presented at Frontiers of Pyrolysis: Biomass Conversion and Polymer Recycling Conference, Breckenridge, 1995, included in written form in Minutes of the 2”*PYR4 Meeting, La Coruna, 1995. Narayan R, Antal MJ Jr, ‘‘Thermal Lag, Fusion and the Compensation Effect during Biomass Pyrolysis”, Ind. Eng. Chem. Res. 35, pp.1711-1721. 1996. Radlein D, “The Production of Chemicals from Fast Pyrolysis Bio-oils”, Fast Pyrolysis of Biomass: A Handbook, A. Bridgwater et al. pp164, CPL 1999. Boutin 0, Lede J, “Use of a Concentrated Radiation for the Determination of the Elementary Mechanisms of Cellulose Thermal Decomposition”, these proceedings. Bridgwater AV and Peacocke GVC, “Fast pyrolysis processes for biomass“, Sustainable and Renewable Energy Reviews, 4 (1) 1-73 (Elsevier, 1999) Gerdes C, Meier D, Kaminsky W, “Flash pyrolysis of industrial biomass wastes”, Meier D, Ollesch T, Faix 0, “Fast pyrolysis of impregnated waste wood - The fate of hazardous components”, Hata T, Meier D, Kajimoto T, Kikuchi H, Imamura Y, “Fate of arsenic after fast pyrolysis of chromium-copper-arsenate(CCA) treated wood”, All, these proceedings Cuevas A, Reinoso C and Scott DS, “Pyrolysis oil production and its perspectives”, in Proc. Power production from biomass 11, Espoo, March 1995 (VTT) PyNe Newsletter 4, September 1997, Aston University, UK Dynamotive Press release, 3 October 2000. Underwood G, “Commercialisation of fast pyrolysis products”, in ‘Biomass thermal processing’, Eds. Hogan E, Robert, J, Grassi G and Bridgwater AV, pp 226-228, (CPL Scientific Press, 1992) Rossi C and Graham RG. “Fast pyrolysis at ENEL”. In: Kaltschmitt MK and Bridgwater AV, editors, Biomass gasification and pyrolysis: State of the art and future prospects, CPL Press, 1997, p 300-306. PyNE Newsletter No. 4, September 1997, Aston University, UK Boukis I, Gyftopoulou ME, Papamichael I, “Biomass fast pyrolysis in an airblown circulating fluidized bed reactor”, these proceedings LCde J, Panagopoulos J, Li HZ and Villermaux J, “Fast Pyrolysis of Wood: direct measurement and study of ablation rate”, in Fuel, 1985, vol. 64, pp. 15 14-1520. LWC J, “The Cyclone: A Multifunctional Reactor for the Fast Pyrolysis of Biomass”, Ind. Eng. Chem. Res., 2000,39, 893-903. Kovac R.J. and ONeil DJ, ’The Georgia Tech Entrained Flow Pyrolysis Process,“ Pyrolysis and Gasification, G.L. Ferrero, K. Maniatis, A. Buekens, and A.V. Bridgwater, eds., Elsevier Applied Science 1989, pp. 169-179. Maniatis K, Baeyens J, Peeters H and Roggeman G, ”The Egemin flash pyrolysis process: commissioning and results”, pp 1257- 1264 in Advances in thermochemical biomass conversion, Ed. AV Bridgwater (Blackie 1993) Prim W and Wagenaar BM, In Biomass gasification and pyrolysis, Eds. Kaltschmitt MK and Bridgwater AV, pp 3 16-326 (CPL 1997) Wagenaar BM, Venderbosch RH, Carrasco J, Strenziak R, van der Aa BJ, “Rotating cone bio-oil production and applications”, these proceedings Yang J, Blanchette D, de Caumia B, Roy C “Modelling, scale-up and demonstration of a vacuum pyrolysis reactor”, These proceedings. OFFER announcement, February 1997, (Office of Electricity Regulation, UK) 995
Diebold JP, Czemik S, Scahill JW, Philips SD and Feik CJ. “Hot-gas filtration to remove char from pyrolysis vapours produced in the vortex reactor at NREL”, in: Milne, TA, editor, Biomass Pyrolysis Oil Properties and Combustion Meeting, NREL, 1994, p. 90-108. 32 PyNe Newsletter 3, March 1997, Aston University, UK pp21 33 Ltdt J, Diebold JP,Peacocke GVC, Piskorz J, (1996) “The Nature and Properties of Intermediate and Unvaporized Biomass Pyrolysis Materials,” in Developments In Thermochemical Biomass Conversion, Bridgwater AV and Boocock DGB, (Eds.) (Blackie Academic & Professional, London 1997). 34 Huffman, D.R., Vogiatzis, A.J. and Bridgwater, A.V., “The characterisation of RTP bio-oils”, Advances in Thermochemical Biomass Conversion, Ed. A V Bridgwater, (Elsevier, 1993) 35 Diebold JP, Milne TA, Czernik S, Oasmaa A, Bridgwater AV, Cuevas A, Gust S, Huffman D, Piskorz J, “Proposed Specifications for Various Grades of Pyrolysis Oils”, In Developments in Thermochemical Biomass Conversion; Bridgwater AV and Boocock DGB, Eds. (Blackie Academic & Professional, London 1997), pp.
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Shihadeh AL, Rural Electrification from Local Resources: Biomass Pyrolysis Oil Combustion in a Direct Injection Diesel Engine, Ph.D. Thesis, (Massachussetts Institute of Technology, 1998) 37 h a M, Slamak M, Sawatzky H, Pyrolysis Liquid-in-Diesel Oil Microemulsions, US Patent 5,820,640, (1998) 38 Baglioni P, Chiaramonti D, Bonini M, Soldaini I, Tondi G, “BCO/Diesel oil emulsification: Main achievements of the emulsification process and preliminary results of tests on diesel engine”, These proceedings 39 Diebold, J.P. and Czernik, S., Additives to Lower and Stabilize the Viscosity of Pyrolysis Oils during Storage, Energy & Fuels, (1997), 11, 1081-1091. 40 Radlein D, Piskorz J, Majerski P, Method of upgrading biomass pyrolysis liquids for use as fuels and as a source of chemicals by reaction with alcohols, European Patent # 0718392, (1999. 41 Diebold JP, “A review of the chemical and physical mechanisms of the storage stability of fast pyrolysis bio-oils”, Report for PyNe, 1999. (To be published in 3G
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Maggi R and Elliott D, In Developments in Thennochemical Biomass Conversion, Bridgwater AV and Boocock DGB Eds. (Blackie Academic & Professional, London 1997), pp. 575-588. 43 Bridgwater AV, “Production of high grade fuels and chemicals from catalytic pyrolysis of biomass”, Catalysis Today, 29,285-295 (1996) 44 Bridgwater AV, “Catalysis in thermal biomass conversion”, Applied Catalysis A, 116, (1-2), pp5-47, (1994) 45 Elliott DC and Baker E, In Energy from Biomass and Wastes X, Klass D, Ed., (IGT, 1983), pp. 765-782. 46 Chang C and Silvestri AJ, Catalysis, (1977), 47, p. 249. 47 Bridgwater AV and Cottam ML, “Costs and Opportunities for Biomass Pyrolysis Liquids Production and Upgrading”,Proc 6th conference on Biomass for Energy, Industry and the Environment, Athens, (April 1991). 48 Bandi A, Baumgart F, “Stirling engine with flox burner fbelled with fast pyrolysis with fast pyrolysis liquid”, These proceedings 49 Oasmaa A, Kyto M, Sipila K, “Pyrolysis oil combustion tests in an industrial boiler”, These proceedings
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Pakdel H, Murwanashyaka J N, Roy C, “Fractional vacuum pyrolysis of biomass for high yields of phenolic compounds”, These proceedings Himmelblau A “Combined chemicals and energy production from biomass pyrolysis”, These proceedings Czernik S, French R, Feik C, Chornet E, “Production of hydrogen fiom biomassderived liquids”, These proceedings Dobele G, Rossinskaja G, Telysbeva G, Meier D, Radtke S, Faix 0, “Levoglucosenone - A product of catalytic fast pyrolysis of cellulose”, These proceedings Venderbosch RH, Wagenaar BM, Gansekoele E, Sotirchos S, Moss HDT, “Cofiring of bio-oil with simultaneous SOXand NOx reduction”, These proceedings Leech J, “Running a dual fuel engine on pyrolysis oil”, pp 495-497, In: Kaltschrmtt MK and Bridgwater AV, editors, Biomass gasification and pyrolysis: State of the art and hture prospects, (CPL Press, 1997) PyNE Newsletter No.4; (Aston University, UK) Andrews R., Patnaik PC, Liu Q. and Thamburaj, “Firing Fast Pyrolysis Oils in Turbines”, Proceedings of the Biomass Pyrolysis Oil Properties and Combustion Meeting, Eds. Milne, T., (NREUCP-430-7215, 1994), pp. 383-391. Strenziok R, Hansen U, Kunstner H, “Combustion of bio-oil in a gas turbine”, These proceedings PyNe Newsletter 4,pp 2 1, (September 1997), Aston University, UK Radlein, D. and Piskorz, J., “Production of chemicals from bio-oil”. In: Kaltschmitt M, Bridgwater AV, editors, Biomass Gasification and Pyrolysis, (CPL Press, 1997), p 471-481. Pernikis P, Zandersons J, and Lazdina B, “Obtaining Levoglucosan by Fast Thennolysis of Cellolignin-Pathways of Levoglucosan Use,” in Developments In Thermochemical Biomass Conversion, Bridgwater AV and Boocock DGB, (Eds.) (Blackie Academic & Professional, London 1997) Meier D, Andersons B, Irbe I, Tshirkova J, Faix 0, “Preliminary study of fhgicide and sorption effects of fast pyrolysis liquids used as wood preservative”, These proceedings Radlein, D, Piskorz J and Majerski, P, “Method of Producing Slow-Release Nitrogenous Organic Fertilizer from Biomass” US Patent 5,676,727, 1997 and European Patent Application 07 16056.
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Test Bed to Turnkey - The Introduction of New Thermal Renewable Energy Technologies I Burdon PB Power Ltd (Merz and McLellan), Amber Court, William Armstrong Drive, Newcastle Business Park, Newcastle upon Tyne, NE4 7 Y Q United Kingdom
ABSTRACT. The step between a laboratory-scale research project and a 111-scale, commercially-viable,scheme which will attract financial support from investors is large. Considerable effort needs to be expended by the developer to demonstrate that the risks involved in moving fiom test-bed into a competitive, environmentallysensitive, real world are mitigated to the maximum extent and borne by the project participants most able to control them. The paper considers a generic fuel-energy conversion process and describes the commercial climate in which such a project has to develop, grow and survive. The trends in the world’s electricity supply industry, from state-controlled utilities to liberalised, commercial, regimes in which competition plays an important part are explained. Key ingredients for a successful project are identified and the way in which independently-ownedpower (ie electricity) projects are structured is described. The requirements of the financiers of such projects are highlighted. The author presents his views from the standpoint of an “independent engineer”; his organization often being employed in such a role by project financiers to appraise energy projects employing novel technologies. It is anticipated that the information given in the paper will be of use to individuals or organisations who are seeking financial support to take an advanced-technology process from a proven design concept to commercial operational service.
INTRODUCTION Commercial viability is a fundamental requirement for any process or project which is intended to supply useful energy into the competitive market places which, nowadays, characterise the electricity supply industries in many parts of the world. Most independent power projects whch compete in these markets are financed “off-balance sheet” ie they rely on debt provided by banks and similar financial institutions to provide 70-85 per cent of the project’s funding. The banks will require all risks associated with the project to be identified and mitigated or accepted by others; they will have a thorough appraisal of the project’s technical and commercial aspects conducted by independent engineers and lawyers with all “due diligence”.
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This paper describes the project development process and attempts to hghlight the commercial climate in whch an energy project, based on a thermal conversion process, has to become established and survive. The detail of the thermal conversion process itself, ie the “technical” process, is not relevant to the conception or the understanding of the project development process, ie the “financial” process, although the former process will have a major d u e n c e on the project’s risk profile whch will come under intense scrutiny in the latter process. A GENERIC PROCESS MODEL
Almost all thermochemical biomass conversion processes can be considered and thus analysed as a “black box” with a feedstock input and an energy output (either thermal or electrical). Useful co-products will arise in the operation of the process (eg sulphur, acetic acid, methanol even food flavouring!) from which revenue may be earned. Relatively useless by-products (eg ash, frit and other emissions to air, land and water) which arise may incur a cost in their disposal if a market for their use cannot be found. Certain defined design and operational parameters will be intimately associated with the black box and these are shown on the diagrammatic representation of the overall process shoM in Fig 1. The parameters can be considered of two types: (1) “Hard” - ie those which are amenable to scientific analysis and, therefore, fairly precise calculation or prediction. (2) “Soft” - which may be largely qualitative and, accordingly, subject to opinion and be affected by judgement or simple prejudice.
Fig. I . Black Box Approach to project operational evaluation and risk assessment
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All of the hard parameters will need to be taken into account in the financial model of the scheme which will evaluate the casMows, in terms of costs and revenues to calculate the financial return on the investment in the project. Two vitally important hard parameters associated with any conversion process, which have a critical effect on project viability, are efficiency and availability. These are defined as follows: efficiency = availability =
useful electricity exported heat value of feedstock period time maintenance time forced outage time
-
-
period time
NB efficiency should be measured at defined ISO-type conditions and forced outage time includes time at reduced capacity and the attainment of high availability is critical in maximizing the energy produced from the process and, therefore, the revenue. The financial model is characterized, in summary, as follows: (1) Based on selected technology (2) Includes all project development costs (3) Includes costs of the site, insurances, working capital and contingency allowance (4) Allows for inflation and taxes ( 5 ) ‘Gives IRR/NPV, casMow and cover ratios (ie the ratio of net project revenue to debt service costs) on project investment.
The model invariably involves many iterations before a satisfactory combination of project parameters is assembled; it will set the scene for the business case which the investor or the financier will consider prior to committing funds to the project. Further discussion on the financial model is given in the section on risks and critical issues later.
THE REAL WORLD Once upon a time, electricity supply was a monopoly activity, the industry’s assets were owned by the state in most countries or, where they were not state-owned, such as in the US or Germany or Japan, they were heavily regulated. In the 1980s, fundamental changes took place in the state-owned sector with the so-called “privatization” of the industry. First in Chile, then in the UK, the assets of the publicly-owned electricity authorities were vested in companies whch issued equity shares to the general public. This change in ownership was coupled with a fimdamental shift in the way electricity was produced and sold. New, cheaper, technologies such as gas-fired, combined cycle, gas turbine plant, which could be built more quickly than the traditional coal-fired steam turbine plant, and which is less labour intensive in its operation, was introduced extensively in the UK. The separate producers of electricity competed with each other on price (per kwh) to win market share. It is alleged, with some justification in the UK, that such competition has reduced the price to the end-user. 1000
In such a hlghly competitive market place, expensive technology, and that which is unproven or of a small-scale, will find it very difficult to compete with tried and tested CCGT plant fired on plentiful and (currently) cheap natural gas. In the developing world, the requirement is for large increases in reliable capacity to provide power for industrial development, new townships and agriculture. Governments in such countries, in most parts of the world, are encouraging private sector developers to invest in the national electricity system and sell power to the distribution companies for onward supply to consumers. Again, conventional technology, utilising lowest cost, often indigenous, fuel such as coal, is a powerful imperative in the thmking of private sector investors. Thus, it is suggested, expensive complex technology will be hard pushed to find a secure place in a market place where the emphasis is on low capital cost and technical efficiency for large-scale power production. One area whch, however, does offer some chance of success for advanced conversion technology lies in the field of renewable energy. The ability of a process to accept a feedstock which is sustainable, such as biomass, or can be considered, to all intents and purposes, as a renewable such as waste arisings eg sewage sludge, municipal solid waste or food waste, from large centres of population, can be beneficial for technological development in the context of a national renewable energy policy. Such a policy manifests itself, for example, in the UK at the present time where Government has set targets for 5 per cent of all electricity production to be from renewable energy sources by 2003, rising to 10 per cent in 2010. Thermal conversion of biomass will play a part in meeting these commitments, along with wind energy, hydro, etc. Premiums paid for renewable energy from such sources can compensate for the inevitable h g h capital cost of the technology and the disadvantages arising from dis-economies of scale. A further important issue to be considered in a “real world” environment is not only competition in the market place for the output of the conversion process (ie electricity or heat) but competition between rival technologies for the attention and sponsorship of project developers. Factors whlch are likely to be considered by developers in selecting a particular technology will include: Previous operating experience - where the “technical” process has been developed from laboratory prototypes, data will need to be gathered on the performance of the process in t e r n of its thermal efficiency in recovering useful energy from the feedstock and the production of by-products (and co-products if relevant), particularly discharges to air, water and land which will be of keen interest to the environmental regulatory agency. Inevitably, such operating experience will, generally, relate to a small scale plant and predictions on performance and operating experience will have to be made where scaling-up is envisaged (see below). Scale of process - those investing in a project where significant scaling-up is needed to attain critical “commercial” mass will require evidence that such a scale-up is a practical proposition and that the relevant factors have been taken into account, both from a mechanical engineering aspect and from the thermochemical conversion standpoint. Sometimes, a multiple parallel-flow process may be adopted to avoid having to stretch the design too far. The risks lnherent in the scaling-up will be of special interest to those involved in the independent project appraisal.
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Current status of technology - the .development status of the technology is interrelated with the feedstock type.to a large extent. For example, whilst the gasification of a coal feedstock is a relatively well understood and well tried process, the gasification of many biomass-type wastes is not, certamly outside of a research laboratory. The question generally asked under this heading in the appraisal process is “how long has the technology been around and where can we see it working?” Cost/MW of capacity - given that the process will be competing for developers funds against a number of other technologies, there will have to be significant advantages, in terms of efficiency or environmental emissions, to be gained from relatively expensive processes if the attention of a developer is to be caught. Potential application - it should be remembered that a conversion process can, fundamentally have one of three potential uses:* electricity generation 0 feedstock utilization or disposal production of co-products. Each of these aspects needs to be quantified in the project evaluation process.
FINANCING A PROJECT The majority of independent energy projects which have been promoted in countries such as the UK, which have a non-parastatal electricity supply industry, have been funded on the so-called “project finance” basis. Under such an arrangement (which was developed originally in connection with North Sea oil projects in the 1960s) the majority (in the region of 80 %) of the finance required to implement the project comprises debt borrowed over a period of, typically, 12-15 years. Bond financing can extend this period to 25 years in certain cases. The debt is repaid from the net revenues of the project. The remainder of the finance required for the project comprises equity provided by the project sponsors.
SPECIAL PURPOSE COMPANY In order to provide a vehicle to take the project forward on a project finance basis, a special purpose company is established by the project sponsors. This company (the “energy company”) becomes the counter-party to the various agreements for goods and services whch, essentially, comprise the project (see Figure 2). The primary agreements in a typical renewable energy project are: (1) (2)
(3) (4)
Fuel or feedstock supply heat and electricity off-take design and construction of the facility operation and maintenance.
There are other secondary agreements covering lease or sale of land, finance etc as well as statutory consents such as planning, integrated pollution control, consents to build a power station, licence to generate electricity etc.
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Fig 2 Contractual arrangements for independently financed energy conversion
projects The primary agreements are written in such a way that each complements one another eg the specification and quantity of feedstock supplied corresponds with that required by the plant to be built under the design and construction (D&C)contract to deliver the electrical energy output specified in the electricity off-take contract. At the same time, the obligations contained within the operation and maintenance (O&M) contract require specific levels of fuel conversion efficiency and plant
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availability to be attained by the operator in line with the terms and conditions of the electricity off-take contract. The essential elements of each of these primary contracts are summarized below: FUEL SUPPLY
Specific quantities of feedstock, of a defined quality, are to be made available under specified conditions at a particular price. A "gate fee" ie a payment from the feedstock supplier to the project company is generally paid in the case of a project in which the disposal of a waste (eg a by-product from a food processing industry) is a fundamental purpose of the project. Some form of "put or pay" arrangement is generally included in the contract in such circumstances. Where the feedstock has an intrinsic value e.g. it can be used as a fuel in other processes (and, thus, there is some competition between procurers), the independent assessors of the project will require to see evidence of its long term, reliable, delivery in sufficient quantity to satisfy the needs of the project. Some measure of price stability will be necessary, for example by linking fuel purchase price to recognized and appropriate cost indices. Diversity of supplier is generally desirable, whether there is a gate fee payable by the supplier or the feedstock has to be paid for by the project company, so as to mitigate the risk of a single supplier going out of business. The duration of the fuel supply agreement usually exceeds the term of the finance. ELECTRICITY OFF-TAKE An electricity off-take contract in a renewable energy project in many countries which offer support to renewable energy is usually a pro-forma agreement with one or other of the generating companies (or perhaps electricity supply companies) for a fixed term sometimes at a premium price. In the UK, for example, the Non-Fossil Fuel Obligation under the 1989 Electricity Act obliged the electricity distribution companies to procure a certain amount of the electricity that they supplied to their customers from renewable sources. Similar obligations exist under the recently introduced Utilities Act.
DESIGN AND CONSTRUCTION
The D&C contract in any project financed on a project finance basis is, invariably, based on a turnkey arrangement. This ensures that a single contractor assumes an undivided responsibility for completion date, final cost and performance of the facility. Liquidated damages are payable by the contractor for delay in completion or shortfall in performance. Such a contract is based on a carefully-written performance specification issued by the project company against which the contractor submits a fixed price tender in competition with other contractors. The contract is generally arranged so that the turnkey contractor accepts all the major construction risks. OPERATION AND MAINTENANCE
Given that the payments made to the O&M contractor are llkely to be a small fraction of the revenues earned by the project as a whole from electricity sales, it is not usually possible to secure damages from the O&M contractor for failure to achieve set performance targets for output (ie efficiency and availability) which are 1004
commensurate with the value of the lost electricity output. The essence of an O&M contract, therefore, is the establishment of, and commitment by the O&M contractor to, an agreed quality assurance system covering the activities of the contractor such as organization, documentation, safety, training, reporting, planning etc. Regular audits of the contractor’s activities and procedures give confidence to the owner that the contractor is likely to achieve the required performance targets. There may also be financial incentives included in the contract to encourage the O&M contractor to meet the desired level of performance and, conversely, penalties for nonachievement. CONSENTS
Certain consents will be required by the project owner or operator prior to putting the facility into use. These, typically, include in the UK for example Integrated Pollution and Prevention Control Authorization (IPPCA) for thermal conversion processes, water abstraction licences for hydro electricity schemes etc. Assuming the use of proven technology and competent D&C and O&M contractors, such consents may not be difficult to obtain, unless the project is sited in a particularly environmentally sensitive area. In the case of pyrolysis processes, it will be necessary to convince the environmental regulators that the organic constituents of a pyrolysis oil-fuel (some of which might be carcinogenic) are destroyed in the overall energy production process. An environmental statement is required under planning legislation in certain countries for all electricity generation projects and consultation with appropriate organizations may be necessary as part of the procedure for obtaining planning permission. RISKS AND CRITICAL ISSUES The main categories of risk associated with a renewable energy project are listed in the Appendix. The majority of risks associated with a given project can be covered in the contractual arrangements described in the preceding section. In this way, risks are allocated to the party who is best able to deal with them. The construction contractor, for instance, will typically be called upon to provide a guarantee of the performance of the plant, with appropriate damages for non-performance which protects other participants from financial loss. Project financiers will wish to see a suitable spreading of risks and mitigation of them (perhaps by insurance) over all the project participants. It is highly unlikely that the banks responsible for the provision of debt to the project will agree to bear any of the risks themselves. RISK ANAL ysrs
The interests of lenders, insurers and contractors require a very carell analysis of risk in the context of the various contractual obligations that go to make up the structure of a project. Matters which would typically be regarded by lenders as representing a potential risk include: (1) Technology - lenders will expect an independent engineer to report on the
technical specification for the plant and will require evaluation of the
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appropriateness and track record of the relevant technology in the proposed application. The developer may be enthusiastic about the potential benefits of a new technology whereas the lender would regard untried features as an unacceptable risk; (2) Contractors - lenders’ interests extend beyond just the identity of the principal contractors. They will consider their reputation and financial strength in assessing whether there is any risk of default in relation to any function which is material to the project’s success; (3) Security of cash flow lenders will analyse all factors which may have a bearing on the ability of the project to maintain the project cash flow comfortably beyond the prospective term of the borrowing including, for instance, as mentioned in the previous section, the security of the fuel supply and contingency arrangement for alternative fuel; (4) statutory authorizations - lenders will check on all factors affecting the consents and permits necessary for the commencement and continuation of the project.
-
The success of the developer in structuring and obtaining finance for projects is closely linked with their ability to identi@ areas of risk and to design technical solutions and allocate contractual responsibility so as to create the greatest comfort in the light of perceived risk. Much of this can be achieved by seeking initial expert advice in each relevant field so that there is a coherent strategy from an early stage. Some of the relevant issues which relate to risk analysis include: (1) Financial model - a crucial test is to see how the economics of a project stand up to the perceived areas of risk. The development of any project will involve the production of a financial model which shows how the sensitivities of the various parameters at risk affect the economics. This is a process which will be followed by any prospective lender in relation to a funding proposal. (2) Contractual framework the contractual framework for any project will need to take account of the risk factors which are relevant to the specific features of the project. It involves careful planning on the identity of the contracting parties and assessment of their capacity to discharge their obligations, the scope of the positive obligations under each contract and the effect of limitations of liability. (3) Financial structure - the funding structure will require a detailed evaluation of the capital requirement and the balance between equity or “risk” capital and loans. The proposed “gearing” will indicate the potential reward for the providers of equity and will also be an important indication in any proposal for project finance of the way in which the developers have considered the allocation of risk. Lenders like to see proposals which show that these principles have been understood and will gain confidence in the developers if the proposed financial structure of the project takes them into account. (4) Insurance - insurance cover will be available for many of the areas of perceived risk and it is important that the scope of the cover is planned with reference to the specific detail of the project. This will, in any case, be a requirement of the lender. Specialist insurance brokers can assist in creating an effective insurance propamme covering relevant interests with an indication of the sort of scope of cover that lenders will wish to see.
-
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As a final note on this subject, the difference between risk and uncertainty is frequently misunderstood and the terminology, accordingly, mis-applied. The following note gives a clear enunciation of the difference:
Risk has its place in a calculus of probabilities. It applies to a specific course of action. The risk of an action is the likelihood that it will produce an unwanted result. Risk lends itself to quantitative expression, as when we say that the chances offailing to strike oil in afield are better than fiftyf i f v or that the chances offinding a defectivepart in a batch are two out of a hundred. In the framework of benefit-cost analysis, the risk of an innovation is how much we stand to lose if we fail, multiplied by the probability offailure. Uncertainty is quite another matter. A situation is uncertain when it requires action but resists analysis of risks. A gambler takes a risk in an honest game of blackjack when, knowing the odds, he calls for another card. The same gambler, unsure of the odds and of the honesty of the game, is in a situation of uncertainty. He can act, but he cannot estimate the risks or rewards of his action. Even so, he operates in a risklike situation because, at any rate, he has two well-defined alternatives - to call for another card or not to call for it. He must act even though relevant alternatives are undefined. He must invent what to do. He has no way of calculating with any precision the risks of action. He has only rough guidelines of skill and experience to help him. (1) CONCLUSIONS Translating thermal conversion research into a full-scale, practical, process plant in which all technical risks are understood and capable of quantification is bound to be a difficult task, but every new development has to start somewhere. The commercial world of electricity supply, in which the thermal conversion process hopes to find an application, is changing rapidly in the developed world and the support and sustenance that prototype processes might have received in years gone by from central authorities is now vanishing. In the developing world, the requirement is generally for tried and tested, reliable and easily maintained technology capable of providing a supply of electricity utilizing indigenous fuels. Advanced technologies, not long out of the research stage, are unlikely to satisfy their demands. The dnve to utilize more non-fossil fuels in electricity production, in order to meet climate change accords, is likely to provide the main stimulus to technological development of new thermal conversion processes which can utilize biomass or waste products as a feedstock in the developed world. The financing of such projects, unless heavily underwritten as a stimulus to their commercial application from grant support of their capital cost from public funds, will require intensive analysis of the intrinsic technical risks of the process. The satisfaction of such a risk analysis will require the production of comprehensive data (and its intensive scrutiny) on operational performance and design, particularly in relation to scaling-up from the test-bed to the turnkey, commercial, installation.
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ACKNOWLEDGEMENTS Thanks are due to the Directors of PB Power Ltd (which now includes the power businesses of both M e n and McLellan and Kennedy & Donkin) for permission to publish this paper. It is emphasized that the opinions expressed are entirely those of the author and in no way reflect the corporate attitudes or views of PB Power. BIBLIOGRAPHY 1.
A guide to finance for the UK renewable energy industry. ETSU for the DTI, London, 1997.
REFERENCES 1.
Schon DA (1967) Technology and Change: The New Heraclitus, Delacorte Press, New York
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APPENDIX Categories of risk in a typical renewable energy project Major category Sponsor
0 0 0 0
Construction
0 0
0 0
0
Completion
0 0 0
0
Technology
0 0 0 0 0 0 0
Specific aspect track record and reputation as a developer degree of financial commitment competing interests insurability time, cost, quality sponsor’s ability to arrange construction experience and standing of contractor terms of main construction contract fixed pricelturnkey penaltiesibonuses performance warranties and damages insurances incentives force majeure commissioning subcontracting competency of contractor’s organization site conditions infrastructure competence of project manager labour availability and suitability previous applications competitive options robustness flexibility adequacy of guarantees and warranties achievability of performance targets maintainability
~
Fuel supply
0 0 0
0 0
Operation maintenance
and
0
0 0 0
0 0 0
0 0
availability specification price and exposure to cost increases delivery and deliverability terms and duration of fuel supply agreement experience of operator adequacy of infrastructure maintenance of performance levels access to replacement supplies and parts cost constraints exposure to cost increases quality management system appropriate maintenance schedules terms of O&M agreement
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Major category 0
Electricity and heat offtake
a a
a a
a a 0
a
a a a
Political a a
Specific aspect emission monitoring quantity product reliability capacity chargelenergy charge market price levels and stability competition indexation term of contract predictability of heat and electricity loads take or pay penalties for non-supply status of off-taker potential for export off-site energy policy legislation taxation environmental controls consents and authorizations
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Woody and Herbaceous Biomass Feeds how can we study their composition and their pyrolysis products? B. Krieger-Brockett* and I. Rodriguez, University of Washington, Dept of Chemical Engineering, Box 351 750, Seattle, WA, 98195-1 750 USA
ABSTRACT: Laboratories have renewed the quest for robust measures of alternative feedstock composition that are predictive of the pyrolysis products, particularly oil (sometimes called biocrude). A level beyond that predicted by pyrolysis conditions is sought. By alternative feedstocks we mean agricultural residues and other herbaceous or woody species that are infrequently commercially exploited. We have obtained a unique sample set of over 1100 feedstocks -- collected, pretreated, and characterized by the US Department of Agriculture principally for their ash and extractives (protein, polyphenol, oil, hydrocarbon) content and composition. We subsequently studied the sample set’s pyrolysis behavior -- temperatures, gross product yields and compositions within the product fractions -- using an instrumented single particle reactor according to an experimental design. The composition attributes and pyrolysis conditions were systematically varied over a range of industrial interest. The experiments were designed using Principal Components Analysis to find feedstock compositions that were unambiguously attributable to one of the extractives. The resulting pyrolysis product slate was then similarly (statistically) analyzed for patterns that related to fed composition. Findings indicate that the native protein content, ash, and oil are the most important extractables in determining the pyrolysis products (oil, char, and gases) from these under-utilized herbaceous and agricultural residue species. INTRODUCTION AND BACKGROUND Expanding renewable and biomass feedstocks to non-traditional agricultural residues requires answering the question posed in the title, restated: What are unambiguous ways to study biomass composition and its relationship to pyrolysis rate and products? Are traditional, operational definitions of composition such as cellulose, Klason or Kraft lignin, and extractives weight fraction useful? What other composition metrics are useful especially as under-utilized or transgenic biomass species are examined? The traditional pulp/paper composition metrics (e.g. specific lignin and cellulose contents) are based on chemical reaction (depolymerization) of the native biomass into fractions. Each separated hction is then independently pyrolysed, and the pyrolysis yields are fit to models, and the composite used to predict other woody species’
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pyrolysis product slates. For example, powders from the same species (sweet gum hardwood) were used in the study of cellulose (Hajaligol, et al. 1982), lignin and whole wood (Nunn, et al. 1985a, 1985b) pyrolysis product slates. Each sample’s decomposition products were measured using a heated grid apparatus. Gross product yields from the wood were correlated with parallel independent reaction kinetic parameters for each of the traditional components, lignin and cellulose. This construct was used even though lignin is not a component, but highly dependent in both molecular weight and structure upon the process used for its separation (pulping). This approach makes use of the extensive knowledge base in the forest products industries and has the merit that it is a familiar approach. However, even moderate pulping alters the polymeric constituents sufficiently that separate pyrolysis of wood fractions does not reproduce the native biomass pyrolysis chemistry and nor reproduce the physical microstructure changes that occur within large biomass wood chips during pyrolysis. Thus, a narrow range of applicability limits the traditional approaches. To answer the question posed in the paper title, we offer a way to study composition effects when non-traditional biomass feedstocks are used in pyrolysis or energy production. Our method builds on traditional approaches and the whole or native biomass compositions as studied by scientists (e.g., Bagby et al, 1980) at the US Department of Agriculture (USDA) in a 1Cyear screening program fix “botanochemicals”. Bagby and co-workers’ objectives included identifying and exploiting minor chemicals, such as extractives, (we will use this term for collectively referring to protein, ash, polyphenol, and oil fractions) from agricultural residues with the remainder, mostly cellulose and lignin, available for pulp/paper or keldenergy. We ask a related question, i.e., is extraction, or alternatively leaving the components within the biomass, an advantage to pyrolysis and what pyrolysis product fractions am altered? Our methods and experiments (UW) previously addressed composition effects in pyrolysis of RDF (Lai, et al 1993) and wood (Krieger-Brockett, et a1 1997). In those papers and this one, even minor components are shown to alter pyrolysis slate when appropriate statistical methods are used. This paper briefly summarizes our work on pyrolysis product slates resulting from large- or macro-particle devolatilization (in which heat transfer is a slow process) of native biomass compositions in under-utilized species. The method has general applicability and owing to the limited scope of this article, the reader is referred to Somasundmam (1990), Lai (1991) and Rodriguez (1996) for details and extensive literature reviews with only a few relevant articles mentioned here. SELECTED LITERATURE AND PREVIOUS WORK
Studies have focused on pyrolysis of wood’s separated polymeric constituents’, i.e., powdered wood, cellulose, hemicellulose and lignin (e.g. Antal, et a1 1985; Nunn, et al 1985). The effects that minor components such as extractives’ have on whole biomass pyrolysis products have not often been studied, except for ash (e.g., Richards and Zheng, 1991). Using regression analysis to quantify the intrinsic cellulose, lignin, and ash 6actions’ effects on a few products, MacKay and Roberts (1982) studied the pyrolytic behavior of 20 unmodified or “native” powdered lignocellulosic materials. Linear regression models for total mass and carbon yields of the products were developed, assuming each constituent fraction of the biomass was pyrolyzed independently. They concluded that char yield increased with substrate carbon content; char yield fiwn lignin content was found to be three times that of cellulose content. They did not study the gas nor tar composition in their experiments. In addition their particle sizes
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were inapplicable (powder) to industrial biomass processes wherein large particles and large temperature gradients (“thermally thick” particles) predominate. Davis et a1 (1985) screened 57 different woody and herbaceous biomass feedstocks for their production of liquid fuel constituents fiom indirect liquefaction of powdered materials in a fluidized bed. Composition metsics in Davis’ studies were ash, proteins, polyphenols, oil, and hydrocarbon content as defmed and measured by the USDA (for example, in Adams, et al, 1986; Buchanan, et aI, 1980a; Roth et al, 1984; Swanson, et a1 1979). We will call the USDA-measured compositions “traditional” for purposes of this article. These feed compositions were correlated with pyrolysis gas production of hydrogen, CO, CzH+ total olefms, paraffins, and HJCO ratio using direct regression equations of the quadratic type, but tadchar production was not addressed. To characterize agricultural species with chemical compositions suitable fm conversion to bio-based chemicals and liquid fuels, a project lasting more than 14 years was conducted at the USDA. Over 1100 promising biomass species were systematically harvested and characterized for their content of ash, polyphenol, oil, protein, rubber, etc (and other “traditional” composition measures such as botanical family) (Bagby, 1989). In our current research, we make use of the USDA homoscedastic (based on standard methods performed by the same lab) biomass composition database to select species of contrasting and suitable chemical composition. Then, using statistical analysis to treat co-varying composition variables, we relate the resulting pyrolysis product yields to the traditional compositions of biomass, based on the USDA measurements. Because of the taxonomic information, as well as the extractives’ fktions provided by USDA, the composition metrics are likely to have wide applicability. Other related research was conducted at the National Renewable Energy Laboratory (NREL)to characterize and study the performance (after harvest and storage) of certain herbaceous biomass species with respect to the production of chemicals and hydrocarbons (e.g., Agblevor, et al, 1995,1992; Evans and Milne, 1987a,b). Of course, interest has increased regarding pyrolysis of genetically-engineered (transgenic) biomass species with special traits such as high isoprenoid content or easily-separated lignin. Biomass pyrolysis oils have received attention owing to their potential value once upgraded (Bridgwater and Cottam, 1992) and their similarity to conventional fuels. Biomass oils are a very complex mixture of organic acids, oxygenated hydrocarbons, phenolic compounds and aromatics with precise composition being dependent on pyrolysis process, the rate of reaction and the feedstock chemical composition. Rodriguez (1996) reviews the literature on biomass pyrolysis oil analysis, with only a few studies mentioned here. Workers have used solid phase extraction, gel permeation chromatography (GPC), GC-MS, inhed spectroscopy, elemental analysis and solubility classification for characterization. Solid phase extraction on silica gel is useful for the separation of pyrolysis oils into functional fractions; GPC provides molecular weight distributions. Elliott et a1 (1988) have observed as many as 360 peaks in a pyrolysis oil by using GC-MS. Williams and Home (1995) identify and quantified 161 compounds in biomass pyrolysis oils by using solid phase extraction and GC-MS. Although we too measured oil composition, space limits our discussion to results pertaining to gross product hctions (char, tar or oil, and gas yields) as affected by the native composition as discussed below.
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MATERIALS AND METHODS SAMPLE PREPARATION The USDA characterized a large number of biomass species for their traditional composition fractions of ash, crude protein, polyphenols, oils and hydrocarbons among others. These fractions were defined operationally by the USDA “botanochemical screening” project; their sample analysis and hction partitioning scheme is summarized in Fig. 1. Of interest were extractives, components that can be separatedlpartitioned fiom the plant by solvents. The major extractives included various oils, terpenes, fatty acids, unsaponifiables, aromatic compounds, tannins, and quinones. In exceptional cases extractives composed over 15% of the biomass (especially ash) but generally they did not exceed 510%. The samples collected by the USDA (Fig. 1) consisted of whole plants, and for trees and shrubs, the latest 2 to 3 years‘ growth. All portions were included except the roots. Each sample was dried in still air at ambient conditions. The entire sample (about 500-1000 g, moisture free basis) was milled (Wiley) and screened using l-mmdiameter holes (Carr et al. 1986). A subsample of each milled species was analyzed fix moisture, ash, and crude protein (% Kjeldahl nitrogen x 6.25). Another portion of the sample (approx. 100 g) was extracted in a Soxhlet apparatus, first with acetone for 48 h and then with hexane for another 48 h. The acetone fiom the frst extraction was allowed to evaporate before proceeding with the second extraction. The acetoneextracted hction was partitioned between hexane and watedethanol (1:7) to obtain the oil and polyphenol fiactions. The residue after acetone extraction was extracted with hexane to yield the hydrocarbon fraction. The yields in all cases were determined gravimetrically after solvent removal and oven drying, and are reported as percents CE the moisture-fiee and ash-& plant sample weight (Carr et a f . 1986). We adhere to these operational definitions of the fractions and use the USDA data and samples directly. Scheme for Paltltlonlngwhole-plant Samples
extractives
,A fats fa eclds re2ns terpenes waxes.
residue
&]+d+, polyphenollcs tannins complex lipids
oil and hydrocarbon-free residue isoprene polymera waxes terpenoids
cellulose llgnin k? Gkic (ash)
Hydrocarbon .ndRubber Pmdudng Cmp. EvduaUonot US. Plant. R.A. Buchuvn at. d. EumrmkBo(.ny32:131-146. i978.
Figure 1 USDA scheme for extraction and partitioning of whole-plant material
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SINGLE PARTICLE REACTOR A brief description of the University of Washington single particle reactor, analysis system, and methodology is here, but the reader is referred to the more complete descriptions in Lai et al (1993), Lai (1991) and Krieger-Brockett et a1 (1997), and references cited therein. The glass portion consists of a reactor head with inlet ports fix carrier gas, and an exit port for volatiles from the reaction zone. The fused silica window of the reactor allows for maximum transmission of radiant heat. Helium is fed though the inlet ports to the pellet front to sweep the volatiles away from the reacting zone and to the back face of the pellet to force volatiles “forward“ for analysis. A glass sleeve holds the biomass pellet tightly while it is heated axially by a high intensity Xenon arc lamp. The arc lamp provides a uniform, constant radiant heat flux on the front of the biomass pellet and the lamp is calibrated absolutely, spatially, and fiequently. During the heating, we visually observe the volatiles swept out of the reactor through the top port of the reactor head. The fluid residence time is short (measured by Chan, et al 1988; Lai, 1991). The volatiles pass through an accumulating tar trap at - 40 O C (dry ice-acetone bath) which also traps water. The permanent gases pass through the trap and go to a semi-automatic GC analysis system. Cylindrical pellets are made from the USDA samples expressly to remove consideration of native biomass morphology. We have shown biomass morphology alters pyrolysis slate (Chan et a1 1988). If necessary, the sample is ground fiuther in a Wile mill and reformed in a cylindrical die to the native density, approximately 1 .O Y g/cm . Three holes are made radially in the pellet with a small drill and thermocouples are inserted at distances of 2, 4 and 6 mm from the heated f i e of the pellet. The pellet is then fitted tightly in a glass sleeve joined to the reactor via an 0ring seal. Real time intraparticle temperatures are displayed on the computer monitor and digitized at a rate of 1 Hz. The moisture content of the pellet is constant at a 5.0 wt %, equilibrium moisture except for some data in Chan et al(l988). We operationally define pyrolysis products as gases, liquids (pyrolysis oil or biocrude) and solids (chartunreacted sample). Gases, 13-26 samples taken at preprogrammed times, are sampled on-line then later injected automatically into the GC for analysis. The pyrolysis liquids that condense in the reactor, tar-tmp and lines are each determined quantitatively by weighing before and after the experiment. Thereafter, the pyrolysis oils are recovered by washing with a known amount of THF and analyzing each hction. The char is determined quantitatively by separating the char and the unreacted part of the pellet with a microtome. The mass balance is then closed by knowing the initial weight of the pellet, the unreacted pellet, the quantitative oil, char, and gas produced. Experimental conditions for the pyrolyses are chosen to be similar to industrial conditions in various gasification, liquefaction, and combustion reactors used with large particles or wood “chips”. These process conditions are kept constant for all pyrolysis experiments reported here. The pellet length (thermal thickness) is 1.O cm, 2 the applied heat flux is 4.0 caVcm -sec, chosen to be in the range of both pyrolysis as well as combustion and fire levels (2.8-27.4 caVcm*-sec). The reaction time of 12 min. is chosen because we wish to study the devolatilization process only. After 12 min., the temperature inside the reacting pellet is high enough to induce secondary tar reactions within the char layer that we wish to keep to a minimum. The particle surface experiences rapid pyrolysis, but the particle interior experiences slow pyrolysis, in keeping with most industrial-scale biomass thermal conversion processes. Several new techniques for pyrolysis oil analysis were tried and results are presented in Rodriguez (1996). Low resolution, or low recovery, or both owing to the oil’s complexity (more that 200 compounds are probably present) plagued all analyses.
-
1015
Capillary GC results are pertinent and this technique provided the most convenient method, though recovery was low (a mean of 7% of the oil was analyzable). The oils were injected as a whole sample (0.5 p1-100:l split) into the capillary column without derivatization. A DB-5column (J & W), temperature gradient, and automatic data analysis were used (Rodriguez, 1996). Additional methods we use appear in Lai (1991).
RESULTS SAMPLE SELECTION FOR CONTRASTING COMPOSITIONS Traditional composition data from the USDA partitioning scheme are plotted (directly) in Fig. 2a top. Most species cluster at the origin consistent with low extractives content. Species names at some extreme compositions are bold (Latin names). We have quantified in previous articles (Lai, et al, 1993) that some traditional extractives are correlated; for example ash and protein are positively correlated (1=0.26), whereas ash and polyphenol are negatively correlated (F -0.32). These correlation coefficients imply that extractives’ compositions clearly co-vary, or vary in concert, not independently. This could be seen in Fig. 2a if we rotated it to prominently show the axis along which compositions show the maximum covariance. However, we prefer to identify native species’ compositions in which we can independently contrast the e i k t of high- versus low- ash samples, or high- versus low-oil contents, on the pyrolysis product distribution. That is, we prefer orthogonal contrasts. Thus, ideal sets of species are represented in Fig 2b (with an expanded scale) by the clusters (bold dots) that show clear contrasts of composition. For example on the EWO shaded surfaces marked A and B, we have 4 clusters of species whose comparisons along the axis marked “variable 2, Z” represent 4 independent observations of the “effect of variable 2 or Z” where Z or variable 2 may not be a traditional measure of composition. Cluster analysis and other multivariate statistics methods can provide such clusters and axes. Results of our statistical analysis of the composition database are described later and in detail in Lai (1991)and Rodriguez (1996). Statistical manipulations on the USDA database (cluster analysis, principal component analysis with varimax rotation e.g., Everitt, 1980) revealed “subsets” d representative species, as idealized in Fig 2b, but with d i h n t variables (orthogonal principal components) than traditional fractions as measured by USDA. A set of species fiom each orthogonal subset appears in Table 1. The Latin names, and where available, the common names of the biomass species are given. The extractives’ ranges are: ash content, 4 to 17%; protein content, 5 to 14%; polyphenol, 3 to 11%; and oil content, 1 to 4%. However no species contains extremes of all 4 variables, Nor can species be found, retaining native compositions, at extremes of just one extractive composition, while the other fractions are present at constant levels. Thus we use orthogonal but non-intuitive compositions in this work, then rank pyrolysis effects in terms of traditional extractives content to get an understanding of their impact on biomass pyrolysis.
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Fig. 2a and b - Three of 4 “traditional” composition measures (USDA) crossplotted (above); Idealized, contrasting clusters versus “orthogonal” compositions (below)
1017
PYROLYSIS YIELDS FROM CONTRASTING COMPOSITIONS
We performed repeated single particle pyrolyses on one species to estimate the reproducibility of the entire experiment, i.e., the sampling, preparation, pyrolysis, product and data analysis. That species was a choice whose composition represented nearly the grand average of all the orthogonal composition measures -- Ongrucue Oenothera Biennis, the so-called "centerpoint" [CP] of the composition fractions. It is listed at the bottom of Table 1. The replicates, and some comparison runs, are included in Table 2 wherein only the % reacted and overall product fractions are presented. Note that the char, oil or biocrude, and gas sum to 100% within 0.6% (round off+experimental error). The replicate results, Table 2, show product yields with good reproducibility and low experimental error. The average conversion for the 12-minute one-dimensional heating is 84.6%, with standard deviation of 1.8 and a relative standard deviation of 2.1%. Note that experiment D* was carried out without extractives (the extractives Even though the were removed by using the procedure shown in Figure 1). extractives were removed, the overall yields are comparable with the results of experiments with extractives retained in the ground biomass (experiments A, B and C). The char yield average is 34.5%; the liquids or oil yield average is 47.1% with a relative standard deviation of 4.3%. These results show that on average about 47% of the total biomass converted goes into bio-oil (biocrude) which is of interest to this work. The average gas yield is 18.3% and shows the highest experimental error. Experiment E** was carried out with holocellulose (cellulose + hemicellulose). That is, the Klason lignin was removed by treating the extracted biomass with sulhric acid (TAPPI Standard Method T 222 om-82). These data were not included in the statistics calculations. The pyrolysis experiments using biomass with extractives, without extractives, and with holocellulose only demonstrate that the pyrolysis product slates !?om these starting materials are different to extents greater than the experimental error. Table 1 Biomass species chosen from the 16 orthogonal compositional subsets
[Subset #] Biomuss Species (common name-USDA) [S 11Compositae Haplopappus (Greene Blake) [S2]Leguminosae. Desmodium Cuspidatum (Tick Trefoil) [S3]Compositae Liatris Spicata (Marsh Blazing Star) [S4]Leguminosae Galactia Wrighttii (Wright Milk Pea) [SS]Polygonaceae Polystichum Munitum ( W . Sword Fern) [S6]Magnoliaceae Magnolia Acuminata (Cucumber Tree) [S7]Compositae Solidago Graminifolia (G L Goldenrod) [SSIThymelaeaceae Largerstromia Indica (Crepe Myrtle) [S9]Compositae Dyssodia Acerosa (N/A) [S lO]Ranunculaceae Rananculus Acris (Tall Buttercup) [S 1 l]Balsaminace Impatients Pallida (Yellow Jewelweed) [S 1ZIPapaveraceae Argemone Platycera (N/A) [S 131 Valerianaceae Valerianella Intermedia (Corn Salad) [S 14]Araliaceae Hereda Canariemis (Algerian Ivy) TS 1SlCarnaceae Cornus Drummondi "/A) iS 16jLabiatae Nepeta Cataria (Catnip) [CPIOngraceae Oenothera Biennis (Evening Primrose) .
I
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%ash
4.1C 5.9C 6.1C 6.9C 4.3C 4.2C 4.9C 5.8C 13.4 17.1 11.6 11.2 9.5C 10.5 9.95 11.4 7.41
%pro
%poly-
5.56 12.80 6.57 14.00 5.65 9.01 6.66 8.76 5.29 12.4 6.81 10.7 6.79 9.51 6.93 9.66 6.94
4.89 4.42 3.75 5.95 10.24 10.44 10.37 10.17 3.69 3.89 3.71 4.65 10.05 10.19 11.43 8.77 5.37
o/ooil
1.05 1.50 2.25 2.23 1.24 1.53 2.28 2.63 1.00 1.16 2.21 2.68 1.53 1.42 2.43 2.21 4.44
Table 2 Overall Product yields from species Ongraceae Oenorhera Biennis [CP]
Experiment A
%, Reacted
% Char
% Biocrude
% Gas
83.8 87.0 82.7 84.8
33.3 36.7 31.8 36.1
50.2 45.9 46.6 45.8
16.3 17.3 20.9 18.6
St. Error %Rel. St. Dev.
84.6 1.8 3.2 0.9 2.1
34.5 2.3 5.3 1.1 6.7
47.1 2.0 4.0 1.o 4.3
18.3 1.9 3.6 1.o 10.8
E* *
80.5
33.3
27.96
38.7
B C
D* Mean St. Deviation
Variance
Note: Experiment D* extractives-free and E** holocellulose only
Given in Table 3 are the total yields of pyrolysis products h m the 16 "contrasting" (orthogonal composition) samples of Table 1. The experimental pyrolysis conditions were kept constant to estimate only composition effects. The average % reacted or conversion was 71.76%. The species fiom three subsets, S7 (approximately: low ash, high protein, high polyphenol, and low oil), S9 (-low ash, low protein, low polyphenol, and high oil), and S10 (-high ash, low protein, low polyphenol, and high oil) show conversions of over SO%, while the species from two subsets, S6 (-high ash, low protein, high polyphenol, and low oil), and S16 (-high ash, high protein, high polyphenol, and high oil) show low conversions of 49 and 56% respectively. Pyrolysis oils were the major product in all cases, with an average of -48% (Table 3). The biomass species from subset S7 (-low ash, high protein, high polyphenol, and low oil) shows a char yield of 48.5% and a biocrude oil yield cf 45.2%. This was the only species that produced more char than biocrude. Char yield was the second highest product with an average of 38.4%. Gas was the lowest yield in all cases, with an average of 13.8%, although several species produced gas yields of almost 20%.
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Table 3 Overall yields from pyrolysis experiments on 16 "orthogonal" compositions (Subset) Species
(S1)Compositae Haplopappus (S2)LeguminosaeDesmodiumCuspidatum (S3)Compositae Liatris Spicata (S4)LeguminosaeGalactia Wrighttii (SS)PolygonaceaePolystichum Munitum (S6)MagnoliaceaeMagnolia Acuminata (S7)Compositae Solidago Graminvolia (S8) Thymelaeaceae Largerstromia Indica (S9)CompositaeDyssodia Acerosa (S1O)RanunculaceaeRananculus Acris (S1 1)Balsaminaceae Impatients Pallida (S12)Papaveraceae ArgemonePlatycera (S13) Valerianaceae Valerianella Interme. (S14)Araliaceae Hereda Canariensis (S15)Carnaceae Cornus Drummondi (S16)Labiatae Nepeta Cataria Minimum Maximum Mean
%
%
YO
Reacted
Biocrude
Char
Y O
Gas
79.6 1 69.90 74.53 77.27 71.91 49. 15 81.73 64.69 85.13 87.40 73.46 60.95 75.50 79.32 61.04 56.57
43.95 50.20 47.90 51.58 44.80 56.49 45.21 49.9 I 45.25 43.30 43.92 52.40 46.19 44.20 47.45 52.24
36.66 40.30 32.22 32.44 41.17 31.14 48.53 40.57 39.88 41.60 39.55 38.27 34.67 38.73 43.48 35.83
19.38 9.40 19.76 15.96 14.01 12.35 6.24 9.50 14.86 14.90 16.71 10.95 19.12 17.05 9.05 11.92
49.15 87.40 71.76
43.29 56.49 47.81
31.13 48.52 38.42
6.23 19.76 13.82
EMPIRICAL MODELS FOR PYROLYSIS YIELDS Pyrolysis product slates were analyzed using as regressors the traditional extractives as well as orthogonal principal components, i.e., quantitative variables linearly related to ash, protein, polyphenol and oil. The methods are well presented in Lai (1991) as a model for the analysis by Rodriguez (1996). Empirical models were calculated from the product slates by using linear biased regression and other appropriate methods (Ridge Regression, Principal Components Regression) to summarize the composition eEa% as coefficients or "contrasts" (Box, et al, 1978). As judged fiom Fig. 3 (representative), the agreement between the measured % reacted (circles) and the model prediction (line) is quite good. The replicates' spread joined by a horizontal line 2 shows the small experimental uncertainty. The coefficients of determination (R ) for all models are over 0.8, with F values over 20. Because all the experimental data was used in developing the models, cross-validation of the models was not carried out. The residual plots showed an even distribution of residuals without outliers. Therefore, the models were considered adequate. The traditional variables (regressors) ash, protein, polyphenol and oil, are correlated. The practical implication d correlated regressors is that the model coefficients are inflated by the extent d correlation among the variables, and thus they are ambiguous and not reported here. However, nearly the same trends were found for the fits and model coefficients based on true orthogonal compositioncontrasts (principal components) as regressors but these are not intuitive variables to the unfamiliar user. Because the principal components were not intuitive, we present qualitative rankings of effects of the traditional compositions or extractives on pyrolysis in Table 4 and the reader is referred to forthcoming publications.
-
1020
Experimental % Reacted Figure 3 Experimental vs. Predicted biomass reacted Table 4 Ranking of the traditional compositions on pyrolysis product distributions
Product Fraction % Reacted Biocrude Char Gases
First Rank*
Second*
L Protein 1'Protein same as L Oil
? Ash
t Oil
1Ash ? Ash 1Protein
Third*
Significant Interactions*? 1Polyphenol Yes Yes J Protein Yes 1Ash No
*Greater than 90% confidence.
The qualitative influence of extractive composition on the percent of the particle that reacted, as well as on the hctions of biocrude, char, and gas produced, are given in relative order of their importance, i.e., as rankings. When the extractives a~ increased over the contrasts studied, the indicated gross pyrolysis product fractions also increase (t)or decrease (4). The ranked effects are shown in Table 4 using a 90% confidence level for the significance test. For example, even increasing the minor protein content from 4 to 17% reduced the % reacted for these experiments. Increasing the ash content over the range given in Table 1 increased the % reacted. The oil content of the reactant biomass is the most influential extractive on the char and gas produced fkom pyrolyis of the species in Table 1. However, as seen fiom the arrows,
1021
high oil content reduces char but increases gas produced in the macroparticle pyrolysis. The fourth column indicates that there are significant interactions (non-additive &Us) among the extractives that alter all pyrolysis outcomes except gas production, which appears to be increased by oil in the reactant, but decreased by the presence of high protein or high ash in the biomass reactant. Recall that our experimental design has “hidden replication”. Box, et a1 (1978) discuss the advantages of hidden replication and it is schematically shown in Fig. 2b, by the shaded surfaces A and B. That is, by choosing orthogonal subsets and clusters, we contrast 4 independent (orthogonal) experiments (surface A) with the four experiments represented on surface B. These 4 contrasts give 4 replicates to estimate the effect of Variable 2. For the sixteen species in Table 1, we have eight replications of the effect of a high or low extractable component of biomass (in principal component space). Thus we can detect small but statistically significant changes in pyrolysis slate due to native biomass composition differences. Intraparticle temperature plays a key role in the pyrolysis process. In the single particle reactor the “thermally thick” biomass pellet is heated on one face by the atr lamp at the applied heat flux. Although not shown here, the measured temperature inside the pellet rises at distances of 2, 4 and 6 mm fiom the heated face. The temperature profiles observed in this work are variable fiom species to species, and even within the same species (OenotheruBiennis), owing to natural sample variability even though we homogenize the sample and the experimental conditions are identical. The temperature rises during the first 6-8 minutes of the experiment until it eventually “plateaus” as long as devolatilization is still occurring deeper in the pellet. At 2 mm inside, the temperature plateau is about 600” C, with the deeper temperature plateaus being about 380‘ at 4 mm and 150’ at 6 mm. During the first 2-4 minutes at temperatures around 200-300’ C, the pyrolysis process is most intense. This is confumed by visual observation of the evolution of volatiles from the heated lice of the pellet. After 4 minutes, the pyrolysis reactions are still taking place but the evolution of volatiles from the pellet interior diminishes in intensity. Thus, our results, while representative of conditions experienced by large biomass particles in industrial reactors, cannot be categorized as either slow or fast pyrolysis, but rather as a pyrolysis dominated by large gradients indicative of important heat and mass transfer rate limitations.
CONCLUSIONS AND SIGNIFICANCE We have answered the question posed in the title of this paper, at least for a wellchosen subset of underutilized biomass residues. Our method of studying herbaceous and woody biomass composition effixts on pyrolysis product slates is generally applicable and we again stress the importance of using statistically designed experiments, especially for correlated variables. We gratefully recognize that for the feedstock composition data and samples, we are fortunate to have the cooperative agreement with the USDA and our conclusions rest on their measurements and sample collections, and complement their fmdings in the cited papers. Other biomass sampling protocols and laboratory composition analyses might uncover different members of the subsets, or different pyrolysis product slate results for the same conditions. However, this paper is intended to focus on the question raised in the title. We stress the power of the statistics method and its ability to detect small differences within the range of natural sample VariabiIity. We also recognize that other contributions to the product slates may be significant, such as the 60-85% of the sample that is cellulose and lignin. As found by MacKay and Roberts, these two tiactions strongly influence char yield. However, by the use of orthogonal variables
1022
and subset choices, our findings on the extractives' influence on pyrolysis yields should be valid in spite of this. The importance of orthogonal or uncorrelated independent variables is further illustrated in Lai (1991) and forthcoming publications. ACKNOWLEDGMENTS
The authors wish to thank NREL and the USDA for their support and cooperation. Additionally, BKB acknowledges the inspiration and advice of MKB regarding the analysis and reviewers for their helpful suggestions. CITED LITERATURE
Adams, R.P., M.F. Balandrin, K.J. Brown, G.A. Stone, S.M. Gruel (1986) Extraction of Liquid Fuels and Chemicals fiom Terrestrial Higher Plants, I. Yields fiom a Survey of 614 Western U.S. Plant Taxa, Biomass 9 255-292. Buchanan, R.A., F.H. Otey, G.E. Hamerstrand (1980a) Multi-Use Botanochemical Crops, an Economic Analysis and Feasibility Study, I&EC Prod. Res. and Devel., 29(4) 480-496. Agblevor, F.A., B. Rejai, D. Wang, A. Wiselogel, and H.L.Chum (1992) Thermochemical Conversion of Biomass to Fuels and Chemicals, the Effect CC Storage Conditions on Pyrolysis Products, Proceedings Alternative Energv-Liquid Fuelsfiom Renewable Resources, Dec. 13-15, 1992, Nashville, TN. Agblevor, F.A., S. Besler, and A.E. Wiselogel (1995) Fast Pyrolysis of Stored Biomass Feedstocks, Energy and Fuels, 9, 635-640. Antal, Michael J., Jr. (1985) Biomass Pyrolysis: A Review of the Literature: Part 2 Lignocellulose Pyrolysis, In: Advances in Solar Energy, Karl W. Boer and John A. M e Eds, 2 175-255.: American Solar Energy Society, Inc., Plenum Press,. Boulder, CO. Bagby, M.O., R.A. Buchanan, and F.H. Otey (1980) Multi-Use Crops and Botanochemical Production, ACS Symp. Ser. 144, Biomass as a NonFossil Fuel Source, D.L. Klass, Ed., 125-136. Bagby, M.O.( 1989) personal communication and cooperative agreement with UW & Northern Regional Research Center of the USDA. Box, GEP, J.S. Hunter, W.G. Hunter, Statistics for Experimenters, Wiley, 1978. Bridgwater, A.V., and M.L. Cottam (1992) Opportunities for Biomass Pyrolysis Liquids Production and Upgrading, Energv and Fuels 6(2), 113-120. Can, M.E., M.O. Bagby, and W.B. Roth (1986) High Oil- and PolyphenolProducing Species of the Northwest, JAOCS 63( 1 l), 1460-1464. Chan, W.C.R, M. Kelbon, and B. Krieger-Brockett (1988) Single Particle Biomass Pyrolysis: Correlations of Reaction Products with Process Conditions, Ind. Eng. Chem. Res. 27 2261-2275. Davis, E.A., J.L. Kuester, and M.O. Bagby (1984). Biomass Conversion to Liquid Fuels: Potential of Some Arizona Chaparral Brush and Tree Species, Nature 307(5953), p 726-728. Elliot, DC, LJ Sealock, and RS Burner (1988) Product Analysis fiom Direct Liquefaction of Several High-Moisture Biomass Feedstocks in Pyrolysis Oils from Biomass: Producing, Analyzing, and Upgrading, Eds: Soltes, EJ, and Milne, T.A., ACS Symp. Seri 376, p 55-65,. 179-187. Evans, R.J., TA Milne (1987) Molecular Characterization of the Pyrolysis of Biomass. 1., Fundamentals, Energy and Fuels 1 (3), 123-137. Evans, R.J., TA Milne (1987) "Molecular Characterization of the Pyrolysis of Biomass. 2., Applications, Energy and Fuels 1 (4), 3 11-319.
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Everitt, B.S. (1980) Cluster Analysis 2nd Ed, Heineman Ed. Books, Ltd.,. London. Hajaligol, MR J.B. Howard, JP Longwell, WA Peters (1982) Product Compositions and Kinetics for Rapid Pyrolysis of Cellulose, Znd.Eng Chem. Res., 21, 457-465. B. Krieger-Brockett, W.C. Lai and W.C. Chan (1997) Comparisons in Biomass and Refhe Derived Fuel Pyrolysis.Multivariate Analysis applied to an Experimental Design. Developments in Biomass Thermochemical Biomass Research, Elsevier, p. 43-61. Lai, WC (1991) Ph.D. Thesis, Reaction Engineering of Heterogeneous Feeds: Municipal Solid Waste as a Model, Univ. of Washington, Dept of Chemical Engineering Lai, WC, I. Rodriguez, and B. Krieger-Brockett (1993) Composition Effects on the Devolatilization Behavior of Biomass and Municipal Solid Waste, Advances in Thermochemical Biomass Conversion, Blackie Academic Publishers, London, pp. 818-832. MacKay, DM and PV Roberts (1982) The Dependence of Char and Carbon Yield on Lignocellulose Precurors Composition, Carbon, 20 95- 104. Nunn, TR., J B. Howard, JP. Longwell, and WA. Peters (1985a) Product Compositions and Kinetics in the Rapid Pyrolysis of Sweet Gum Hardwood, Znd. Eng. Chem. Process Des. Dev., 24, 836-844. NUM, TR., J B. Howard, JP. Longwell, and WA. Peters(l985b) Product Compositions and Kinetics in the Rapid Pyrolysis of Milled Wood Lignin, Ind. Eng. Chem. Process Des. Dev., 24, 844-852. Richards, GN and G. Zheng (1991) Influence of Metal Ions and Salts on Products h m Pyrolysis of Woods: Applications to Thermochemical processing of newsprint and Biomass, J. Anal. and Appl. Pyrolysis 21 (1-2) 133-146. Rodriguez, I. (1996) Ph.D. Thesis Composition Related Effects of Thermal Reactivity of Organic Feedstocks, Univ. of Washington, Dept of Chemical Engineering. Roth, WB, ME Carr, EA Davis, and MO Bagby (1984) Evaluation of 107 Legumes for Renewables Sources of Energy,. Economic Botany 38(3) 358-364. Somasundaram, S, (1990) MS Thesis, Biomass Pyrolysis: Temperature Distribution and Composition Effects, Univ. of Washington, Dept of Chemical Engineering. Swanson, C.L., R.A. Buchanan, and F.H. Otey (1979) Molecular Weights of Natural Rubbers from Selected Temperature Zone Plants, J. Appl Poly. Sci. 23 743-748. Williams, PT and PA Home (1995) Analysis of Aromatic Hydrocarbons on Pyrolytic Oil Derived from Biomass, J. Anal. andAppl. Pyrolysis 31 15-37.
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Biomass Selection Criteria Conversion Processes
for
Pyrolytic
Anuradda Ganesh and Raveendran K. Energy Systems Engineering, Indian Institute of Technology, Bombay 400 076, India
ABSTXACT: Composition of biomass is proved to play a definitive role in determining the pyrolysis product distribution and their properties. In this paper, an attempt is made to suggest criterion for selecting biomass to obtain optimum products such as char, liquids, gases and char adsorbent through pyrolysis process. In order to facilitate the implementation of such a guideline, Rating Indices (RI)are defined based on the product yield and their properties which are specific to each of the end use, namely, pyrolysis-, carbonization, liquefaction and gasification. Based on the RI obtained for a particular application, different biomass are ranked for their suitability for each of the conversion process. This ranlung is specifiho endTuseand is different for different application. Further, a stepwise procedure is presented to obtain the RI of any biomass from their initial composition.
INTRODUCTION Biomass pyrolysis though is one of the first process man developed, is being studied extensively since the last two decades to obtain liquid, gaseous and solid fuels and chemicals. It is well recognised that pyrolysis plays a key role in any of the thennochemical conversion process be it combustion, gasification, liquefaction, production of char or active carbon. In this context, selection of feedstock and optimal utilisation of the products can play a vital role [l]. Thls paper suggests a criteria to select an appropriate biomass or find its relative suitability to the conversion processes. Biomass is composed of various components such as cellulose, hemicellulose, lignin, extractives and mineral water. The composition of biomass plays a definitive role in altering the product distribution and their properties [2-31. As is shown in earlier publications [4-81 different biomass, on pyrolysis, give different product yield with different product properties. In order to choose a biomass for a particular process (carbonisation, liquefaction, gasification or adsorbent char) knowledge on the product distribution and properties for various biomass are essential.
1025
Literature does not provide information on (i) suitability of a particular biomass for a particular process or vice-versa, or (ii) a methodology to evaluate the suitability of a biomass or a process. To address the above mentioned "Suitability criteria" this paper introduces the concept of "Rating Indices", which is based on the desirable properties of the products for a particular application. Further, it is also possible to obtain the suitability of biomass lmowing only the initial composition of biomass. For this purpose various correlations developed and reported in previous publications [4-81 are employed in order to obtain the required product yield and their properties so as to enable rating of a biomass for a process. RATING INDEX (RI)
DEFINITION Rating Index (RI) has been defined separately for each of the four processes i.e. carbonisation, liquefaction and gasification and char adsorption (activation), based on the product yield and their properties as follows: Carbonisation
Rating index for carbonisation is defined as the product of char yield, char heating value and total fraction of carbon present in the char, where all values are on asmeasured basis which is shown in equation (1). RIc = Yc x Hc x Cc
(1)
where RIc is rating index for carbonisation, Yc is char yield (as fraction), Hc is higher heating value of char (in MJkg) and Cc is carbon content (percentage weight) of char. Liquefaction
Rating index for liquefaction is given in equation (2) which is defined as product of liquid yield and higher heating value of pyrolysis liquids divided by mole of hydrogen required per mole of carbon in biomass liquids for upgrading to premium kel.
RIC = Ye xHC x 1/Hr
(2)
where RIC is rating index Rating Index for liquefaction, Y.! is pyrolysis liquids yields (in fraction), He is higher heating value of pyrolysis liquids (in MJkg) and Hr is hydrogen requirement for upgradation of premium fuel. More details of Hr are given in section 2.2. Gasification
Rating Index for gasification is defined as product of gas yield, heating value of the gases and the WO ratio of the gases and the same is represented in equation (3). 1026
RIg = Yg x Hg x WOg
(3)
where RIg is rating index for gas, Yg is gas yield (in fraction), Hg is hgher heating value of pyrolysis gases (in MJ/nm3)and WOg is ratio of hydrogen to oxygen of gases (by percentage weight). In addition to the above three, rating index is also defined for the char to be used for adsorption purpose i.e. as activated charcoal. Char Adsorption Rating index for making char adsorbents is defined as product of char yield and char iodine adsorption value (taken in percentage weight on as-measured basis). RIca = Yc x Ic
(4)
where RIca is rating index for utilisation of char as carbon adsorbent and Ic is the percentage of char iodine adsorption on as-measured basis. WHAT IS Hr?
Figure 1 shows the Van Krevelan diagram for liquefaction presented by Chronet and Overend [S]. The atomic WC and O/C ratios of petroleum and also that of biomass pyrolysis liquids are indicated in the same figure. It is observed from the figure that the biomass pyrolysis liquids are characterised by very h g h O/C ratios (more than that of even the original biomass itself). However, biomass pyrolysis liquids have H/C ratios comparable to those of petroleum products. It is important to note fiom Van Krevelan diagram that for the biomass liquids, upgradation possibility rests as a h c t i o n of the methods of oxygen removal so as to achieve products with low O/C ratios which are compatible with the existing highly capital - intensive oil based transportation. For a carbon-rich, and oxygen-, hydrogenpoor fuel such as coal or coal liquids the problem is to add hydrogen on, rather than to remove oxygen. The stoichiometric hydrogen requirement to upgrade biomass pyrolysis liquids to premium fuels has been calculated through the following relation presented by Chronet and Overend [S]: mole of H required Hr = mole of C in feed = 2AR [.product
-
reaction
1
where R is the number of oxygen molecules required to combust completely one molecule of biomass pyrolysis liquids to water and carbon dioxide.
1027
0.0
FIC.1
0.3
0.2
0.1 0.5
-
COMPONENTS
0.6
RATIO
0.8
0.9
BlOUAS L 1 4 U I D S
0.7
DE-ASUaD
~IOMAU/COM~ONLNTS-LI~UIDI DlOMASS/ COMPONENTS-CUR
OIC-ATOMIC
A
m-
0
9-BIOMASS I
1.0
1.1
1.2
1*3
1.4
COMBUSTION O f ONE MOLECULE OF BIOMASS/ LlOUlDS /CHAR XYL
REQLURED FOR COUPL ETE
R SNUUBLR of bXY6EN MOLECULES
VAN K R E V E L E N O l A G R A M FOR LIQUEFACTION
0.4
LEGNlN
WD-WOOD
-
RW- RICE HUSK
LN
NORMALISED RATING INDEX (NH) The Rating Indices obtained from all the biomass studied are normalised to obtain values (NRI) ranging from 0-1 for the purpose of comparison using the following relation.
NRI = [RIb-RImin]/ [RImax-RImin]
(6)
where RIb is Rating Index for a given biomass, RImin is the lowest value of rating index obtained, and RImax is the highest value of rating index obtained among the biomass studied. DISCUSSIONS The calculated Rating Indices are tabulated in Table 1. It can be seen from Table 1 that lignin has the hghest RI for carbonisation process as against xylan whch has low carbonisation index. For liquefaction both cellulose and lignin show similar index. As expected, xylan has the highest RI for gasification process and for makmg char adsorbents. Among the five biomass studied in Packed Bed Pyrolysis (3,4,5), coir pith and groundnut shell have the highest rating index for carbonisation which is about 4.5 whereas groundnut shell and rice husk have the highest rating index for liquefaction having RI values of 2.63 and 2.34, respectively. On the other hand, wood has the highest rating index for gasification (0.70) and rice husk has the hghest rating index (16.4) for making char adsorbents. De-ashed biomass are suitable for liquefaction and for making char adsorbents. Considering liquefaction, groundnut shell shows the highest rating (2.63) followed by corn cob (1.65). Both cellulose and lignin shows higher rating for liquefaction (about 0.94). Under de-ashed condition also, groundnut shell and coir pith have higher RI for carbonisation. Wood has a higher RI for gasification, however on de-ashing shows a hgher RI for carbonisation. As may be seen from Table 1, groundnut shell, coir pith and corn cob are very attractive feedstock for making carbon adsorbents.
STEP WISE PROCEDURE TO CALCULATE RATING INDICES
In order to calculate Rating Indices, the data required are those of char yield, liquid yield, gas yield and that of respective heating values and elemental composition of each hydrogen requirement for upgradation of liquids is necessary. Also, to calculate RI for char adsorbent making, char iodine adsorption value is required. Correlations have been developed to obtain each one of the above candidate values, as functions of the biomass compositions and are detailed in previous publications [4-71. The correlations used for this purpose are presented here for the sake of completeness. Based on these, the stepwise procedure to calculate the different rating indices can be laid down as follows:
1029
Table I Rating Indices of biomass for char, liquid and gas production. Biomass/ Components Components Cellulose Lignin Xylan Biomass Coir Pith Corn Cob G'nut Shell Rice Husk Wood De-ashed Biomass Coir Pith Corn Cob G'nut Shell Rice Husk Wood
Ratings RIC
RI&
RIfz
RIca
3.22 11.95 5.13
0.96 0.94 0.17
0.03 0.18 1.04
3.8 3.7 10.2
4.62 4.21 4.39 1.70 3.61
0.7 1 1.65 2.63 2.34 0.85
0.09 0.03 0.08 0.22 0.70
9.9 5.3 5.7 16.4 2.3
5.63 2.49 6.34 1.77 2.77
1.52 2.66 2.23 2.32 2.42
0.02 0.10 0.07 0.22 0.20
13.4 8.6 11.6 14.9 2.3
(i) determine chemical composition of biomass, particularly, cellulose, lignin, hemicellulose, ash and silica content, (ii) calculate the product yield using the equation (7),
where Yb is the product yield and suffix i takes the values for char, liquids and gases; X is a percentage of each component suffixes c, 1, a, si represents cellulose, lignin, ash and silica respectively values of constants c and M are given below: Product
c 3
Ml
M Z
M 3
char liquids gases
49.90 102.86 13.93
-0.0258 0.7028 -0.4 168
0.258 0.335 -0.337
-0.044 0.033 -0.002
~~
(iii) calculate the heating values of products using the equations (8), (9) and (10). 0.19275 -0.18165 -0.134073;sE =0,063andR2 = o.984 H , =34.0138X1 Xu xsi
1030
(8)
where H,, HI and H, are higher heating values of biomass pyrolysis char, liquids and gases respectively and SE is standard error value and R2is R2 value. (iv) calculate the elemental composition of the products using the equations (1 1) & (12) H -0.96 Xc-2.533x1.7493.&a-0 6143 ;!&f6l08andR2 =0.988
i.37
-(Ii/C\g
30.2334X;
Xa
XSj
where WC and O K are percent weight ratios in percentages and suffixes c, 1 and g -0.486 -4.4284 4.962 0.388 (0IC)g =x c xi Xa Xsj represents char, liquids and gases respectively.
(v) calculate the carbon content of char using equation (13) -0.3350 -0.1612 -1.0225 C , = 0.4433 % Xc Xi Xa where C, is the carbon content in the char.
(13)
(vi) calculate the hydrogen required for upgrading the liquids to premium hels from van Krevelen’s diagram (Fiqure 1 and equation (5)) (vii) calculate the char iodine adsorption values using the equation (14), and
r = 49.93 x (yVx R,x xsi)o.2m1
(14)
where I is iodine adsorption, Y , is volatile yield in fraction, & is maximum rate of volatilisation, XSiis silica content, (viii) use equations (1 5) and (1 6) to generate the weight loss curve biomass to obtain volatile yield and maximum rate as de-volatilisation.
are the experimental weight I derivative weight loss data for any given and biomass and X, , X,, and XIare initial silica fiee ash, silica and lignin content, 1031
respectively, in fraction, present in dry biomass on weight basis and CI = 0.5, nl = 8.5 and n2 = 7.0 Substituting the values of product yield and their corresponding product properties in equation (1) through (4) rating indices can be calculated. BIOMASS CHARACTERISATIONBASED ON NRI
Using the aforesaid procedure, the RIs for the other biomass (whose product properties are not experimentally studied) have been calculated. The RI values obtained are normalised to obtain NRI as per equation 5 and are presented in Table 2. The values of the maximum rate of devolatilisation is obtained from the actual TGA curves. (These may also be calculated with a high level of accuracy using the chemical analysis of biomass as described in steps (ii) and (iii) in Section 3). The last column indicates the remarks regarding their utilisation or end-product based on the ranking systems suggested. It may be seen from Table 2 that for the purpose of charcoal production, coir pith, groundnut shell, corncob and cashewnut shell are better suited candidates than the other biomass considered for the present study. Similarly, in the case of liquid he1 production, coconut shell, cotton gin waste, groundnut shell, coconut coir and rice husk show better properties. Coconut shell, groundnut shell, wood, coconut coir and rice husk are better suited candidates compared to the other biomass studied for pyrolytic gasification applications. For active carbon manufacture, rice husk, corn stalks, wheat straw, coir pith and cotton gin waste are better rated candidates. The method, presented here is the initial step in choosing or selecting a biomass from various available sources. However, fbriher studies have to be conducted for optimising process parameters and also to obtain the suitability in terms of physical properties (e.g. solid flow properties of biomass).
CONCLUSIONS
Criteria for evaluating the suitability of biomass for a pyrolysis conversion process to obtain solid, liquid and gaseous hels are developed based on the properties of the pyrolysis products. Based on this method, suitability of various biomass for pyrolytic conversion is studied and is found that based on the composition of biomass, different biomass are suited for different applications such as carbonisation, liquefaction, gasification and making of char adsorbent.
1032
Table 2 Ranking Biomass for Pyrolysis Conversion Processes. Biomass Bagasse Cashewnut Shell Coconut Coir Cococnut shell Coir Pith Corn cob Corn Stalks Cotton gin waste Groundnut shell Millet husk Rice Husk Rice Straw Wood Wheat Straw
Char A B .21 7 .22 4 .22 5 .20 8 .31 1 .27 3 .13 9 .03 11 .29 2 .03 12 .04 10 .OO 14 .22 6 .05 13
Liquids A B .32 11 SO 6 -62 4 1.oo 1 .15 14 .46 7 .37 9 2 .92 .79 3 .33 10 5 .69 .39 8 .19 12 .17 13
A - NRI for each of the product B- Rank for biomass for each product C- Carbonisation
Gases A B 6 .13 2 .80 4 .38 1 1.00 7 .09 9 .02 10 .01 14 -04 8 -07 .01 11 .20 5 12 .01 3 .67 .OO 13
Adsorbent A B .44 6 .14 12 .19 10 13 .04 4 .54 9 .22 .58 2 .48 5 .24 8 .30 7 1.oo 1 .16 11 .oo 14 .55 3
Remarks good feedfor
C&G L&G L&G c c CA L&CA C&L CA
-
G GA
-
G Gasification L Liquefaction CA - Carbon Adsorbent
REFERENCES 1. Antal Jr., M. J., "Biomass Pyrolysis: A Review of the Literature, Part I : Carbohydrate Pyrolysis (vol. 1, p175-239, 1982); Part 11: Lignocellulose Pyrolysis (vol. 2, p 1-96, 1985), In Advances in Solar Energy, Eds. K. W. Boer and J. A. Duffie, American Solar Energy Society, New York. 2. Ganesh, A., "Studies on Characterisation of Biomass for gasification", Ph. D. Thesis, 1990, Department of Chemical Engineering, Indian Institute of Technology, New Delhi, India. 3. Raveendran, K. "Studies on Influence of Biomass Composition on Pyrolysis", Ph. D. Thesis, 1995, Interdisciplinary Program in Energy Systems Engineering, Indian Institute of Technology, Bombay, India. 4. Raveendran, K., Ganesh, A. and Khilar K. C., "Pyrolysis Characteristics of Biomass and Biomass Components", Fuel, v-75, n-8, pp.987-997, 1996. 5. Raveendran, K., Ganesh, A. and Khilar K. C., "Influence of Mineral Matter on Biomass Pyrolysis Characteristics", Fuel, v-74, n-12, pp.1812-1821, 1995. 6. Raveendran, K. and Ganesh, A. "Adsorption Characteristics and Pore Development of Biomass - Pyrolysis Char", Fuel, v-77, n-7, pp.769-781, 1998. 7. Raveendran, K. and Ganesh, A. "Heating Value of Biomass and Biomass Pyrolysis Products", Fuel, v-75, n-15, pp.1715-1720, 1996. 8. Chronet, E. and Overend, R. P., "Liquid Fuels from Biomass", In Biomass Regenerable Energy, (Eds. D. 0. Hall and R. P. Overend), John Willey and Sons (1987).
1033
Use of a Concentrated Radiation for the of Cellulose Thermal Determination Decomposition Mechanisms 0.Boutin and J . Lkdk Laboratoire des Sciences du Gknie Chimique LSGC-CNRS-ENSICINPL, 1, rue Grandville - BP 451 - 54001 Nancy Cedex France
ABSTRACT The purpose of h s paper is at first to summarize the main experimental results obtained with an original device (image furnace) in which pellets of pressed cellulose are subjected to a concentrated radiation. The time of pyrolysis and also the absorbed heat flux densities are perfectly controlled. The fractions and chemical compositions of all the products are determined as a tinction of the reaction time. It is hence possible to distinguish the primary products from those resulting from secondary reactions. The results bring new evidences for the validity of the Broido ShafEadeh model at atmospheric pressure as well as indications about the locations where the gases, liquids and solids are produced. It is then suggested to take advantage of these results for studying more usual pyrolysis reactors, on the basis of the heat flux densities that are locally exchanged. It is for example possible to predict the operating conditions in which no char is formed and hence almost only water soluble products are obtained.
INTRODUCTION A great number of papers have been and continue to be published in the field of biomass thermochemical conversion. They are related to:
0
fundamental studies (kinetic schemes, Arrhenius and thermodynamic constants [1 -61) the optimization of reactors (cyclone, fluidized bed, rotating cone, ... [7-121) the modelling of the phenomena (chemical, transfer and hydrodynamic processes) occuring at the particle and at the reactor levels [13-181.
All the processes of biomass thermal conversion (pyrolysis, gasification and combustion) begin with elementary steps of decompositions of each of the components of the starting material (cellulose, hemicellulose and lignin). It is hence necessary to well understand the kinetics of the corresponding fast elementary
1034
reactions of pyrolysis. At the present time, no consensus has been reached, even for a model component such as cellulose. Most of the authors consider several kinds of competitive steps giving rise to gases, tar and char followed by subsequent consecutive reactions. The most often mentioned schemes rely on the BroidoShafizadeh model [l]. However, according to some authors, a step giving rise to intermediate species (sometimes called "active cellulose") is not necessary for the prediction of the global kinetics of cellulose pyrolysis [19]. Other authors completely deny the existence of such species and just consider a single step giving rise to vapours from virgin cellulose [20-211. Recently, LCdt et al. [4]gathered a great number of strong further evidences for the existence of these intermediate species. Several reasons may explain these controversies. Indeed, the studies are usually performed in systems (mainly thermogravimetricanalysis) having several drawbacks : they are incompatible with the rapidity of the reactions. they are unable to evidence reactions occuring without mass loss. the heat and mass transfer efficiencies are often poor, and hence the pure chemical processes are not rate controlling. secondary reactions (crackings, recondensations ...) are possible and do not permit studying the initial short life time species. The main purpose of the present paper is to report several new (qualitative and quantitative) results obtained in an original set up, in order to better understand the mechanisms of cellulose pyrolysis. EXPERIMENTAL, METHODS
A description of these methods has been given by Boutin [22] and we shall merely remind the basic principles of the experiments. A small cylinder of cellulose (obtained from pressed cellulose powder) is placed in the middle of a quartz reactor fed at steady state by a flow of argon at room temperature. The end cross section of the pellet is positioned in the focal volume of an image furnace. Such a device relies on the use of a 5 kW xenon high pressure arc lamp associated with two elliptical mirors. The available flux density obtained in the focal volume and on the surface of the pellet can be quantitatively changed through the interposition of diaphragms [ 6 ] . The values of the absorbed heat flux densities used (2.1 x lo5 Iq (W m-2)I 7.4x lo6) are known through calibration measurements. The cellulose sample is subjected to the concentrated radiation during perfectly controlled flash times ( 5 x lo-*I t(s) I 5). After the flash, the mass loss of the pellet is measured. The fractions of all the products (remaining on the pellet and escaping from the reactor) are determined as well as their chemical composition (HPLC, HPLCMS, GC and elementary analysis). The reprodwAbility of the experiments and the mass balances is very good: around 98 %. NATURE AND CONDITIONS OF FORMATION OF THE PRODUCTS
The products are recovered at the surface of the pellet and downstream of the reactor (condensed vapours and gases).
1035
PRODUCTS RECOVERED AT THE SURFACE OF THE PELLET
After the reaction, the surface of the pellet is covered with a solid brown matter. Its microscopic observation reveals that the fibrillar structure of virgin cellulose (figure 1) has disappeared and that the products have the appearance of large networks of more or less agglomerated melted particles (figure 2). These observations clearly show that, in our conditions, the thermal decomposition of cellulose passes through an intermediate stage involving liquid compounds (ILC). However, although ILC appears to be liquid at the temperature of reaction it is solid at room temperature. It is soluble in water proving that it is not melted cellulose. From the temperature of glass transition of cellulose (510 K) it is possible to calculate a theoretical melting temperature of cellulose close to 720 K [23]. This temperature is very close to its pyrolysis temperature (around 740 K) and hence, even if cellulose melts, this possible liquid state would have a too short life time for being detectable in our device. ILC could be compared to the “active cellulose” postulated more than 20 years ago by Bradbury et al. [l] even if these authors had few chances to observe a liquid phase because their experiments were made at too low temperatures. Direct and non destructive analysis (HPLC, and HPLC coupled with MS) shows that ILC contains anhydro-oligosaccharides with degrees of polymerization ranging fiom 7 to 2 and levoglucosan. These species result from the depolymerization of the cellulose molecules. It is important to note that the chemical composition of ILC does not depend on the operating conditions (heat flux densities and flash times).
Fig.1 SEM photograph of the surface of a cellulose pellet before the reaction (x 800)
1036
Fig.2 SEM photograph of the surface of a cellulose pellet after a flash of concentrated radiation (x 1000) Studied as a function of the flash times of pyrolysis, the thickness of the layer of ILC first increases until it reaches a constant value ranging from 100 and 200 pm (steady state conditions of pyrolysis). The mass loss from the pellet that begins just after the appearance of ILC, linearly increases as soon as the extent of the layer of ILC has reached the steady state. CHAR For absorbed heat flux densities lower than approximately 9 x lo5 W m-*, the layer of ILC becomes progressively covered with amorphous black products (char) that are not soluble in water. The quantity of char increases with the flash time. Its mass fraction may range between 0 and 10 % according to the absorbed heat flux densities. The elementary analysis of char gives the following mean molar composition : C : 6 ; H : 7.6 ; 0 : 3.3, indicating that the char is far from being pure carbon. The image furnace provides a given flux density under the form of a concentrated radiation and not of a given temperature. Actually the effective temperature of the products formed inside the focal zone depends on a great number of chemical and physical (ex. optical) characteristics. It is hence very difficult, without the help of an accurate mathematical model, to estimate, a priori, the temperature reached by the char. CONDENSED VAPOURS
The condensed matter (including water) trapped by the filters placed downstream of the reactor contains qualitatively the same species as those previously analysed in ILC. The presence of heavy molecules can be explained by mechanical ejections from the surface of the pellet (in the probable form of aerosols). We can expect that the highest fractions of levoglucosan and cellobiosan result from their vaporization as 1037
soon as they are fonned from secondarycrackings of the heaviest molecules inside the layer of ILC. The quantity of the products trapped by the filters linearly increases with the flash time after the beginning of the steady state conditions of the pellet pyrolysis. The total mass fraction of these condensedvapours may reach 80 %. GASES
A non negligible hction of products (up to 10 %) is recovered in the form of gases. Their composition is given on a dry basis by the tables 1 and 2. The results do not depend on the flash time for a given heat flux density, but seem8 to change with this flux.No noticeable quantity of COz is detecied. Table I Molar composition (%) of the gases for several absorbed heat flux densities (W m-’) NoChar Absorbed flux 4.0 x lo6 co 41.8 25.8 H2 67.6 Hz + CO 15.5 Methane 1.5 Propylene 2.2 Ethane 8.7 Ethylene 1.1 n pentane Propane 1.7 Isobutane 1.4 0.3 n butane
formation 3.4 x lo6 48.4 17.0 65.4 16.1 0.5 1.1 8.2 7.3 0.4 0.7 0.3
Char formation 2.7 x lo6 50.5 22.5 73.O 13.2 1.4 2.0 6.5
1.1
1.1 0.9 0.8
2.2 x I d 51.9 31.3 83.2
4.6 x 10s
63.5 21.3 84.8 9.7 1.7 1.7 0.8 0.3 0.1 0.5
10.3
1.3 3.O 1.2 0.2 0.4
0.3 0.1
0.4
Table 2 Mass composition (%) of the gases for several absorbed heat flux densities (Wm-’)
NoChar Absorbed flux 4.0 x lo6 Hz + CO 55.0 Hydrocarbons 45.0
formation 3.4 x lo6 54.0 46.0
char 2.7 x lo6 62.0 38.0
4.6 x 10’ 82.0 118.0
folmatim 2.2 x lo5 79.0 21.0
when no char is formed, the mass fiactions of the hydrocarbons may represent almost 50 % of all the gases. When char is formed, the molar fraction of H2+ CO reaches 85 %. It is however difficult to determine if such behaviour is due to the differences in the values of the flux densities or to the presence or not of char, the two phenomena being connected each others.
1038
Table 3 gives different results of the values of CO / H2 and H2 + CO (on molar basis) reported by the literature. They are obtained in very different experimental conditions, and their comparison with those presented in this paper is difficult. They reveal that the ratio CO / H2 may vary between 1.5 and 37 even if the total fraction CO + H2 is relatively constant (roughly between 60 and 80 %). However we can observe that the results of the present work are similar to those obtained by Antal[24] and also in cyclone reactors. In each case, the vapour phase formed near the hot walls and / or near the reacting particles is rapidly mixed with a colder carrier gas (quench). Another interesting observation is that in our image h a c e experiments, the gases are detected only after the beginning of the steady state conditions of pyrolysis of the pellet.
Table 3 Molar composition of the gases as reported in the literature for several fast pyrolysis processes References Antal [24] Li [25] LCdt et al. [7] Diebold et al. [26] Graham et al. [27] Peacocke [28] Boutin et al. [29] This study (no char) This study (char)
Process Solar furnace Spinning disc Cyclone Vortex reactor Vortactor Ablation Rotating cylinder Image furnace Image furnace
T (K) 973 873 893 900 700 824 973
-
CO / H2 2.9 12.5 1.5
2.9 37 14.2 15.5 2.1 2.2
(CO + HJ~.(%) . . 70 81 76 70 76 61 66 69 84
OVERALL INTERPRETATIONS All these observations show that after a period of simple heating, cellulose primarily decomposes into intermediate and short life time species (ILC) resulting from partial thermal depolymerization processes. After a short transient period these species remaining on the sample reach a constant thickness. This steady state regime indicates the existence of an equilibrium between the rates of ILC formation and disappearance (through probable ejection of aerosols and vaporizations). It is difficult to know the exact nature (i.e. degree of polymerization) of the primary molecules formed at the first stages of cellulose decomposition. For example, we cannot d e f ~ t e l yconclude if the analysed species having DP ranging from 3 to 7 are of primary nature or not. However, it is likely that the lighter products such as levoglucosan and cellobiosan (analysed on the pellet and in the filters) are formed from secondary cracking, occurring inside the ILC layer. These products rapidly vaporize after their formation. Hence, we can expect that the lighter molecules, trapped by the filters, are not of primary natures as often stated in the literature. Gases and char are also formed but only after the fKst appearance of ILC indicating that they also result fiom secondary reactions of the vapours and of ILC.
1039
Above flux densities of about 9 x lo5 W m-2,no char is formed indicating that under these circumstances the pyrolysis of cellulose gives rise to about 90 % of water soluble products (ILC and vapours) and about 10 % of gases. The late formation of gases indicates that, under our conditions, no "prompt gas" would be formed as sometimes suggested in the literature [27]. The large hydrocarbons and H2 contents of the gases, as well as the absence of C02, bring another proof of the secondary nature of the gases (formed by thermal cracking reactions). All these results provide evidence for the validity of the Broido Shafizadeh type model for representing the elementary processes of cellulose pyrolysis (Fig. 3).
Fig. 3 Simplified cellulose thermal decomposition pathways at atmospheric pressure
The question of the mechanisms and locations of gas formation (reaction 4) deserves to be raised. In a recent bibliographical review, LBdd [S] gathers several data related to the kinetics of the thermal cracking of vapours formed during cellulose and biomass pyrolysis. These values can be associated to the short (around 0.5 s) residence time and low (around 500 K) theoretical mixing temperature of the cold carrier gas and hot vapours flowing through the quartz reactor in the present work. The calculations show that, in our conditions, these vapours have only few chances to thermally crack during their crossing of the reactor (quench effect). The cracking of the vapours may, a priori, occur inside or downstream of the layer of ILC. If we expect that the coalescence of the small droplets forming ILC is not perfect, this layer has a porous structure (as proved by its low density) through which the vapours have chances to partially crack before they escape outside of this layer. After liberation from ILC, the vapours do not mix immediately with the cold carrier gas and probably form a more or less developed volume (boundary layer or free jet) inside which they can also partially crack. Such a situation in which pyrolysis products evolved from a high temperature reacting solid, must mix with a much colder medium is conceptually similar to that existing in a cyclone reactor [S] heated at its walls. In this last case, it has been shown that the vapours are partially cracked into gases inside the thin and hot boundary layer existing close to the hot inner surface of the cyclone. We could hence assume that in the present work, the gases have chances to be formed partly inside the porous ILC but also during the brief crossing of the vapours through a hot volume separating the outer surface of the ILC fiom the colder gas flowing through the quartz reactor. These phenomena are still more complicated when char is formed because these vapours can also react during their transit through the thin porous layer of char. These additional processes could bring a possible
1040
explanation for the different gas compositions observed in the presence or absence of char. These mechanisms of gas formation are probably valid for many other experimental devices where solid biomass particles react inside a colder medium. Hence, it seems difficult to completely avoid gas formation in most pyrolysis reactors because of the bad micro-mixing efficiencies existing in the vicinity of the particles.
APPLICATION OF THESE RESULTS TO THE COMPARISON OF PYROLYSIS REACTORS ON THE BASIS OF HEAT TRANSFER MECHANISMS The performances of pyrolysis reactors are often compared on the basis of the fractions of the different recovered products (gases, oils, char). In addition to the maximum biomass throughputs, the comparison criteria usually rely on the values of the residence times and temperatures, two parameters that are often poorly known and defined, and that may be very different for each of the phases involved. The hydrodynamics are often very different from plug flows and generally the true reaction temperatures cannot be accurately measured. It could be interesting to consider another comparison parameter relying on the exchanged heat flux densities. As already analysed by Reed et al. [30], among the main elementary processes of biomass heating, we can mention : radiation from a hot source, direct contact with a hot surface and exchange with a gas. It is possible to estimate the corresponding heat flux densities q in each of these elementary cases and to compare their values as a function of the heat source temperature Tw. The following rough calculations are made on the basis of a cellulose reacting temperature TR = 740 K [4] and of 400 pm particle sizes. 0
Heating by radiation (particles flowing through a reactor having a wall temperature Tw, but without contact with it) :
The view factor and emissivity are expected close to 1, in order to calculate an upper limit for qr. 0
Ablative pyrolysis (a piece of biomass is pressed against a hot moving surface at Tw, [I31 ) :
The calculations are made with P = lo6 Pa. A special case is that of a cyclone reactor in whch the reacting particles flow against its heated walls. The heat transfer coefficient Lyis estimated close to 2700 W rn-’ K-’ [313 and hence : qcy= 2700 (Tw - TR) (3) Heating by exchange with a hot gas :
1041
The Nusselt number representing the exchange between a particle and a hot flowing gas is often calculated by [32] : N k o = 2 + 0.6 Re, '3 Pr 0.33
(4)
Assuming a low particle Reynolds number Re,, N k o is close to 2 and hence :
The values of the thermal conductivityh are taken from Franck [33]. Fig. 4 reports the theoretical variations of qa, qco ,qcyand qr, as a function of Tw. The limit of char formation determined in the present work (q1h = 9 x lo5 W m-*)is also represented. It can be seen that ablation is the most efficient process of heating during which char would be formed only at low wall temperatures (5 800 K). Gas exchange and non concentrated radiation lead to the lowest values of q and as a matter of fact, char is usually observed when each of these two situations occurs separately. However, the available flux densities often result from several contributions as for example the sum of radiation fiom a hot wall and exchange with a hot gas. Having in mind that the equation ( 5 ) related to this last process represents a minimum (qcocan be much higher in the case of high particle Reynolds numbers), the values of q, + qco can approach the limit flux density defined in this work. In these conditions, we can expect that low char fiactions will be formed. The results for the cyclone reactor are also interesting because the values of qcy are not far from q1;n. In that case, the formation of char can be expected only for Tw lower than about 1000 K [7] while only condensible vapours and gases would be observed at higher temperatures. These results for the cyclone are only qualitative indications, the heat transfer coefficients hcy depending on dp and biomass flowrate. 1
I
Present work
M
3
E
* 1Et03 I X 1EM 700
800
900
loo0
1100
1200
1300
1400
1500
T C ~ I I P WTSy~(K) ~
Fig. 4 Heat flux densities transferred to a particle for different elementary processes of
heating
1042
These very simple calculations give only first indications. A more sophisticated approach would need to take into account the fact that lignin gives rise to lugher fractions of char than cellulose. Also, the existence of mass transfer resistances must not be ignored: if the primary products are not efficiently removed from the reacting zone, they have more chances to take part in subsequent reactions, with formation of char.
CONCLUSIONS Cellulose can be pyrolysed under well characterised conditions of absorbed heat flux densities by using a concentrated radiation. Experiments have been performed with cellulose pellets submitted to controlled times of irradiation at the focal zone of an image furnace. The composition and rate of formation of all the products have been studied as a function of flash times. The mass balances are excellent. The results clearly show that the reaction begins with depolymerization processes giving rise to a short life time intermediate compound (ILC) that is liquid at reaction temperature but solid at room temperature. This product partially decomposes into condensible vapours. A fraction of them undergoes a thermal craclung producing gases inside or in the vicinity of the sample. For heat flux densities higher than approximately 9 x lo5 W m”, no char is observed. With the help of these results, a comparison of the performances of pyrolysis reactors can be made on the basis of the types and extents of the elementary heat exchanges processes (radiation, ablation, convection). The calculations show that in some practical cases (ablative pyrolysis and cyclone reactor), it is possible to reach high flux densities and hence avoid the formation of char.
NOTATION particle diameter (pm) heat transfer coefficient (particle / walls of a cyclone) (W m-’ K-I) Intermediate Liquid Compound Nusselt number (exchange with a gas) pressure (biomass against a wall) (Pa) Prandtl number (close to 0.7 for a gas) heat flux density (W m-*) heat flux densities (ablation, gas exchange, cyclone, radiation) (w m-’) limit of the heat flux density for char formation (W m-’) particle Reynolds number flash time (s) temperature (cellulose pyrolysis, heat source) (K) gas thermal conductivity (W m-l K-’) Boltzmann constant (5.67 lo-*W m-’ K-4)
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ACKNOWLEDGMENTS The authors want to thank ADEME (Agence de l'Environuement et de la Maitrise de 1'Energie) who has h d e d this research in the framework of AGRICE (AGRIculture pour la Chimie et 1'Energie) under the contract no 96 01 048.
REFERENCES 1. 2. 3. 4.
5.
6.
7. 8. 9. 10.
11. 12.
13. 14.
15.
Bradbury A.G.W., Sakai Y . and Shafnadeh F. (1979) A Kinetic Model for Pyrolysis of Cellulose. J. Appl. Polym. Sci., 23,3271-3280. Milosavljevic I. and Suuberg E.M. (1995) Cellulose Thermal Decomposition Kinetics : Global Mass Loss Kinetics. Znd. Eng. Chem. Res., 34 , 1081-109 1. GrBnli M. (1996) A Theoretical and Experimental Study of the Thermal Degradation of Biomass, Thesis, The Norwegian University of Science and Technology, Faculty of Mechanical Engineering, Trondheim (N). Ltdt J., Diebold J.P., Peacocke G.V.C., Piskorz J. (1999) The Nature and Properties of Intermediate and Unvaporized Biomass Pyrolysis Materials. In Fast Pyrolysis of Biomass : A Handbook, (Ed. by A.V. Bridgwater ; S. Czernik ; J. Diebold ; D. Meier ; A. Oasmaa ; C. Peacocke ; J. Piskorz ; D. Radlein), pp. 51-65, CPL Press, Birmingham. Antal M.J., Varhegyi G.and Jakab E. (1998) Cellulose Pyrolysis Kinetics : Revisited. Ind. Eng. Chem. Res., 37, 1267-1279. Boutin O., Ferrer M. and LCdC J. (1998) Radiant Flash Pyrolysis of Cellulose Evidence for the Formation of Short Life Time Intermediate Liquid Species. J. Anal. Appl. Pyrolysis, 41, 13-31. LCdC J.,Verzaro F., Antoine B., Villermaux J. (1986) Flash Pyrolysis of Wood in a Cyclone Reactor. Chem. Eng. Process., 20,309-3 17. LCdC J. (2000) The Cyclone : A Multifunctional Reactor for the Fast Pyrolysis of Biomass. Ind. Eng. Chem. Res., 39, 893 - 903. Bilbao A., Millera A. and Arauzo J. (1988) Product Distribution in the Flash Pyrolysis of Lignocellulosic Materials in a Fluidised Bed, Fuel, 67, 1586-1588. Bridgwater A.V. (1999) Principles and Practice of Biomass Fast Pyrolysis Processes for Liquids. J. Anal. Appl. Pyrolysis, 51,3 -22. Bridgwater A.V,. Peacocke G.V.C (2000) Fast Pyrolysis Processes for Biomass. Renewable and Sustainable Energy Reviews, 4, 1-73. Venderbosh R.H., Janse A.M.C., Radovanic M., Prins W., Van Swaaij W.P.M. (1997) Pyrolysis of Pine Wood in a Small Integrated Pilot Plant Rotating Cone Reactor. In Biomass Gasification and Pyrolysis, (Ed. by M. K a l t s c h t t & A.V. Bridgwater), pp. 345-353, CPL Press, Newbury. Ltdt J,, Panagopoulos J. , Li H.Z., Villermaux J. (1985) Fast Pyrolysis of Wood : Direct Measurement and Study of Ablation Rate. Fuel, 64, 1514-1520. Ltdt J. (1994) Reaction Temperature of Solid Particules Undergoing an Endothermal Volatilization. Application to the Fast Pyrolysis of Biomass. Biomass and Bioenergy, 7,49-60. LtdC J. (1996) Influence of the Heating Conditions on the Thermal and Chemical Behaviour of Solid Undergoing a Fast Endothermic Decomposition. J. of Thermal Anal., 46,67-84.
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16. Diebold J.P. (1994) A Unified, Global Model for the Pyrolysis of Cellulose. Biomass and Bioenergy, 7, 75-85. 17. Di Blasi C. (1996) Heat, Momentum and Mass Transport through a Shrinking Biomass Particle Exposed to Thermal Radiation. Chem. Eng. Sci., 51, 11211132. 18. Miller R.S. and Bellan J. (1997) A Generalized Biomass Pyrolysis Model Based on Surimposed Cellulose, Hemicellulose and Lignin Kinetic. Combust. Sci. Tech., 126, 97-1 12. 19. Lanzetta M, Di Blasi C., Buonanno F. (1997) An Experimental Investigation of Heat Transfer Limitations in the Flash Pyrolysis of Cellulose Ind. Eng. Chem. Res., 36, 542-552. 20. Varhegyi G., Jakab M., Antal M.J. (1994) Is the Broido-Shafuadeh Model for Cellulose Pyrolysis True? Energy and Fuels, 8, 1345-1352. 21. Antal M.J. and Varhegyi G. (1995) Cellulose Pyrolysis Kinetics : the Current State of Knowledge. Ind. Eng. Chem. Res., 34,703-717. 22. Boutin 0. (1999) Analyse des processus primaires de dbgradation thermochimique de la biomasse. INPL thesis, Nancy, France. 23. Nordin S.B., Nyren J.O. and Back E.L. (1974) An Indication of Molten Cellulose Produced in a Laser Beam. Textil Research J., 152-154 24. Antal M.J. (1983) Effects of Reactor Severity on the Gas Phase Pyrolysis of Cellulose, and Kraft Lignin, devired Volatile Matter. Ind. Eng. Chem. Prod. Res. Dev., 22, 366-375 25. Li H.Z. (1984) Pyrolyse kclair de baguettes de bois - Modile de fusion. DEA INPL, Nancy, France. 26. Diebold J.P., Power A.J. (1988) Engineering Aspects of the Vortex Pyrolysis Reactor to Produce Primary Pyrolysis Oil Vapours for use in Resins and Adhesives. In : Research in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater & J.L. Kuester), pp. 609-628, Elsevier Applied Science Publishers, London and New York. 27. Graham R.G., Bergougnou M.A., Free1 B.A. (1994) The Kinetics of VapourPhase Cellulose Fast Pyrolysis Reactions. Biomass and Bioenergy, 7,33-47. 28. Peacocke G. (1994) Ablative Pyrolysis of Biomass. PhD Thesis, Aston University, Birmingham (England). 29. Boutin O., Ltdt J. , Li H.Z. and Kiener P. (1997) Temperature of Ablative Pyrolysis of Wood. Comparison of Spinning disc and Rotating Cylinder Experiments. In : Biomass Gaszjkation and Pyrolysis, (Ed. by M. Kaltschrmt & A.V. Bridgwater), pp. 336 - 344, CPL Press, Newbury. 30. Reed T., Diebold J., Desrosiers R. (1980) Heat Transfer Perspectives. In Specialists ' Workshop on Fast Pyrolysis of Biomass, SERVCP-622-1096, pp. 720, Copper Mountain. 31. Ltdt J., Li H.Z. , Soulignac F., Villermaux J. (1992) Le cyclone rtacteur. IV: Mesure de l'efficacitt des transferts de chaleur entre les parois et les phases gazeuse et solide. Chem. Eng. J., 48, 83-99. 32. Levenspiel 0. (1972) Chemical Reaction Engineering, 2"* Edn, John Wiley & Sons, Canada. 33. Franck M.W. (1984) Heat Transfer. Adson-Wesley Publishing Company, London (England).
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Modelling and measurements of drying and pyrolysis of large wood particles Larfeldt J.*, Lecher B.* and Melaaen M. Chr.** *Departmentof Energy Conversion, Chalmers University of Technology, 9 4 1 2 96 GSteborg, Sweden, +46 31 7721430, +46 31 7723592 fax), Telemark Institute of Technologv, Porsgrunn, Norway, +47 35 575286, +47 35 575100 f a ) . *I
ABSTRACT: Internal temperature distribution and mass were measured during drying and pyrolysis of cylindrical samples of wood of 0, 14 and 44% moisture. A onedimensional model of drying and pyrolysis is modified to reflect the anisotropy of the wood. Inclusion of an instant axial convective mass flow is shown to reduce the time of conversion compared to simulations with no axial flow. This mass flow, contrary to a convective mass flow through a porous structure, is not in thermal equilibrium with the solid phase. The gas leaves at a lower temperature compared to the temperature of the solid phase and is thus neglected in the energy equation. Intrinsic gas permeabilities from literature were found to produce unrealistically high interior pressures. Simulations show that the permeability of liquid has a large influence on the time of conversion of the moist wood samples. Liquid permeabilities lo4 times lower than the gas permeability were found. There was a discrepancy between measured and simulated temperature gradient found for moist samples probably explained by radial variations in material properties and structural dependence of the heat transfer properties.
BACKGROUND This work concerns w i g and pyrolysis of thermally thick wood particles. The work applies to, for instance, log firing in small scale boilers for house heating. In these boilers the rate of combustion is governed by each wood log. The boundary condition for a log is given by the design of the furnace and by the surrounding logs. In previous work structural changes during pyrolysis were included, which was shown to reduce the time of pyrolysis substantially [I]. Furthermore, the heat transfer properties of charcoal, produced from the same samples as those treated here, were investigated [2, 31. These properties become important for a correct description of the rate of pyrolysis in the interior of large biomass particles. In earlier work the influence of anisotropy of wood on the convective flow of gas was neglected. In this work the drying process is modelled together with the pyrolysis and compared with measurements. The onedimensional model is modified to reflect the anisotropy of the convective flow in wood, especially important during drying.
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EXPERIMENTAL WOOD SAMPLES AND EXPERIMENTAL PROCEDURE Cylindrical pieces of birch wood with a radius of 25 mm, a length of 300 mm, moisture contents of 44%, 14% and 0% and an average dry density of 41 0 kg/m3 were pyrolysed in an inert atmosphere using a furnace wall temperature of 700 ‘C. Initially the radiative heat flux to the sample was approximately 50 kW/m2, decreasing as the sample surface temperature increase. The wood pieces were turned from stems of approximately 150 mm diameter, the centre of the stem coinciding with the centre of the resulting cylinders. The samples were heated in an electrical furnace with an inner diameter of 15 cm and a height of 110 cm. The experiments were separated into two types carried out in sequence, measurements of internal temperature distribution and mass, since it was found that the thermocouples mounted in the wood samples interfered with the mass measurement. The temperature was measured in five radial positions in the interior of the wood piece and at the surface by interior thermocouples mounted in axial1.y drilled holes with a length of 150 mm. Due to difficulties to accomplish straight bore holes, the positions of the thermocouples were determined by X-ray analysis [4].The normalised mass was given by an amplified voltage signal from a strain gauge.
RESULTS FROM EXPERIMENTS The measured temperature distributions for 44, 14 and 0 % moisture are shown in Figure la-c. The profiles represent three or, in the case of dried wood, two measurements. In case of 44% moisture the total time of heating was approximately 2000 seconds, for 14 % moisture 1000 seconds and for dry wood 600 seconds. Drying plateaus develop in the wet samples at a temperature of around 100°C. The wood cylinders were converted into charcoal samples of the same length, approximately 300 mm, with a radius of 20+2 mm for all moisture contents. The samples cracked at relatively regular intervals. The cracking was more intense in the case of 44 % moisture, and the charcoal structure “fell apart” completely as the sample holder was removed. This was not the case for the dry sample and the sample of 14% moisture, where cracks were formed across the cylinder, distributed in the axial direction at about 20 mm distance. The structural changes during drying and pyrolysis ‘are related to the initial moisture content of the sample and to the structure of the wood. The intense cracking in the case of 44 % moisture is also seen as irregularities in the interior temperature profiles at 21, 18 and 8.5 mm radii in Figure la. Also in the case of dry wood an irregularity is seen for the profile at 20 mm radius. All irregularities occur in the temperature range of 200 to 300 “C, where the pyrolysis is initiated. The irregularities are explained by sudden changes in the position of the thermocouples as cracks are formed during pyrolysis. The instantaneous mass of a total of nine samples, normahsed to their initial mass, are shown in Figure Id. The time of pyrolysis is in agreement with the time of heating reported above; 500 seconds for dry wood, 1000 for 14 % moisture and 2000 seconds for samples of 44% moisture. In the case of dry samples and samples of 44 % moisture the measured masses are in complete agreement but at 14 % moisture there is a certain
1047
scatter. The dry samples have a char yield of approximately 20% of the initial dry mass. Assuming a char yield of 20%, as in the dry case, the total mass remaining in the case of 14 and 44 % moisture should be 11 and 17% respectively. This agrees with Figure Id. MATHEMATICAL MODEL A model describing one-dimensional heat transfer in porous material [S] is used and modified to study drying and pyrolysis of large wood particles. The model solves the conservation equations of energy, solid phase, liquid water, bound water and the concentrations of species in the gas phase of the porous material. The phases are assumed to be in thermal equilibrium and the variables are calculated as volume averages [ 6 ] .
Time [seconds]
Time [seconds]
d) I
Y
v1 v1
8
a
z8
.-2
0'
500 750 Time [seconds]
250
1000
O*O;
.
so0 Id00 1500 2000 25bo '
'
'
Time [seconds] 1. Measured internal temperature in wood samples (sample no. I - solid, I1 dashed and 111- dotted lines) of; a. 44% moisture at the radial positions of 25,21, 18, 17, 14, 14, 8.5 and 6 mm. b. 14 % moisture at the radial positions of 25,21, 18, 15, 10,9,7 and 2 mm. c. 0 % moisture at the radial positions of 25,23,20, 16, 15, 11, 5 , 2 and 2 mm. d. Measured normalised total mass of nine wood samples; three of 0%) three of 14% and three of 44% moisture
1048
CONSERVATION EQUATIONS The conservation of solid phase is written:
where char contributes to the solid phase and is formed from the reacting wood during pyrolysis, o+, at a given char yield, x. The third term on the right hand side accounts for structural changes during pyrolysis (described below). The concentration of species j ('j=gas and tar from pyrolysis, water vapour from drying and air) in the gas phase is calculated from:
v
at
H
where the last two terms on the right hand side estimate the shrinking and the axial (i.e. 2D) flow of gases in the samples (discussed elsewhere [7]). The diffusion of gas species is neglected in the axial direction [7]. Water is described by convection of liquid water and diffusion of bound water in the wood structure as:
Also here terms are included to account for shrinking and axial flow of liquid water.
The axial diffusion of bound water is neglected in the equations [7]. The temperature, T, is solved from the non-conservative form of the energy equation:
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It has been shown that shrinking will not appear in the energy equation [l]. The axial flow of gas is included in the energy equation as the last term on the right hand side.
CONVECTIVE FLOWS The velocities of liquid water and gas are calculated from Darcy's law, and the radial velocities are expressed as:
The axial velocities are given by:
DRYING
Liquid water, sometimes referred as free liquid water, is present in the porous structure at moisture contents above the fibre saturation point, here assumed constant at all temperatures: 0.2 kg water per kg dry wood. The rate of evaporation in Equation (3), ~ + qis,indirectly solved from Equation (2) for water vapour, j=Hz0. In the presence of liquid water equilibrium vapour pressure is assumed in the gas phase described as
[81:
(P~~~)= ' ' 101325exp(17.58'~ 5769/T- 5.686. lo-")
(9) At moisture contents below the fibre saturation point the vapour pressure sinks as [9]:
The capillary pressure in Equation (6 and 8) is calculated as [lo]:
Pw = 1.364.lo5(0.128- 1.85 104T)(X, + 1.2.10-4)0'63 *
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(1 1)
PYROLYSIS Four independent, parallel reactions with kinetic data, KI-K4, empirically derived from measurements of the mass loss of small samples of birch wood [9] have previously been shown to describe the pyrolysis reaction [l].
wood
KI+K2+K3+K4
> f(1- X)GUS+ (1 - f)(l - x ) T u ~+%Char (12)
The four reactions reflect the constituents of birch; two types of hemicellulose (13% and 36%), cellulose (39%) and lignin (1 1%) [9]. The rate of pyrolysis, q,is given as the sum of these four reactions. The pyrolysis reactions result in gas, tar ( H . 6 8 ) and char according to the right hand side in Equation (12) and the production terms for Equation (2) are found as wgas = f( 1-x)q, and o,= ( 1-f)( 1-x)(op. STRUCTURAL CHANGES A shrinking model based on three empirical shrinking parameters describes the evolution of the solid and gas volumes during pyrolysis [ 1 11 as,
The shrinking parameters; a, p and y (1 .O, 0.5 and 0.66 in all cases treated) are related to the final sample radius, rc, and char yield found in the experiments [l]. The volumes of gas and solid are related to the initial volumes of the wood sample. The shrinking during drying is neglected, motivated by results from [12]. The degree of pyrolysis, q, is calculated as the mass of virgin wood minus the char fraction divided by the initial mass of volatiles.
ADDITIONAL EQUATIONS The gas pressure is calculated for an ideal gas and the gas density from the sum of gas species:
The sum of the volume fractions of all phases is equal to 1:
105 1
E,
+ E, + E, + €,
=1
(17) The liquid and bound water volume fractions are calculated from the water content divided by the intrinsic density of water (1000 kg/m3). The volume fraction of gas, i.e. the porosity, is given both by the gas law (i.e the gas density in Equation 14) and the volume of gas (i.e. structural changes arising from Equation 13). The solid volume fraction is calculated from Equation (1 2).
SOLUTION PROCEDURE
From Equations (1) to (16) the solid density, liquid water density, bound water density, the concentrations of species in the gas phase, temperature, gas density, pressure, solid, liquid and gas phase volumes in a cylinder of wood are calculated versus drying and pyrolysis time. The convective terms in Equations (2) and ( 6 ) are discretised by firstorder upwinding, while central differentiation is used for the diffusive terms. Grid independence was tested. The time integration is managed by the code DASSL [13] that implicitly solves a set of differential and algebraic equations. The effect of the shrinking sample volume on the gas velocity was found to be negligible and is therefore not included in the equations. The temperature of the sample's surface is measured during the experiments and is given as the boundary condition in the simulations. The pressure at the surface equals the surrounding atmospheric pressure, also given as a boundary condition.
RESULTS AND DISCUSSION AXIAL FLOW OF GAS
The permeability of hardwoods (such as birch treated here) is lo4 to 1O8 times higher in the axial direction than in the radial and tangential directions [ 141. Therefore the axial flow was included in the conservation equations (Equation 2, 3 and 4). A discussion on the anisotropy of wood, the inclusion of axial convective terms and the omission of axial difhsive terms are presented in [7]. In the conservation of gas species and of liquid phase the axial flow was estimated by Equation 7 and 8 assuming, for simplicity, a linear pressure gradient in the axial direction of the sample. For the energy equation (Equation 4) two cases were studied, q,=O and an axial heat loss according to:
(T'" 4, =4c,(pg)"uf
) -( T ) H
where Tsurfis the temperature of the gas leaving the sample. Figure 2a shows the measured case of dry wood and three simulated interior temperatures for Case 1, 2 and 3 in Table 1. Case I corresponds to a case without axial flow. Including axial flows in Equation 2 and 3 and using Equation 18 in Equation 4 the time of conversion is seen to increase, Case 2 in Figure 2a. If, instead, the axial flow in the energy equation is neglected the time of conversion is decreased and the simulated case shows a better fit to experimental data. Figure 2b shows the measured case of 14% moisture and two
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simulated interior temperatures for Case 4 and Case 5 in Table 1. Neglection of the axial flow in the energy equation also here is shown to agree with measured data, while Equation 18 overestimates the conversion time of the sample. In theory, all convective flows will increase the conversion time since energy is released from the sample. A convective flow of instantaneously released gas through cracks formed during pyrolysis is illustrated in Figure 3. In these cracks the gas is no longer in thermal equilibrium with the solid structure. The gas temperature is thus lower at instant convection compared to convective flow through the porous structure, which explains the negligible influence on the energy equation. Another explanation may be that the energy loss due to axial flow of gas is compensated for by an axial conduction of heat in the samples not accounted for in the model. From comparison of Case 1 and 3 in Figure 2a it is shown that including an instantaneous convective mass flow will reduce the time of conversion compared to the case of no axial flow. Table 1. Simulated cases Case Moisture Axial flow in Eq. 2 and 3 1 2 3 4 5 6 7
No Yes Yes Yes Yes Yes Yes
0% 0% 0% 14% 14% 44% 44%
0
18 18
Eq. 18 0 0 0
20 20 20 20
800
I
10 1 3
10-14
I
B 600
600
O -
E
400
f
200
g!3 200
Q
t-”E
K‘, m2
b) 800
2
f
Charcoal radius, rc [mm] 18
Eq. 18
a)
-u
Axial flow in Eq. 4; qz
0
200
400
600
800
400
0
1000
250 500 750 100012501500
Time [seconds]
Time [seconds]
Fig. 2. Influence from axial flow. Comparison of a. measured temperatures (solid lines) from Figure Ic and simulated centre temperatures of the wood sample for Case 1,2 and 3, Table 1, 0% moisture. b. measured temperatures (solid lines) from Figure l b and simulated centre temperatures of the wood sample for Case 4 and 5, Table 1, 14% moisture.
Fig. 3. Illustration of convective flows in wood during pyrolysis [ 151.
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INFLUENCE FROM GAS P ERMEABLIL TY ON VAPOUR PRESSURE
Figure 4a shows simulated temperatures, Case 5, compared to measurements. The model fails to predict the drying plateaus at the sample's surface - but in the interior of the sample the drying plateaus agree with the measured ones. Figure 5a shows the drying plateaus in a larger scale. All measured plateaus fall in the range of 100 to 105"C, which corresponds to vapour pressures of 1 to 1.2 times the atmospheric pressure in the structure. The intrinsic permeability of gas in Case 5 was 10" in the axial direction and lo4 times lower in the radial direction. This resulted in interior pressures in the same range as seen in measurements. Also for the case of 44% moisture the measured drying plateaus, shown in Figure 5b, fall in the 100-105°C range. The intrinsic gas permeability is thus similar in the case of 44 % moisture. The intrinsic gas permeability of 10" found in the simulations here is, however, higher than to 10'' in in the axial direction and reported values in literature of 10" to the radial direction [9]. The lower intrinsic permeability resulted in unrealistic high vapour pressures in the structure (for instance 6 times the atmospheric pressure). It is noted that the temperature of the drying plateaus tend to increase towards the centre of the 14% moisture samples which is not the case for the 44% m d t u r e samples. The simulated mass is in agreement with measured data as shown in Figure 4b. INFLUENCE FROM LIQUID PERMEABILITY ON TIME OF CONVERSION
In the case of 44 % moisture there is liquid water present in the structure and also the permeability of liquid influences the conversion time of the sample. Figure 6 shows the measured temperature profiles of Figure 1 and the two simulated Cases 6 and 7 in respectively. In both Table 1, simulated for liquid axial permeabilities of 1O-I3 and cases the radial permeability is assumed to be lo4 lower. It is seen from Figure 6 that the intrinsic permeability of liquid has a large influence on the pyrolysis time. A higher permeability leads to a larger transport of water through the wood and less water evaporates inside the sample, which reduces the time of conversion.
-
130
u
0,8 E
.
. Simulated
TI 0,6
.-P
7 0,4
€
2 "0
250 500 750 1000 1250 1500
02 O'OO
Time [seconds]
250 500 750 1000 1250 1500 Time [seconds]
Fig. 4. 14 % moisture a. Measured temperatures (solid lines) previously shown in Figure Ib and 2b and simulated temperatures (dotted lines) Case 5, Table 1, 14% moisture. b. Measured normalised total mass, previously shown in Figure Id, compared to simulated mass Case 5, Table 1, 14% moisture.
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Comparison with measured data in Figure 6 shows that the liquid permeability is in the range of lo-'' and Case 7 for the lower liquid permeability of I O l 4 is shown together with measured data in Figure 7. Also here the drying plateaus close to the surface of the sample were not well predicted, similar to the case of 14 % moisture. The simulated mass has a larger loss in the beginning and a slower at the end compared to the measured ones. Possibly the drying model included will overestimate the axial flow in the simulation. It is concluded that the transport of water in the wood structure has a large influence on the drying and pyrolysis processes. Experimental data on transport properties of birch wood are needed for more accurate simulations of wet fuels. Also the permeability may vary with the radius of the sample.
INFLUENCE FROM VARIOUS FACTORS
The model uses material properties and models of these properties, such as; thermal conductivity, permeability, diffusivity, specific heats, heat of pyrolysis, final sample radius and so on. The appendix gives an overview of the data used in the simulations.
Time [seconds] Time [seconds] Fig. 5. Measured drying plateaus a. and simulated (dotted lines) at 14 % moisture from Figure 4a b. at 44 YOmoisture from Figure 1a.
Time [seconds] Fig. 6. Influence from liquid permeability. Measured temperatures (solid lines) from Figure l a and 5b, and simulated centre temperatures (dotted lines) of the wood sample for Case 6 and 7, Table 1,44% moisture.
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From Figure 4a and 7a it is seen that after drying, above the drying plateau, the measured temperature gradients are steeper than the simulated ones. The material properties used in simulations are based on previous work where dry wood and charcoal were investigated. Using the data according to the Appendix, a case of dry wood is simulated, Case 3 in Table 1, and compared to measured data in Figure 8. Here the simulated and measured temperature profiles show better agreement. Also the simulated mass of the sample agrees with the measured mass of the samples. There are several explanations for the discrepancy in measured and simulated temperatures in the initially moist samples. For instance the structural dependence of the heat transfer properties. The charcoal porosity can be divided in pores, originating from the virgin wood structure, and cracks formed during pyrolysis [3]. The formation of these pores and cracks, i.e. the structural changes, are known to vary with the heating rate during pyrolysis and the initial moisture content. The heating rate is known to vary with the sample radius. No radial variations in the structural changes or any effects of variations in the distribution between large cracks and small pores were accounted for. Large cracks may increase the radiation heat transfer in the porous structure while a high yield of pores may reduce the specific heat of char. Also the finite sample length leads to axial heating of the ends that in turn cause initial rapid drying of the ends which is not accounted for in the model.
MODEL SENSITIVITY In earlier work a sensitivity analysis of simulations of dry wood shows an 8% reduction in time of pyrolysis at 1 mm reduction of the final charcoal radius [l]. This was shown to be in the same range as the influence from an exothermic heat of pyrolysis (150 kJkg). In this paper the inclusion of an axial convective flow is shown to influence the time of pyrolysis in the same range, a reduction of 6% (evaluated from Figure 2a where the times of conversion of Case 1 and Case 3 are estimated to 750 and 708 seconds respectively). For wood with a moisture content above the fibre saturation point the axial liquid permeability is shown to dominate the influence on the conversion time. a) 800
G 600
Simulated
L
2
f
400
@ 200
g
0 '
500
1000 1500 2000 2500
Time [seconds]
Time [seconds] Fig. 7. 44% moisture a. Measured temperatures (solid lines) previously shown in Figure la, 5b and 6 and simulated temperatures (dotted lines) for Case 7, Table 1. b. Measured normalised total mass, previously shown in Figure Id, compared to simulated mass for Case 7, Table 1.
1056
-
1,o
3
03
‘CI
0,6
Y
oz 3
600
E
2
$ 400 E g 200
.-8
E z”
0,4 0,2
n
“0
250
500
750
Time [seconds]
1000
o’oO
250
500
750
1000
Time [seconds]
Fig. 8. Dry wood. a. Measured temperatures (solid lines) previously shown in Figure l c and 2a, and simulated temperatures (dotted lines) for Case 3, Table 1. b. Measured normalised mass, previously shown in Figure Id, compared to simulated mass for Case 3, Table 1.
CONCLUSION
The internal temperature distribution and the mass were measured during drying and pyrolysis of cylindrical samples of wood of 0, 14 and 44% moisture. A onedimensional model of drying and pyrolysis is modified to reflect the anisotropy of wood. Inclusion of an instant axial convective mass flow is shown to reduce the time of conversion compared to simulations with no axial flow. This flow, contrary to a convective mass flow through a porous structure, is not in thermal equilibrium with the solid phase and is thus neglected in the energy equation. Measured drying plateaus occurred in a temperature range of 100 - 105°C corresponding to vapour pressures of 1 -1.2 times atmospheric pressure. Intrinsic gas permeabilities from literature were found to produce unrealistically high interior pressures. Simulations show that the permeability of liquid has a large influence on the time of conversion of the moist wood samples. The simulations agreed with measured data for a liquid permeability lo4 times lower than the gas permeability of lo-’’ found. For the moist samples, the drying plateau close to the surface of the sample and the temperature gradient above this plateau were not well predicted in the simulations. On the other hand a comparison of simulation and measurements on dry wood shows good agreement. The discrepancy for moist samples could be explained by structural dependence of the heat transfer properties.
ACKNOWLEDGEMENT
The Swedish contribution to this work was financed by the Small Scale Combustion Programme, Swedish Energy Administration.
1057
NOMENCLATURE
:P
D
ih F H K K M N P
Q R
Ro
r r
U
V X
Specific heat (Jkg K) Gas diffusivity (m2/s) Enthalpy of formation (Jkg) Fraction (-) Length of samples (m) Thermal conductivity (W/mK) Reaction rate (11s) Molecular mass (kg/mole) Number of species Pressure (N/m2) Heat loss (W/m3) Radius (m) Universal gas const. (J/mole K) Temperature (K) Time (s) Velocity ( m / s ) Volume (m3) Moisture (kg waterikg wood)
Greek
o!
P X
Porosity (volume fraction) Phase Shrinking parameter Degree of pyrolysis Permeability (mZ) Dynamic viscosity (kg/ms) Production rate (kg/m3s)
E
Q Y rl K
CL
w Index b C
eff fSP g i j 1 P r S
Shrinking parameter Shrinking parameter Char yield (kg charikg wood)
sat surf W
Bound Char Effective Fibre saturation point Gas Initial Gas specie Liquid Pyrolysis Radial Solid Saturated Surface Dry wood
APPENDIX. DATA USED IN SIMULATIONS Diffusivity of gas species in gas phase, Dgr=0.1 5*104(T/298)’.75(P,JP)/1000, Dgz=20D, Diffusivity of bound water, Db,r=exp(-9.9-4300/T+9.8(
/
k c,
Wood 0.20 2439
Char 0.09 1000
Water 0.658 4200
1058
Gas 0.026 1000
Table A2. Data of gas and liquid permeability used in simulations.
K:
K:
Intrinsic Wood Char 10E-14 10E-11 1OE- 10
10E-18 K1 K/, 10E-14
1OE-08
X*=O 1
Relative [ 161 O<X‘
I 0
0.95(X/0.48)2
0.05(X-0.48)/(X,,-0.48)+0.95 (X&d3
0
REFERENCES 1 Larfeldt, J., Lecher, B. and Melaaen, M.C., ”Modelling and measurements of the
pyrolysis of large wood particles”, Fuel, 79, 1637-1643,2000. 2 Suleiman, B.M., Larfeldt, J., Leckner, B., Gustavsson, M., ”Thermal conductivity and diffusivity of wood”, Wood Science and Technology, 6, (33), 465-473, 1999. 3 Larfeldt, J., Leckner, B. and Melaaen, M.C., ”Modelling and measurements of heat transfer in charcoal from pyrolysis of large wood particles”, Biomass and Bioenergy, 18,507-5 14,2000. 4 Green A. and Larfeldt J., ”Results from measurements of drying and pyrolysis of wood cylinders”, Report A98-222, Chalmers University of Technology, Goteborg 1998. 5 Melaaen M.C., “Numerical Analysis of Heat and Mass Transfer in Drying and Pyrolysis of Porous Media”, Numerical Heat Transfer, Part A, 29,33 1-355, 1996. 6 Whitaker S., “Simultaneous Heat, Mass and Momentum Transfer in Porous Media: A Theory of Drying”, Advances in Heat Transfer, 13, 119-203, 1977. 7 Larfeldt J., “Drying and Pyrolysis of Logs of Wood”, PhD Thesis, Department of Energy Conversion, Chalmers University of Technology, Goteborg, 2000. 8 Lai, W.-C., ”Reaction Engineering of Heterogeneous Feeds: Municipal Solid Waste as a Model”, PhD thesis, University of Washingtonm Seattle, 1991. 9 Gronli M., ”A theoretical and experimental study of the thermal degradation of biomass”, PhD Thesis, Division of Thermal Energy and Hydro Power, Trondheim, 1996. 10 Spolek G.A. and Plumb O.A., ”Capillary Pressures in softwoods”, Wood Sci Technol., 15, 189-199, 1981. 11 Di Blasi C., “Heat. Momentum and Mass Transport Through a Shrinking Biomass Particle Exposed to Thermal Radiation”, Chem. Sci. Eng., 51, 1121-1 132, 1996. 12 Fyhr C., ”Superheated Steam Drying of Wood Chips in Pneumatic Conveying Dryers”, PhD Thesis, Department of Chemical Engineering Design, Chalmers University of Technology, Goteborg, 1996.
1059
13 Petzold L. R., ”A Description of DASSL: A DifferentiaVAlgebraic System Solver”, Scientific Computing, Stepleman et al. (eds.), 65-68, IMACS, North-Holland, Amsterdam, 1983. 14 Siau J. F., ”Transport Processes in Wood”, Springer-Verlag, Series in Wood Science, 1984. 15 Roberts A.F., “Problems Associated with the Theoretical Analysis of the Burning of Wood”, 131h Symp. (Int.) on Combustion, The Combustion Institute, Pittsburgh, 893-903,1971. 16 Perrd, P. And Degiovanni A., “Simulation par Volumes finis des Transferts Couple’s en Milieux Poreux Anisotropes: Sdchage du Bois A Basse et a haute Temperature”, Int. J. Heat Mass Transfer, 33, 2463, 1990.
1060
Thermal Analysis and Kinetic Modelling of Wheat Straw Pyrolysis M. Stenseng, A. Jensen & K. Dam-Johansen Department of Chemical Engineering, Technical University of Denmark, Lyngby, Denmark
ABSTRACT Simple kinetic models have been tested to fit the weight loss during pyrolysis of wheat straw under different temperature programs. The models include a single first order reaction, a three parameter nucleation model, a distributed activation energy model and a superposition model based on parallel first order reactions. Although the distributed activation energy model appears to give the best fit to data, it is still concluded that a biomass pyrolysis model valid over a wide range of process conditions is not available in the literature. INTRODUCTION The use of straw in power plants in Denmark has resulted in a need for research in the pyrolysis of straw. Very little regarding straw pyrolysis and straw pyrolysis kinetics are reported in the literature. A few thermogravimetric studies of the pyrolysis of biomass are reported in the literature [1,2,3,4,5],which include wheat straw. In those studies, the pyrolysis experiments were carried out at one set of experimental conditions only. No investigation of the influence of e.g. heating rate and final temperature has been found in the literature. In some of the studies [1,3,5] demineralized (washed) wheat straw was investigated as well. The washing was carried out in order to remove most of the salts and minerals in the wheat straw. Varhegyi et al. [5] washed the straw in order to determine the pyrolysis kinetics by a superposition model and wished to eliminate the influence of the salts. Jensen et al. [l] investigated the influence of washing the straw and adding salts to the washed straw in order to get a better understanding of the way the salts are bound in the straw. In practice, washing of straw may become a process used in the power plants in order to remove mainly potassium to reduce the deposition and corrosion in straw fired power plants [6]. Few straw pyrolysis models have been found in the literature. A few investigators [4,5] tested a superposition model for straw pyrolysis, in which the pyrolysis of straw is described by a number of parallel reactions, typically two or three, representing the pyrolysis of the main biomass constituents. The model was found to fit the weight loss data reasonably well when fitted to data at a single heating rate. Lanzetta and Di Blasi [7] made a model in which the pyrolysis of wheat straw is described by competing parallel and consecutive reactions. They tested their model for high heating rates (1 500-
1061
4200"C/min) and temperatures below 375"C, and found reasonable good agreement with data. The main objective of this paper is to test simple kinetic models for the pyrolysis of wheat straw. In order to obtain well-defined conditions, especially regarding the sample temperature,low heating rate experimentsin a thermogravimetric analyzer was preferred.
MATERIALS AND EXPERIMENTAL PROCEDURES SAMPLES
The samples used in this study were Danish wheat straw pellets which were milled on a laboratory ultra centrifugal mill (Retzsch ZM 100) mounted with a 1 mm screen. In addition, washed wheat straw was used. The washing procedure was as follows: 1 g of milled straw was placed in 200 ml deionized water and stirred for 20 hours at room temperature, then filtered and washed with 300 ml deionized water, before drying [ 1,6,8]. EXPERIMENTAL SETUP AND PROCEDURES
The experiments were carried out in a Netzsch STA 409 C (Simultaneous Thermal Analysis - STA) in the TGA/DSC configuration. The STA has a vertical sample carrier with a reference and a sample crucible, and in order to account for buoyancy effects, a correction curve with empty crucibles was first conducted and then subtracted from the actual experiments. PlatinumlRhodium crucibles were used in order to get the best possible heat transfer. The thermocouple for each crucible was positionedjust below and in contact with the crucible. The temperature obtained from the measurement is the temperature in the reference side. This temperature is converted to the temperature in the sample side by using the DSC-signal in pV and a temperature-voltage table for the thermocouple.The product gases were swept away by 100Nml/min nitrogen which exited the top of the STA. The STA was calibrated for temperature and sensitivity (DSC) with metal standards at each heating rate. Samples of 5 mg straw were pyrolyzed to a maximum temperature of 700°C. The influence of sample mass for the pyrolysis of wheat straw and washed wheat straw has previously been investigated for samples between 2 and 20 mg at a heating rate of 40"C/min [9], and it was found that the sample mass had no influence on the pyrolysis process. Heating rates of 5 , 10 and 40"C/min was applied. Two types of temperature programs were used. In both types of temperature programs the sample was first heated to 110°C and kept at that temperature for 30 min in order to remove the moisture. In the first temperatureprogram type, the sample was heated at a specifiedheating rate to 600°C. In the second temperature program the sample was heated at 10"C/min to 300"C, kept isothermal at that temperature for 60 min before heating at lO"C/min to 600°C. In both types of temperature programs the sample was then kept isothermal at the final temperature for 30 min before burning off the char in a nitrogedoxygen mixture in order to determine the ash content.
RESULTS The pyrolysis of wheat straw and washed wheat straw show different pyrolysis behavior. Besides a lower onset temperature of the pyrolysis and a higher char yield of the wheat straw, the main difference is the shape of the mass loss rate curves. The pyrolysis of wheat
1062
straw show an asymmetrical peak with a long tail, see Fig. 1. However, the pyrolysis of washed wheat straw shows two peaks, corresponding to hemicellulose and cellulose decomposition, and a tail, corresponding to lignin decomposition. 1
0.003
0.9
..,.....Washed wheat straw
\
- 0.0025
Xmo5
.*."......_..._"
i...
....._.
200
250
300 350 400 Temperature ("C)
450
Fig. I The pyrolysis of wheat straw and washed wheat straw at lO"C/min.
The shape of the mass loss and mass loss rate curves determine the types of kinetic models that can be applied to fit the data. Simple kinetic models, which only include one reaction, may be applied to describe the pyrolysis of wheat straw, but cannot describe the two peaks observed in the pyrolysis of washed wheat straw. Different kinetic models have been applied to fit mass loss data, including (with the parameters of each model in parenthesis): 1. a single first order reaction (A, E) 2. a three parameter nucleation model (A, E, m) 3. a distributed activation energy model using a natural logarithmic distribution (A, a, PI 4. a superposition model based on n first order reactions (c, A,, Ei, i=l,n) All four models were used to fit the pyrolysis of wheat straw, but only the superposition model was used to fit the pyrolysis of washed wheat straw. All four models have the flaw that they are not able to predict the char yield, which then has to be given as input to the models. Prediction of char yield can only be done by more complicated models. The most simple model is the single first order reaction, which is described by: dX = k ( l -X) dt
in which X is the fractional weight loss normalized with the ultimate weight loss, and the rate constant (k)is given by the Arrhenius equation (k=A exp(-E/R,T)). An algorithm to fit the experimental data to Eqn. (1) was made in Fortran using a least squares method by Fletcher [ 101, in which the rate constant was expressed as follows:
1063
This parameterization gives a lower coupling between the parameters E and k* than E and A in the normal Arrhenius equation. Note that k* is the rate constant at T*. The algorithm was used to fit kinetic constants to the pyrolysis of wheat straw at 5,lO and 4O"C/min (one data set per heating rate). The algorithm use the local temperature and does not rely on a constant heating rate. The data from an experiment were converted to dry ash free basis and the mass loss rate was normalized by the maximum mass loss rate. The data in the range where the normalized mass loss rate was above 0.1 was then used. This excludes the lignin tail from the data. The mass data were then converted to degree of conversion and normalized so the conversion of the final data point was 1.300 points were used per data set. Kinetic parameters were fitted to the individual data sets as well as to all three data sets simultaneously . The kinetic values are listed in Table 1.
Table 1 Kinetics derived for individual heating rates as well as for all three heating rates for the pyrolysis of wheat straw using a single first order reaction model. E (kJ/mol)
A (s-')
All
137.9
2.40.10'"
1.92.10-3
5 10 40
80.9 78.5 69.5
6.11.104 5.42.1O4 2.37.104
1.25*10-3 1.91 5.95.
Heating rate (OClmin)
k277-c (s-')
1
0.8
5Y? 0.6 0.4
0.2 0
200
240
200
320
360
400
Sample temperature ("C)
zoo
280 320 360 Sample temperature ("C)
240
400
(B)
(A)
Fig. 2 The pyrolysis of wheat straw at 5"C/min, 1 O"C/min and 4O"C/min.Single first order kinetics fitted to the individual curves (A) and simultaneously to all curves (B).
Table I shows that there is some difference between the kinetics fitted to the individual data sets. The activation energies for the pyrolysis at 5 and 10"C/min are close to 80
1064
kJ/mol, but the activationenergy for the pyrolysis at 40"C/min is approximately 10 kJ/mol lower. The conversion for all three data sets as well as the individual fits are presented in Fig. 2A. The model agrees reasonably well with the data. When the kinetics are fitted to all three data sets simultaneously, both the frequency factor and the activation energy increases dramatically. Fig. 2B shows the conversion for all three data sets as well as the kinetic from the simultaneous fit. The activation energy increases approximately by 60 kJ/mol to 140 kJ/mol and the frequency factor increases by a factor 105-106.The fit does not show as good agreement with the data as the individual fitted kinetics. In order to test the applicability of the model over a wider range of conditions the kinetics derived at lO"C/min (from Table 1) was used to predict the pyrolysis of wheat straw heated to 300°C at lO"C/min, then kept isothermal for 60 min, before heating to 600°C at 10"C/min, see Fig. 3. The experimental data show that the addition of a 60 min hold time at 300°C results in somewhat different mass loss curve. The initial mass loss follows that of the data heated directly to 600"C, but when 300°C is reached, the mass loss rate becomes much lower, and after 60 min at 300"C, the conversion is only about 0.8. Full conversion is reached when the temperature is raised to 600°C. The char yield for the two experimentswas approximately the same. The model, which fit the direct heating to 600°C reasonably well, does not work for the case with a 60 min isothermal segment at 300°C is inserted. The model predicted that the pyrolysis was far too fast, and a full conversion is predicted under the isothermal segment.This shows that first order kinetics only works well on the data on which it was fitted. A single first order model cannot predict the pyrolysis behavior under different conditions. 600
100
0
2000
4000 Time (s)
6000
8000
Fig. 3 The experimental data for the pyrolysis of wheat straw at 10"C/min to 600°C (#1) and to 300°C for 60 min before 600°C (#2).Model data fitted to #1 and used to model #2. Reynolds and Burnham [ l l ] found that a three parameter nucleation model fitted the pyrolysis of cellulose well. Here, the model was tested for the pyrolysis of wheat straw. The conversion rate was given by:
1065
dX = k(l -X)(1-0.99X)"'
(3)
dt
An algorithm similar to the one used for a single first order reaction was used to fit data to this model. The kinetics derived by Eqn. (3) is listed in Table 2 for fits to the same data as used to fit the single first order reaction. Fig. 4A and 4B show the fit to the individual curves and the fit to the curves simultaneously.The fits are reasonable good, but not better than the fits using a single first order reaction. In fact, the fit to all three heating rates simultaneously gave a fit with a value of m close to zero,in which case the model reduces to a first order reaction.
Table 2 Kinetics derived for individual heating rates as well as for all three heating rates for the pyrolysis of wheat straw using the model by three parameter nucleation model.
E (kJ/mol)
A (s-I)
All
148.2
2.22-10"
5 10 40
86.4 82.8 76.6
Heating rate ("C/min)
k,,., (s-'1
m
1.86,
4.1 *
2.39.10' 1.69.105 1.25*1OS
0.169 0.177 0.124
1.48*10-3 2.30*10-3 6.67-10"
1
0.8
.- 0.6 2
F
3 0.4 0.2 0 200
240
280
320
360
400
Temperature ("C)
200
240
280
320
360
400
Temperature ("C)
(B)
(A)
Fig.4 The pyrolysis of wheat straw at 5"C/min, 10"C/minand 40"C/min. Three parameter nucleation kinetics fitted to the individual curves (A) and simultaneously to all curves (B).
1066
Yet another model previously used to fit the pyrolysis of biomass is the distributed activation energy model. This model was originally used to model the pyrolysis of coal [ 121. Niksa and Lau [ 131 found that the distributed activation energy model was the best for coal pyrolysis and that it should be included as an essential element in any pyrolysis model. The pyrolysis is assumed to occur via a number of parallel first order reactions which has the same frequency factor but varying activation energy. The pyrolysis rate for the reactions is: dai
k,(l -ai) (4) dt in which the rate constant (ki) is given by an Arrhenius equation and the fraction of material which pyrolyses at a certain activation energy is given by a natural logarithmic distribution. In the original model of Anthony et al. [12] a normal distribution was used instead of the natural logarithmic distribution. - - -
f(E)
=
T)
1 1 exp[ -.-( 1 ln(E)-a
6 PE
*]
for which
7
f(E)dE
=
1
0
This results in the following expression for the conversion:
An algorithm to fit the experimental data to this model was made in Fortran using a least square method by Fletcher [lo]. In order to achieve the best result, the model was constructed to fit the frequency factor, A, and a, while the influence of p on the residual sum of squares had to be tested manually. As for the previous models, the model was used to fit data at a single heating rate and to fit data at three heating rates simultaneously.The best fit for the simultaneous fit to three heating rates was obtained with p = 0.055. This value was subsequently used to fit the data individually. Table 3 lists the kinetic parameters and Fig. 5A and 5B shows the fit to the data curves for fits to individual curves and simultaneous fit to all three curves. This model deviates from the previous described models by having frequency factors which are within a factor lo3 between the kinetics derived for individual data sets and simultaneously to all three data sets. Another feature for this model is that the fit to all three data sets simultaneously gives a reasonable result. This shows that this model is better equipped to handle different experimental conditions.
1067
Table 3. Kinetics derived for individual heating rates as well as simultaneouslyfor all three heating rates for the pyrolysis of wheat straw using the distributed activation energy model with p = 0.055.
A (d)
Heating rate ("Chin)
a
Em, (kJ/mol)
0
(kJ/mol)
all
2.56*10"
12.50
268.8
14766
5 10 40
6.07.10''
12.34 12.39 12.46
229.4 241.6 259.7
12627 13298 14296
8.33.10'' 4.57.10''
1
0.8
.-5 0.6 LI
s
8 0.4 0.2 0 200
250
300
350
400
Temperature ("C)
(A)
200
250 300 350 Temperature ("C)
400
(B)
Fig. 5 Pyrolysis of wheat straw at 5 , 10 and 40"C/min. Kinetics fitted to the individual curves (A) and simultaneously to three curves (B) by the distributed activation energy model.
To test the applicabilityof the model, the kinetics derived for the pyrolysis of wheat straw using the kinetics derived at 10"C/min (from Table 3) was used to predict the pyrolysis of wheat straw when the heating was delayed by inserting a 60 min isothermal segment at 300°C in the temperature program. Fig. 6 shows the experiments and the fits. Compared to the other models the distributed activation energy model performs better. It must, however, be remarked that given enough time at 300°C, also this model would predict full conversion.
1068
0
4000
2000
Time
6000
8000
(s)
Fig. 6 The experimental data for the pyrolysis of wheat straw at 10"C/min to 600°C (#I) and to 300°C for 60 min before 600°C (#2). Model data using the distributed activation energy fitted to #1 and used to model #2. The pyrolysis of wheat straw could be expected to be described by a superposition model, i.e. described by a number of parallel first order reactions, representingthe decomposition of the biomass constituents (cellulose, hemicellulose and lignin) [4,5]. In this work, a superposition model has been used, in which the pyrolysis is assumed to be described by N independent first order reactions (for i=l, 2, ..., N):
where X is the total conversion and the conversion rate is given by:
The reacted fraction of component i is then given by: a. =
'
- mi
%,i l%,i
(10)
-mchar,i
The components are assumed to decompose individuallyaccording to first order kinetics:
_da,_
- ki(l -a,) dt where the rate constant is given by the Arrhenius equation:
ki
=
Aiexp[
-+]
The coefficient ci expresses the contribution of the partial process to the overall process:
1069
c. =
%i
- mchar,i
mo
The set of equations can be fitted to one data set (either TGA or DTG) by the least square method. An algorithmhas been developed by Austegaard [ 141and programmed in Matlab. Further information regarding the algorithm can be found in Grcdnli [ 151. When the superposition model is used to model the pyrolysis of wheat straw, difficulties may arise because the mass loss rate curve shows an asymmetricalpeak with a long tail. The absence of several distinct peaks makes the fit difficult. It is possible to achieve a good mathematical fit; however, the fit may not have a physicakhemical meaning. This may limit the fitted kinetics to be valid only at the conditions of the experiment they were fitted to. The pyrolysis of wheat straw at 10"C/minhas been fitted to both mass loss data and mass loss rate data using both two and three parallel reactions. The kinetic parameters are listed in Table 4. All the fits describe the pyrolysis well, however, there is a large difference between the kinetic constants. Fig. 7 shows the fit to the mass loss rate using three reactions. The activation energies are very different depending on whether they are determined from mass loss or mass loss rate data. This indicates that the values of activation energy are not well determined by this method and the values of all the parameters may not have a physical meaning. 1
0.000030
0.9 v)
0.000025
0.8
2
0.7
.-2
0.6
0.000020 $2
! I
'0
v)
0.000015
m
3
0.5
z
4
0.000010 5
0.4 0.000005
0.3
0.2 0.000000 150 200 250 300 350 400 450 500 550 Temperature ("C)
Fig. 7The pyrolysis of wheat straw at 10"C/min fitted by three parallel reactions using the mass loss rate data.
The model contains many more parameters than the single first order reaction, and so it can only be justified by a wider range of applicability. To test this, the kinetics derived for the pyrolysis of wheat straw using three reactions and based on mass loss rate was used to predict the pyrolysis of wheat straw when the heating was delayed by inserting a 60 min isothermal segment at 300°C in the temperature program. Fig. 8 shows the experiments and the fits, and it is clear that the superposition model does not predict the case with the isothermal segment well. It is, however, slightly better than a single first order reaction (Fig. 3) due to the slower kinetics of the lignin tail.
1070
Table 4 Kinetics for the pyrolysis of wheat straw at 10"C/minusing 2 and 3 parallel reactions (for both mass and mass loss rate data).
El (kJ/mol) E& (kJ/mol) E3 (kJ/mol)
2 reactions (mass)
2 reactions (mass loss rate)
3 reactions (mass)
3 reactions (mass loss rate)
23.8 54.4
27 51.2
24.2 44.5 9.14
28.5 34.4 14.4
22 108.9
20.6 117.8
23.8 139.1 114
24.5 181.1 139.8
8.16.10.' 3.80.10'
6.96.10.' 4.13.10'
1.24.10.' 1.96-10'0 5.94.10'
1.78.10-' 1 . 1 6*1014 1.34.10"
600 500 400
0
-$ c
Ea
300 a
I-
200 100
0
2000
4000 Time (s)
6000
8000
Fig.8 The experimental data for the pyrolysis of wheat straw at lO"C/minto 600°C (#1) and to 300°C for 60 min before 600°C (#2). Model data fitted to #1 and used to model #2 using the kinetics derived from the superposition model with three reactions (based on mass loss rate data). The superposition model was also applied for the pyrolysis of washed wheat straw, where a multi-reaction model is required to fit the data. Experiments show that the DTG-curves have two distinct peaks, which makes the kinetics determinationeasier than for raw straw. Fig. 9 shows the fit of three reactions to the mass loss rate data from the pyrolysis of
1071
washed wheat straw at lO"C/min. The fits show good agreement with the experimental data. The kinetic data are listed in Table 5 . As for the pyrolysis of wheat straw, it may be expected that these kinetic values will fail to predict the pyrolysis at different conditions.
Table 5 Kinetics for the pyrolysis of washed wheat straw at 10"C/min 3 parallel reactions (for both mass and mass loss rate data). 3 reactions (mass)
3 reactions (mass loss rate)
7.79 44.6 32.7
6.57 46.2 31.6
68 189.8 114.4
90.4 172.7 111.6
El (kJ/mol) E, (kJ/mol) E, (kJ/mol)
4.67.10' 4.33.1013 1.77.1Ox
3.33.1O4 1.59.1012 1.03.1Ox
0.002
1
0.8
- - - Reaction 2
0.0016
v)
I
al
c
0.0012 !!
E 0.6 73
v) v)
-0
.d 0.4
0.0008
8 0.2 0 200
I
0.0004 0
250
300
350
400
450
500
Temperature ("C)
Fig. 9 The pyrolysis of washed wheat straw at 10"C/min fitted by three parallel reactions using the mass loss rate data.
1072
The model for straw pyrolysis by Lanzetta and Di Blasi [7] was applied to the data in this study even though the conditions of the TGA-experiments are very different from the conditions used to fit the parameters in the model. The heating rate in the present study (5-4OoC/min)is much lower than in the study by Lanzetta and Di Blasi [7], who used heating rates of 1500-420OoC/min,and the maximum final temperature was higher in this study (700°C) compared to 375°C by Lanzetta and Di Blasi [7]. The reaction scheme used by Lanzetta and Di Blasi [7] is given in Fig. 10.
Fig. 10 Reaction scheme for the model by Lanzetta and Di Blasi [7]. Fig. 11 shows the experimentaldata from this study and results from simulationsusing the kinetic model by Lanzetta and Di Blasi [7] at 10"C/min. Only the normalized mass and mass loss rate are shown from the model. The model predicts that the pyrolysis starts at lower temperatures (below 150°C) than in the experiment and that the char yield is 65% compared to the experimentally determined char yield of 25%. This shows that the model cannot be extrapolated to work at TGA-conditions.This is an inherent problem of parallel reaction schemes.
- ..
1
-'"..:.,.."x
..
0.90.8-
2rb
x
'. ....
x xx"
....
0.7-
'cxx
h
......Yx......x.................................
E 0.6-
-0.002
x
-2 0.5.-
X
x X
- 0.0015 2
#
0.4-
2 0.3-
-0.001
,.'.....
0.2-
/ x.., . x ..;: .f
0.1 --
../$X
x x x x x x x-
-0 3 v)
5
- 0.0005
X
...
c
a c !!
X
-0
-(?
xX
....jar ........ I I I 1 I I I 0 100150200250300350400450500550 Temperature ("C) g..i
1
I
I
I
I
I
I
Fig. 11 Modelling of wheat straw pyrolysis at 10"C/min using the model by Lanzetta and Di Blasi [7].
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CONCLUSION Pyrolysis of wheat straw and washed wheat straw has been carried out by simultaneous thermogravimetric analysis and differential scanning calorimetry at heating rates of 5 , 10 and 40"C/min. Simple kinetic models have been tested against the weight loss and weight loss rate data from the pyrolysis experiments. The models include a single first order reaction, a three parameter nucleation model, a distributed activation energy model using a natural logarithmdistribution and a superpositionmodel based on first order reactions. All of the models were used to fit an individual set of experimental data, while some were also used to fit a set of experimental data with different heating rates. All the models could fit a single experimental data set reasonably well. However, a simultaneous fit to several experimentaldata sets at different heating rates did not give good results for a single first order reaction and the three parameter nucleation model. These models also failed to predict the pyrolysis behavior when an isothermal segment was introduced during the heating. The superpositionmodel also failed to predict this. None of the models, however, could be extrapolated safely to conditions very different from those they were fitted to. The distributed activation energy model performs better for a set of experimental data with different heating rates. This model could also show the correct trend for the mass loss curve when an isothermal segment at 300°C was introduced during the pyrolysis of wheat straw at 10"C/min to 600°C.The distributed activation energy model gave the overall best results for the pyrolysis of wheat straw. In addition a two-step model describing the pyrolysis of straw at high heating rates was extrapolated to predict the thermogravimetric data. The model did not work well at these heating rates. A widely applicable biomass pyrolysis model is still not available in the literature. ACKNOWLEDGEMENTS This work is part of the CHEC research program (Combustion and Harmful Emission Control) funded by the Technical University of Denmark, the Danish Technical Research Council, Elsam, Elkraft, the Danish Energy Research Program, the Nordic Energy Research Program and the European Union. NOMENCLATURE frequency factor (s-') contribution of the partial process to the overall process (kgkg) activation energy (J/mol) rate constant (s-') parameter in the three parameter nucleation model (-) initial mass of component i (kg) mass of char for component i (kg) mass of component i (kg) gas constant (8.314 J/molK) time (min) temperature ("C or K) temperature of maximum mass loss rate ("C or K) overall conversion (kgkg)
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Greek a parameter in the distributed activation energy model (-) degree of conversion for component i (kgkg) Qi parameter in the distributed activation energy model (-) P REFERENCES
1. Jensen, A., Dam-Johansen, K.,Wcjjtowicz, M.A., Serio, M.A. (1997). Energy & Fuels, 12,929-938. 2. Kobs, M., Rep& M., Kosik, M., Reiser, V., Mihglov, V., Ciha, M. (1983). Chem Zvesti, 39, 399-408. 3. Raveendran, K., Ganesh, A,, Khilar, K.C. (1995). Fuel, 74, 1812-1822. 4. Raveendran, K., Ganesh, A., Khilar, K.C. (1996). Fuel, 75, 987-998. 5. Varhegyi, G., Szabo, P., Antal Jr., M.J. (1994). In A.V. Bridgwater (Ed.), Advances in thermochemical biomass conversion. (pp. 760-770). Blackie Academic & Professional. 6. Knudsen, N.O., Jensen, P.A., Sander, B. & Dam-Johansen, K. (1998) Proceedings of the 1Olh European Conference and TechnologyExhibition, Wurzburg, Germany, June 8- 1 1, pp. 224-228. 7. Lanzetta, M. and Di Blasi, C. (1998). Journal ofAnalytica1andAppliedPyrolysis, 44, 181- 192. 8. Stenseng, M., Jensen, A., Dam-Johansen, K.& Grflnli,M.(1999) Proceedings to the 2" Olle Lindstrom Symposiumon renewable Energy, Bioenergy, Stockholm,June 811, pp. 97-104 9. Stenseng, M., Jensen, A. & Dam-Johansen, K. (2000) submitted to Journal of Analytical and Applied Pyrolysis 10. R. Fletcher (1971) A modified marquardt subroutine for non-linear least squares, Harwell Report, AERE R. 6799. 11. Reynolds, J.G. & Burnham, A.K. (1997) Energy and Fuels, 11, 88-97. 12. Anthony, D.B., Howard, J.B., Hotel, H.C. & Meissner, H.P. (1975) Fifteenth Symposium(International)on CombustiodI'he Combustion Institute, Pittsburgh, Pa, 1303-1317. 13. Niksa, S. & Lau, C.-W. (1993) Combustion and Flame, 94, 293-307. 14. Austegaard (1997) Experimental and numerical study of a je@re stop material and a new helical flow heat exchanger. DPhil thesis, The Norwegian University of Science and Technology. 15. Grgnli, M. (1996) A theoretical and experimental study of the thermal degradation of biomass. DPhil thesis, The Norwegian University of Science and Technology.
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The potential of multivariate regression in determining formal kinetics of biomass pyrolysis S. Volker, 42 Engineering, von-Behring-Str. 9, 0-34260 Kaufiitgen, Germany
Th.Rieckmann, University of Applied Sciences Cologne, Department of Chemical Engineering and Plant Design, Betzdorfer Str. 2,D-50679 Cologne, Germany
ABSTRACT.- Multivariate regression has been applied to derive formal kinetic models for the mass loss of beech wood and wood componentsduring pyrolysis. STA experimentshave been performed with constant heating rates between 0.14 and 108 Wmin. It was demonstrated that our experimentalconditions provided (1) reliable process temperatures, (2) linear increase of temperature with time, and (3) negligible temperaturegradients inside small samples. This allows the simultaneousfitting of a kinetic model with a single set of kinetic parameters to experiments at all investigated heating rates. Multivariate regression is an important tool for deriving reliable kinetic models with validity over the investigated parameter range. Using this approach, an appropriate model representing the best compromise between accuracy and number of parameters can be selected for the desired engineering application.
INTRODUCTION
The demand for sustainable development has led to growing interest in renewable energy and chemical feedstocks. Biomass is a source with global availability and its thermal conversion has attained a lot of attention during the past decades. The pyrolysis of biomass to yield valuable liquid and solid products is now on its way to commercialization.The development and optimization of technical processes bases on data from laboratory equipment as well as mini plants and scale-up is supported by process simulation. To be successful, this task needs reliable reaction models for the desired operation conditions. Since the molecular mechanisms during the thermal conversion of biomass are too complex to be modelled in full detail, reduced formal reaction models are used for engineering purposes which are characterized by lumping of compounds and mechanisms with limited numbers of reaction steps. Nevertheless, these formal reaction models have to comprise enough parameters to satisfy the requirement of validity over the whole technical parameter range to be evaluated and the reaction kinetics have to be separable fiom heat and mass transport influences. Although the thermal decomposition reactions of biomass have been systematically investigated for many years the biomass community is still debating the best formal reaction model and kinetic parameters for the prirnary thermal decomposition. Even for 1076
the chemically well defined component cellulose, the proposed models 4-10 include onestep frst order models postulating or rejecting a change of the rate limiting reaction step going fiom low to high heating rates, as well as multi-step models with competitive and series reactions leading to volatiles and char. Thermogravimetric analysis is a common tool for deriving formal reaction models for low heating rates which are relevant for e.g. carbonization processes. TGA experiments are usually performed in a dynamic mode with constant heating rates between around 100 Wmin and less than 1 Wmin. Although the simultaneous evaluation of results fiom experiments with different heating rates according to the technique of multivariate regression has been strongly recommended 11*'2, this approach is still not commonly applied in publications on biomass pyrolysis. This might be due to the general assumption that TGA experiments with high heating rates cannot be performed without a thermal lag between the sample and the thermocouple measuring the process temperature lo. Therefore in this study, conditions for TGA experiments have been investigated under which the process temperature is reliable and the temperature gradient inside the sample is negligible. The results of those experiments allow the fitting of formal kinetic models to the experimental mass loss curves according to the technique of multivariate regression. The intention was to work out guidelines for the development of formal reaction models with validity over a broad range of operation conditions which can be applied for design and scale-up of biomass reactors. EXPERLMANTAL AND MODELLING THERMOGRAVIMETRICANAL YSIS
Experiments have been performed with a combined TGA/DTA (STA 503, Buhr GmbH, Altendorfer Sir. 12, 0-32609 Hiillhorst) applying constant heating rates p between 0.14 and 108 Wmin. The maximum load of the STA is 1000 mg, and mass variations o f f 200 mg can be detected. The resolution is 1 pg. The STA apparatus has a horizontal weighing beam which generates an unusually low drag in the weighing direction and the effect of drag was further reduced by using high purity helium (99.999 %, Linde) as a low viscosity purge gas with a gas flow rate of 1.7 Lh. The thermocouples measuring the sample temperature and the reference temperature, respectively are integrated in the sample holder and are located directly beneath the platinum plates under the crucibles. This configuration provides an intense and reproducible thermal coupling to the sample. A temperature calibration has been performed by analyzing the DTA signals from the melting peaks of the pure substances In (T, = 429.8 K); KC104(T, = 572.6 K); and Ag2S04 (T, = 699.6 K). This temperature calibration has been performed for all heating rates, using open crucibles made 6om A1203and the same conditions as in the experimental runs. Therefore, the corrected temperatures of the thermogravimetric analyses represent the sample surface temperatures well. The linearity of the T(t) curves was excellent for all heating rates. Linear regression of the T(t) curves to estimate the individual heating rate for each experiment resulted in correlation coefficients between 0.999997 for p = 0.14 Wmin and 0.999832 for p = 108 Wmin. Samples from beech wood as well as from biomass components have been investigated using the same crucibles as for the temperature calibration. Microcrystalline cellulose (Avicel, Merck); lignin (Organosolv, Aldrich); and xylan
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fiom birchwood (extracted with NaOH, max. 5 % sulfate ash, Roth) have been chosen as biomass components. Samples from beech wood (Riiuchergold, J. Rettenmaier & Siihne) were prepared by grinding the coarse particles and separating the fiation with particle diameters < 0.5 mm. Additionally, this fraction was extracted two times with hot water (335 K) to remove water soluble minerals. The beech wood contained 46.4 ‘YO cellulose; 22.3 % hemicellulose; 25.4 % lignin; 5.1 % water solubles; and 0.7 % ash. Initial sample masses m,, between 1 and 20 mg have been used in the investigation of cellulose. The experiments with beech wood, xylan, and lignin were performed with mo = 3 mg. All samples were dried 1.5 h at 38 1 K in a helium stream at the beginning of the experiments.
MODELLING OF THE SAMPLE TEMPERATUREDISTHBUTION Derivation of kinetic models from TGA analyses bases on the assumption of a homogeneous temperature throughout the sample. To check this assumption, the spatial temperature distribution in the sample has been modelled by a simplified heat transport reaction model. The modelling has been performed for cellulose, because its heat demand per time interval is the highest of the investigated biomass samples due to its high reaction rate and endothermic reaction enthalpy. The crucible contents of the TGA were modelled as a dynamic system with distributed parameters assuming plate geometry. The radius of the sample was taken as the characteristic length for the heat transport. The overall rate of reaction was approximated by an irreversible reaction of 1”order: biomass + solids + volatiles, assuming the validity of the Arrhenius expression for the temperature dependency of the rate constant.
The kinetic parameters were derived from the simultaneous evaluation of two TGA experiments with heating rates of 0.14 and 0.5 Wmin, which can be assumed to be independent of heat transport influences. The material balance and the enthalpy balance resulted in a set of two meshed partial differential equations (PDE).
Material balance
As a worst case scenario, the pyrolysis of cellulose was assumed to be endothermic during the entire decomposition with a reaction enthalpy of AHR= 340 kJ/mol. This is the most endothermic value measured for cellulose under vacuum by Mok and Antal l3 in experiments with open sample holders. The measured reaction enthalpy comprises the vaporization of volatiles and represents the total heat demand of the pyrolysis process.
Enthalpy balance
ar a2T (1 - % )Poc,, -= 3L-+(-ARH) at
ax2
r(p, T)
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(3)
Initial conditwns T(t = 0) = TO
P(t = 0)= Po
Boundary condilions
gl
(left boundary)
=0
T(x=h(t) = Pt + To
f% x=o
The temperature Tx=h(t)at the right boundary is the corrected temperature representing the sample surface temperature (as described above). Therefore, external heat transfer limitations between STA oven and crucible had not to be considered. The PDE system was transposed into a set of 2-n ordinary differential equations (ODE) by the finite differences technique (calculated with n = 128 finite differences).
for j
= 2, n-1
finite element.
At the left boundary (i = 1)
8 ’ ~ (-Tj + Tj+l) dx2 Ax2
(9)
-5%
and at the right boundary (i = n)
The simultaneous numerical solution of the ODE system was performed using Matlab (The MufhWorks).The kinetic and model parameters are summarized in Table 1. Table I Parameters for the heat transport reaction model.
I
14.72
this work
197.5 kJ/mol
this work
340 k J k g
reference 13
h
2.5 kJkgK 0.243 W/m/K
reference 14 reference 14
Po
1500 kg/m3
Merck
Eo
I
0.7
I
Merck
I
FORMAL KINETIC MODELS
Formal kinetic models have been derived to describe the experimental mass loss of the biomass samples. The material balance of the crucible contents was described as a dynamic system with concentrated parameters which resulted in a single or a set of ordinary differential equations (ODE). The corrected sample temperature Tx=h(t)was
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used for the time/temperature integration at a constant heating rate p which was calculated 6om the experimental data. The simultaneous numerical solution of the ODE system and the estimation of the kinetic parameters by the least squares (LSQ) method were performed using the software package Thermokinetics (Netzsch). The kinetic parameters according to different reaction models were calculated by fitting single experiments as well as by the multivariate regression technique I*. TGA curves were used for the least-squares analysis because the TGA curve is less sensitive to noise and round-off error than the DTG curve. Additionally, this method uses the important data obtainable 6om the start and finish of the TGA curve which helps selecting an appropriate model Is. RESULTS AND DISCUSSION
SAMPLE TEMPERATURE DISTRIBUTION
The modelling results of the temperature distribution in cellulose samples during TGA experiments are displayed in Figures 1 - 3. At a constant heating rate of 108 Wmin, the temperature in the interior of a cellulose sample of 1 mg follows the temperature of the sample surface closely with a maximum temperature difference of less than 0.15 K between the different spatial layers of the cellulose sample (Fig. 1). Under the same conditions, the spatial layers of a cellulose sample of 3 mg exhibit. a maximum temperature difference of 0.8 K (Fig. 2). This is mainly due to unsatisfied heat demands of the endothermic reaction, since a similar calculation with ARH= 0 kJ/mol yielded a maximum temperature difference of 0.2 K between the different spatial layers of the cellulose sample. In a cellulose sample of 20mg at a constant heating rate of 108 K/min, a maximum temperature difference of nearly 25 K is predicted between the different spatial layers of the cellulose sample (Fig. 3). It is evident that such sample sizes are unsuitable for kinetic analyses applying the assumption of a homogeneous sample temperature, as stated in the literature before lo.
675.0,
700 750 temperatwe sample surFace / K
650
=
I
800
Fig. 1 Calculated temperature distribution for the different spatial layers in a cellulose sample of% = 1 mg at p = 108 Wmin according to the heat transport reaction model.
1080
600
. . . .
I
650
.
.
676
675
614
600-
.
'
700
I
750
'
.
*
.
$
0
temperatme sample surface / K
Fig. 2 Calculated temperature distribution for the different spatial layers in a cellulose sample of Q = 3 mg at 0 = 108 Wmin according to the heat transport reaction model. 750
I
700-i
I
I
1
Fig. 3 Calculated temperature distribution for the different spatial layers in a cellulose sample of Q = 20 mg at p = 108 Wmin according to the heat transport reaction model.
In Figure 4, the calculated mass losses for cellulose at different constant heating rates and initial sample masses are compared to experimental TGA results. The TGA curves at heating rates of 0.14 Wmin and 0.5 Wmin had been used to evaluate the kinetic parameters for the one step fmt order reaction model which was incorporated into the model to calculate the sample temperature distribution. Since the temperature gradients in those samples are nearly zero, the results of the heat transport reaction model represent simultaneously the best fit for the assumed reaction model. At a heating rate of 108 K/min, the initial sample mass influences the temperature at which a given mass loss is attained. Cellulose samples with IQ = I - 3 mg are affected only to a minor 1081
extend by heat transport phenomena, and the heat transport reaction model represents the experiments very well. The temperature shift of the curve for 3 mg compared to the curve for 1 mg is less than 2 K. In a cellulose sample of 20 mg, the temperature inhomogeneity is pronounced and generates a significant temperature shift of the mass loss curve which is underestimated by our simplified model. This is due to mass transport effects which can not be neglected for a comprehensive evaluation l6 but are not included in our model. The results c o n f m that TGA experiments are not significantly affected by heat transport phenomena if low initial sample masses as well as the described TGA configuration and experimental procedures are applied. The temperature gradients inside the samples are sufficiently small to allow the fitting of formal kinetic models to the experimental mass loss curves assuming a homogeneous sample temperature. Cellulose samples with initial sample masses of around 5 mg and higher can only be submitted to kinetic analyses under consideration of the enthalpy balance.
Fig. 4 Modelled mass losses (lines) for cellulose at different heating rates and initial sample masses and comparison to TGA results (symbols).
To minimize heat transport influences on our kinetic analyses, we have chosen an initial sample mass of 1 mg for cellulose at heating rates of 41 and 108 Wmin. For cellulose at low heating rates as well as for all runs with xylan, lignin, and beech wood, initial sample masses of 3 mg have been chosen to improve the signal / noise ratio. The use of 3 mg samples at the high heating rates was justified by the observation of the DTA signals which showed that the endothermic heat of reaction for beech wood pyrolysis is at least four times smaller than that for cellulose pyrolysis. According to our heat transport reaction model, a beech wood sample with m o = 3 mg at p = 108 Wmin would exhibit a maximum temperature difference of 0.4 K between the different spatial layers of the sample. The pyrolysis of lignin showed almost no overall reaction enthalpy whereas the pyrolysis of xylan was very slightly exothermic. The temperature gradients inside the samples can be neglected for kinetic analysis.
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KINETIC ANA YLSIS A vice1 cellulose
The kinetic analysis of cellulose pyrolysis is extensively described in the literature with varying results regarding the proposed formal reaction mechanisms as well as kinetic parameters for identical reaction models4-". One of the most widely used reaction models is an irreversible 1" order single step reaction. This model is grossly simplified with the major drawback, that it is unsuitable to describe the last stages of the pyrolysis process as well as the different char yields observed for different heating rates. Another disadvantage of this model is that evaluations of data sets fiom single experiments result in kinetic parameters varying over a broad range due to a compensation effect 16," which compensates influences of heat and mass transport and the inadequacy of the one step model. A further problem with evaluating data sets of single experiments is that the formal reaction models contain enough parameters to achieve an equally good fit for a great number of model types. The choice of an appropriate reaction model is therefore complicated considerably. The technique of multivariate regression requires the identity of kinetic parameters for all data sets submitted to the analysis. This provides a very strong criterion and the number of applicable models as well as the range of resulting kinetic parameters narrows substantially. To demonstrate this, kinetic analyses basing on a single experiment with p = 3 Wmin have been compared to kinetic analyses basing on the simultaneous evaluation of the data sets for all five heating rates. A single step reaction has been chosen for the mechanism and different reaction types have been tested. The results are summarized in Table 2. The kinetic analyses of the single experiment yielded best fit models describing the experimental mass loss well with correlation coefficients between 0.99929 and 0.99993. Although each of the models could be used as formal reaction model for a process with a constant heating rate of 3 Wmin, these models would predict widely differing behaviour for other heating rates due to their activation energies lying between 12 and 552 kJ/mol. On the other hand, the temperature shift of the mass loss with heating rate determines the activation energy resulting fiom the multivariate regression. This restriction reduces the number of formal reaction models fiom nine to six which describe the experimental mass loss well with correlation coefficients > 0.999. It is especially noteworthy that the kinetic parameters for those six models fall in a very narrow range with loglo (k&-') varying between 14.69 and 15.08 and EA varying between 197.8 and 198.7 kJ/mol. Kinetic parameters for the model with a single step first order reaction have also been calculated by a multivariate regression analysis of the DTG results for all heating rates. The best fit kinetic parameters are logLo(k,-,/s-') = 14.75 and EA= 196.7 kJ/mol demonstrating that the result for the kinetic parameters is not influenced by choosing either the TGA or the DTG curves for the multivariate regression. The comparison of the results of a round-robin study published by Gmnli et al. Is to our results for p = 40 Wmin (the only heating rate investigated in both studies) shows that the correspondence between the experimental mass loss curves is very good. Nevertheless, the activation energy of 198.4 kJ/mol for the ifreversible Is' order reaction fiom this work is considerably lower than the 21 1 - 232 kJ/mol found for the results fiom the different laboratories participating in the round-robin study. Instead,
1083
our results correspond better to an investigation performed by Reynolds and Bumham '. Table 2 Comparison of results fiom the evaluation of a single experiment with results fiom a multivariate regression for different formal kinetic models
I
I Evaluation of single experiment I
Reaction type
~~~~
Multivariate regression
I
~
lstorder reaction
I nthorder reaction
1
log10
EA/
(k,,/s-') 16.87
kf/rnol 220.6
reaction order 1
cc: 0.99968;mr: 0.02201 21.30 I 268.8 I 1.52 cc: 0.99989; mr: 0.01256
3D Jander's diffusion
31.39
3D phase boundary reaction
13.90
1 order
- 1.19 I
but-Tompkins
396.7
cc: 0.99929; mr: 0.03 156
I
193.6
I I
(k&)
14.85
I I
1
cc: 0.99989; mr: 0.01230
reaction order
198.4
1
cc: 0.99927; mr: 0.03035 14.89 1 198.7 1 1.05
196.2
cc: 0.99860; mr: 0.04193
15.78
I
195.3
I
1
cc: 0.96431; mr: 0.21006
I
21.29
1* order autocatalytic reaction
16.87
cc: 0.99968; mr: 0.021 17
cc: 0.99924; mr: 0.03080
nth order autocatalytic reaction
15.42 206.2 1.65 cc: 0.99993; mr: 0.00982
n dimensional nucleation Avrami-Erofeev
46.23
198.0 1.68 14.69 cc: 0.99968; mr:0.02009 14.80 197.8 Dim.: 1.04
1.52
cc: 0.99989; mr: 0.01255
I
220.6
552.1
I
1
Dim.: 0.41
cc: 0.99968; mr: 0.021 14
I
15.37 217.7 cc: 0.98784; mr: 0.12352
nthorder autocatalytic reaction Prout-Tornpkins
268.6
I
cc: 0.99928; mr: 0.03012
14.11
cc: 0.99929; mr: 0.03161
11.9
EA/ kJ/mol
loglo
15.08
198.2
1.46
cc: 0.99958; mr: 0.02302
14.77
197.8
I
1
I
cc: 0.99930; mr: 0.02970
To improve the formal reaction model, multi-step models have been tested by multivariate regression with data sets fiom all heating rates. The pyrolysis of Avicel cellulose can best be represented by a formal reaction model comprising three reaction steps of n"order with two parallel reactions, one of them being a series reaction. Models with reactions (1) and (2) being competitive reactions yielded worse fits than models with parallel reactions. This might be due to different reaction pathways for crystalline and non-crystalline cellulose or an undetected formation of anhydrocellulose. However, for a mechanistic investigation a broader data base would be required. The kinetic parameters for the best fit model A&B
C Z D ~ E
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are loglo06/s-') ( I ) = 14.97, EA(1) = 198.7 W/mol, n( 1) = 0.95 ; loglo0ds-I) (2) = 1 1.44, EA (2) = 157.9 kJ/mol, n(2) = 0.95 ; and log,, (k,Js-') (3) = 35.96, EA (3) = 413.5 kJ/mol, n(3) = 4.60 . The total mass loss is the s u m of the individual mass losses from the three reactions. Reaction (1) contributes 62.3 %, reaction (2) contributes 4.5 YOand reaction (3) contributes 33.2 % to the total mass loss. The best fit model has a correlation coefficient of 0.99993, a mean of residues of 0.00966 and is displayed in Figure 5 together with the experimental results and the best fit for a single step 1" order reaction. The three step model fits nearly perfectly and could be interpreted as comprising a single tar forming reaction (1) and a serial char + gas forming reaction (2 and 3), as proposed by VArhegyi '. Reaction (1) has similar kinetic parameters as those resulting 6om the best fit of a single step mechanism, indicating that this reaction dominates the pyrolysis process of cellulose. The activation energy EA (1) matches the value given by Bradbury et al. for the volatilization step during cellulose pyrolysis 19. Reaction (2) has an activation energy of 157.9 kJ/mol, comparing well to data fiom the literature ',*'. The kinetic parameters for reaction (3) can be interpreted as covering a great number of parallel reactions taking place during the carbonization of the primary char.
temperature / K
Fig. 5 TGA results for cellulose at different heating rates and best fit models for a single step mechanism ( - - - ) and a three step mechanism ( -). The reaction order of the first two reactions is 0.95, indicating that their reaction profile is narrower than that of a first order reaction. This could be a hint for a serial reaction or the presence of a preceding activation step respectively. A multivariate regression fitting a four step mechanism to our experimental data did not improve the fit significantly due to the already high accuracy of the three step mechanism. To verify the presence or absence of an activation step, additional experiments under differing conditions have to be performed. The kinetic parameters fiom this work are valid under our investigated experimental conditions. They will probably change if the model is fitted to experiments with differing celluloses or experimental conditions due to influences of catalysts on reaction (1) and influences of mass transport on reactions (2) and (3). Especially mass transport plays a very important role in the pyrolysis of cellulose. This item demands further research to enable the evaluation of formal reaction models with improved transferability.
',
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Xylan from birch wood Xylan fiom birchwood decomposes over a broader temperature range than cellulose and the shape of the TGA curves suggests a mechanism with more than one reaction step. Consequently, one step formal models are not very successful. Applying multivariate regression, the best fit for this case is achieved with a single reaction of n* order with the kinetic parameters log,, (kds-')= 19.63, EA = 220.8 kJ/mol and n = 5.45. A correlation coefficient of 0.99423 and a mean of residues of 0.06806 were calculated for this fit. To represent xylan pyrolysis under our experimental conditions, a formal reaction model with three reactions in series was most successful:
A - L B ~ C ~ D The following reaction types were used: A 2-dimensional diffision reaction for reaction (1); a 3-dimensional diffusion reaction of the Jander's type for reaction (2); and a reaction of n* order for reaction (3). The best fit model has a correlation coefficient of 0.99955, a mean of residues of 0.0191 1, and the kinetic parameters fiom a multivariate regression are log,, (kds-I)(1) = 19.14, EA(1) = 203.5 kJ/mol; loglo (kds-I)(2) = 16.0, EA (2) = 203.2 kJ/mol; and loglo Ob/s-') (3) = 54.61, EA(3) = 625.4 kJ/mol, n(3) = 15.02. Reaction (1) contributes 10.6 %, reaction (2) contributes \ 61.6 % and reaction (3) contributes 27.8 % to the total mass loss. The best fits for the one step and the three step formal reaction model are displayed in Figure 6. Formal reaction models with parallel reactions always yielded worse fits than equivalent models with series reactions. This confirms results !?om an investigation of Vhrhegyi et al. on the thermal decomposition of 4-methyl-DglUCtUOnO-D-Xylan 21.
0 -10
s .
-20
8
-30 b3
3 -40 -50 -60 400
450
500
550
600
650
700
7
temperatureJK
Fig.6 TGA results for xylan at different heating rates and best fit models for a single step mechanism ( - - - ) and a three step mechanism ( -).
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Organosoh,lignin Organosolv lignin decomposes over an even broader temperature range than xylan. A formal reaction model with a single reaction step gives only a rough estimate for the pyrolysis behaviour of lignin. The best reaction type for a one step mechanism is the n-dimensional nucleation reaction (Avrami-Erofeev), especially for higher heating rates. Applying multivariate regression, the best fit is achieved with the kinetic parameters loglo0ds-I) = 11.99, EA = 174.4 kJ/mol and a dimension of 0.34. The correlation coefficient for this fit is 0.98833 and the mean of residues is 0.08904. To represent lignin pyrolysis under our experimental conditions, a formal reaction model with two parallel series reactions had to be set up:
AAB-L'C D
~
E
~
F
The following reaction types were used: A 3-dimensional diffusion reaction of the Jander's type for reactions (1); (2); and (3) and an n-dimensional nucleation reaction (Avrami-Erofeev) for reaction (4). The best fit model has a correlation coefficient of 0.99972, a mean of residues of 0.01362, and the kinetic parameters from multivariate regression are loglo(kds-') (1) = 5.49, EA(1) = 95.4 kJ/mol; loglo(kds-I)(2) = 15.92, EA (2) = 236.8 kJ/mol; loglo(kds-')(3) = 8.44, EA(3) = 138.5 kJ/mol; and loglo0ds-I) (4) = 46.22, EA (4) = 635.1 kJ/mol, dim. = 0.09. Reaction (1) contributes 14.6 %, reaction (2) contributes 13.5 %, reaction (3) contributes 23.2 % and reaction (4) contributes 48.7Xto the total mass loss. The best fits for the one step and the four step formal reaction model are displayed in Figure 7.
Fig. 7 TGA results for lignin at different heating rates and best fit models for a single step mechanism ( - - - ) and a four step mechanism ( -).
Water washed beech wood The pyrolysis of water washed beech wood can only be described by multi-step formal models. The most successlid approach is a mechanism with three parallel reactions, as commonly used throughout the literature: 1087
ALB C A D
ELF A formal reaction mechanism with three 1" order reactions describes the experimental results very satisfyingly over the whole temperature range in case the experiments at different heating rates are evaluated separately. For each single heating rate, kinetic analysis results in kinetic parameters which can not be applied to describe the pyrolysis behaviour at any other heating rate. If multivariate regression is applied to TGA results from all heating rates simultaneously, the main parts of the pyrolysis process are still described very well, but the slow mass loss at the final stages of the decomposition can not be re resented. The kinetic parameters for this formal reaction model are loglo(k&- ) (1) - 15.62, EA (1) = 188.4 M/mol; loglo (Ws-')(2) = 17.18, EA(2) = 218.7 kJ/mol; loglo(k&-') (3) = 14.99, and EA(3) = 209.8 W/mol. Reaction (1) contributes 17.2 %, reaction (2) contributes 18.9 % and reaction (3) contributes 64.0 YO to the total mass loss. A correlation coefficient of 0.99856 and a mean of residues of 0.03802 were calculated for the best fit. To improve the formal reaction model, a mechanism comprising three parallel n* order reactions was fitted to the experimental TGA curves from all heating rates. The best fit was achieved for the following kinetic parameters: logl0&Js-') (1) = 17.52, EA (1) = 213.7 kJ/mol, n(1) = 4.61 ; loglo &Js-') (2) = 14.90, EA(2) = 208.3 kJ/mol, n(2) = 0.75 ; and loglo(Ws-') (3) = 15.29, EA (3) = 169.0 kJ/mol, n(3) = 3.28. Reaction (1) contributes 46.9 %, reaction (2) contributes 50.5 % and reaction (3) contributes 2.5 % to the total mass loss. A correlation coefficient of 0.99938 and a mean of residues of 0.02448 were calculated. The best fits for the formal reaction models with 1 st order reactions and n* order reactions, respectively are displayed in Figure 8.
P -
i0
temperature / K Fig. 8 TGA results for beech wood at different heating rates and best fit models for a three step mechanism with Is' order reactions ( ) and n' order reactions ( -).
---
1088
CONCLUSIONS
Thermogravimetric analysis can be performed with heating rates up to 108 Wmin without significant temperature gradients inside small samples. An STA apparatus enabling a temperature calibration under the exact experimental conditions will give reliable process temperatures. This allows kinetic modelling without consideration of the enthalpy balance. Formal kinetic models have to be valid for different heating rates which will occur in all reactors with a distributed temperature. To obtain reliable kinetic models, experiments covering the desired range of heating rates have to be performed. A simultaneous evaluation of all experiments applying multivariate regression results in the description of the pyrolysis behaviour in the analyzed range of heating rates by a formal reaction model with one set of kinetic parameters. Multivariate regression is an important tool for deriving an appropriate model representing the best compromise between accuracy and number of parameters for the desired engineering application. The results of our kinetic analyses demonstrate that formal kinetic models can be derived which describe mass loss during pyrolysis of biomass and its components quantitatively for heating rates varying over more than two decades. Nevertheless, care should be taken when transferring kinetic parameters from the literature. The exact chemical composition of the biomass samples is crucial to their pyrolysis behaviour. Additionally, mass transport phenomena influence the char forming reactions dramatically thus changing the observed macro kinetic parameters and the contribution of the individual reaction steps to the overall mass loss. To derive reliable reaction models for engineering applications, experiments with the desired technical particle sizes and bulk phase properties should be selected or mass transport models should be incorporated into the kinetic modelling. NOTATION
heat capacity, J/(kg K) correlation coefficient apparent activation energy, kJ/mol characteristic length of sample, m enthalpy of reaction, J/kg rate constant, l/s pre-exponential factor, 11s mean of residues reaction rate, kg/(m3s) ideal gas constant, J/(mol K)
time, s t T, To temperature, initial temperature, K Tx=h temperature at sample surface, K X spatial coordinate, m heating rate, Wmin P initial void fraction, m3.JmPbulkphase €0 heat conductivity, W/(K m) h density, kg/m3 P final density, kg/m3 PF, initial density, kg/m3 Po
REFERENCES 1. Antal M.J. (1982) Biomass pyrolysis: A review of the literature. Part 1 carbohydrate pyrolysis. Advances in Solar Energy, 1,61- 111 2. Antal M.J. (1985) Biomass pyrolysis: A review of the literature. Part 2 lignocellulosepyrolysis. Advances in Solar Energy, 2, 175-255 3 . Hajaligol M., Howard J.B., Longwell J.P. and Peters W.A. (1982) Ind. Eng. Chem. Process Des. Dev., 21(3), 457-465
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4. Antal M.J. and Vhrhegyi G. (1995)Cellulose pyrolysis kinetics: The current state of knowledge. Ind. Eng. Chem. Res., 34,703-17. 5 . Vhrhegyi G.,Jakab E. and Antal M.J. (1994)Is the Broido-Shafizadeh model for cellulose pyrolysis true? Energy Fuels, 8, 1345-52. 6. Milosavljevic I. and Suuberg E.M. (1 995)Cellulose thermal decomposition kinetics: Global mass loss kinetics. Ind. Eng. Chem. Res., 34,1081-91. 7. Reynolds J.G. and Bumham A.K. (1 997)Pyrolysis decomposition kinetics of cellulose-based materials by constant heating rate micropyrolysis. Energv Fuels, 11,88-97. 8. Antal M.J. and Vhrhegyi G. (1997)Impact of systematic errors on the determination of cellulose pyrolysis kinetics. Energy Fuels, 11, 1309-10. 9. Di Blasi C.(1998)Comparison of semi-global mechanisms for primary pyrolysis of lignocellulosicfbels. J. Anal. Appl. Pyrolysis, 47,43-64. 10. Antal M.J., Vhrhegyi G. and Jakab E. (1998)Cellulose pyrolysis kinetics: Reviseted. Ind. Eng. Chem. Res., 37, 1267-75. 1 1. Conesa J.A., Marcilla A., Caballero J.A. and Font R. (2000)Comments on the validity and utility of the different methods for kinetic analysis of thermogravimetric data. Pyrolysis’2000, April 2-6,2000,Seville, Spain. 12. Kaisersberger E. and Opfermann J. (1991)Thermochim.Acta, 187, 151. 13. Mok W.S.-L. and Antal M.J. (1983)Effects ofpressure on biomass pyrolysis. 11.Heats of reaction of cellulose pyrolysis. Thermochim.Acta, 68, 165- 186 14. Curtis L.J. and Miller D.J. (1988)Transport model with radiative heat transfer for rapid cellulose pyrolysis. Ind. Eng. Chem. Res., 27, 1775-83. 15. Hoare I.C.and Hurst H.J. (1992)The evaluation of kinetic parameters fiom thermogravimetric curves. Thermochim.Acta, 203,127-35. 16. VOlker S.,Rieckmann Th.and Klose W. (2000)Thermokinetic investigation of cellulose pyrolysis - impact of final mass on kinetic results. Pyrolysis’2000, April 2-6,2000,Seville, Spain. 17. Chornet E. and Roy C. (1980)Compensation effect in the thermal decomposition of cellulosic materials. Thermochim.Acta, 35,389-93. 18. Grranli M., Antal M.J and Vhrhegyi G. (1 999)A round-robin study of cellulose pyrolysis kinetics by thermogravimetry.Ind. Eng. Chem. Res., 38,2238-2244 19. Bradbury A.G.W., Sakai Y.and Shafizadeh F. (1979)A kinetic model for pyrolysis of cellulose. J. Appl. Polym. Sci., 23,3271-80. 20. Lanzetta M., Di Blasi C.and Buonanno F. (1997)An experimental investigation of heat-transfer limitations in the flash pyrolysis o f cellulose. Ina! Eng. Chem. Res., 36,542-52. 21. VBrhegyi G.,Antal M.J., Jakab E. and SzaM P. (1997)Kinetic modeling of biomass pyrolysis. J. Anal. Appl. brolysis, 42,73-87.
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A Modeling Study On Cellulose Particle Pyrolysis Under Fluidized-Bed Conditions Chunjiang Yu', Wennan Zhangl and Kefa Cen2 I Department of Applied Science, Mid Sweden University, 871 88 Harnosand, Sweden 2 Institute for Thermal Power Engineering, Zhejiang University, 3I002 7 Hangzhou P. R. China
ABSTRACT To study biomass pyrolysis processes, a numerical model is developed in this work focusing on the cellulose constituent. A single cellulose particle is supposed to experience pyrolysis at a fairly high heating rate in the case of typical fluidized bed conditions. The model involves the pyrolytic lunetic scheme and the detailed heat and mass transfers caused by radiation, conduction, diffusion and convection with respect to the solid, liquid and gas products from pyrolysis. From the solution of the model, the central role of the reaction heat versus a slow heat supply through the particle is identified. The dynamic characteristic of the pyrolysis and the evolutions of the internal pressure and the flow of the products are analyzed quantitatively. Computation of the model shows the importance of particle permeability, whereas the liquid phase of active cellulose and inter-particle secondary reaction of volatile play a negligible role during the pyrolysis under a typical fluidized bed conditions both for small and large particles. Various predictions by the model provide a good understanding of the complex process of biomass pyrolysis, which can be useful for applications of thermochemical conversions of biomass. INTRODUCTION Fluidized bed application is fairly popular in biomass thermochemical conversions. The one of reason is due to the high efficient heat and mass transfers between particles and gases in the reactor. The initial and essential step of biomass thennochemical conversions such as combustion and gasification in a fluidized bed is pyrolysis. Directly experimental study on the pyrolytic process is limited. Numerical simulation can be an effective way for understanding of the first step of biomass combustion and gasification in applications. From literature, a great amount of work on biomass pyrolysis, including modeling of pyrolysis, has been found over the last two decades. However, the most of modeling works deals with low heat flux conditions or an extremely high heating rate (flash pyrolysis), or focus on the prediction of product distribution [ 1-31. The comprehensive 1091
pyrolytic process inside a biomass particle with median intensive heating under fluidized bed conditions remains to be studied in detail for a better understanding of the process and optimal design of related industrial units. The aim of the present work is to develop a mathematic model describing the behavior of a single cellulose particle when heated up under a fluidized bed condition. This may give an insight into the temperature profile of the particle in connection to the reaction heat, the characters of product gas and liquid flows, the secondary reaction of primary volatile and other important phenomena of biomass pyrolysis.
THE MATHEMATIC MODEL The model is one-dimensional over a spherical, dry cellulose particle. The particle is supposed to expose to an inert, high-temperature environment of a fluidized-bed subject to an intense heating immediately. A typical constant overall heat transfer coefficient is adopted.
Celloluse-b K1
Primary K4 .-b Volatile
x Char + y Gas + ( 1-x-y ) Sec. Tar
celluloseh Active
Char + (1-z) Gas
Fig. 1 Modified Broido-Shafizadeh Scheme [4]. Table I Kinetic data of the modified B-S scheme. Reaction Ki 1 2
3 4
Ai [s”]
Ei [kj/mole]
2.8e19 242.8 3.17el4 198.0 1.26e18 238 1.32e10 150.7 3.16e9 147 4.28e6 107.5 n=0.35 i=0.35 ;j=0.3
AHi[kj/Kg]
References
40 418
Bradbury, [4] Bradbury, [4] Varhegyi et. al. [5] Bradbury, [4] Varhegyi et. al. [5] Liden’s [6] Colomba Di Blasi [7] Estimated
418 -42
The kinetic scheme that has been chosen is basically the Broido-Shafizadeh [4] as shown in Fig.1. The most widely accepted Broido-Shafizadeh scheme provides a more reasonable description of the involved kinetics than other schemes. However, in the present model, some modifications are made. First, some more reliable kinetic data obtained by Varhegyi et. al. [5] is employed as seen in Table 1. Secondly, the intermediate product “active cellulose” is included, according to recent experimental study of 0. Boutin et. a1 [8]. The experimental study not only proved the existence of such “active cellulose”, but also showed the short-life intermediate product would experience a phase change from solid to liquid when temperature rises up to about 740K. A fast heating may raise the temperature of active cellulose over 740K before it consumed, which leads a
1092
liquid fluid present in the pyrolytic process. Third, the secondary reaction includes char and secondary tar as additional products to constitute a complete reaction. The physical aspects involved are heat transfer and mass transfer. The external heat and mass transfers are free of resistance under the environment of fluidized beds. Only the internal heat and mass transfers are treated here. To develop the mathematic model, the following assumptions and statements are made: 1.
2. 3. 4. 5. 6.
7. 8.
9.
There is a local thermal equilibrium among the solid phase, liquid phase and gaseous phase. Negligible enthalpy flux due to fluid species diffusion. Negligible lunetic and potential energy, negligible body forces. No particle shrinkage and surface regression. The inert gas and gaseous products resulted fiom reaction behave ldce ideal gas, with a constant mean molecular weight. The secondary reaction quenches as soon as volatile products leave the solid surface. The initial porous structure of cellulose is filled with inert gas before pyrolysis. Physical properties (porosity, permeability, thermal conductivity, specific heat and fluid viscosity ) depend on temperature and conversion extent. The secondary reaction is homogeneous gas reactions.
Based on the above assumptions, The mass conservation through a differential control volume of a layer of spherical cellulose particle are thus formulated for cellulose (l), char (2), active cellulose (3), volatile (4), secondary tar ( 5 ) , gas (6) and inert gas (7), respectively.
The variables in 0 mean they are based on the whole control volume, in <>f mean they are based on porous volume, in Oamean they. are based on volume occupied by the active cellulose. The whole control volume is occupied by solid phase (cellulose and char), liquid phase (active cellulose) and gas phase (porous volume). Considering the flow of gas and liquid through the porous cellulose, Darcy's law (Eqs
1093
8 and 9) and ideal gas law (Eq. 10) are introduced to govern the superficial velocity of the fluid with respect to the pressure and the particle permeability: = ---(P) Ka a
(24,)
Pa
%=: - Kf a (P)
(9)
(Ur)
where K, and Kf are intrinsic permeabilities of porous matrix for the liquid phase and the gas phase, respectively. The energy conservation is written as Eq. 11.
(1 1)
where
describes the instantaneous energy absorption associated with chemical reactions. %tr is effective thermal conductivity of porous material, which takes into account thermal conductivity of the solid matrix and gas in pores and radiation in the pore system, as shown in Eq. 12: K, = E x K+7;1~K,+(l-7;1)~K, ~ + - x& ~ X d p l-& The following equations also hold true:
1094
x4xT’
(12)
();:
6 = 1,2,3,4)
Ki = Aiexp --
The following statements are applied to the boundary conditions and initial conditions:
0
The fluidized-bed is supposed to have atmospheric pressure and constant temperature of T, = 800°C. The cellulose particle submerged in the fluidized bed is heated with an overall heat transfer coefficient: HOut=250 Wm-2K.There is no diffusion and convection at the center of the particle. The concentrations of all released species are equal to 0 outside the particle. The initial temperature of the particle is 300K. The initial concentration of cellulose is 450 kgm-3. The appearance density of inert gas in pores is 0.76 Kgm-3. The solution of the 10 partial differential equations allows to determine temperature, composition and velocity as a function of both time and space. The system of equations is solved by the method of finite difference.
COMPUTATION RESULTS AND DISCUSSION TEMPERATURE INSIDE THE PARTICLE
The spatial temperature profiles versus the time are shown in Fig. 2 for a 2 mm particle and in Fig. 3 for a 10 mm particle. In the figures, the particle is divided equally into 10 layers along the radius. At the first stage of temperature build-up (in 1 second in the case of 2 mm particle), the particle temperatures at different positions from the center to the surface rise in the shape of normal exponential temperature-rising profile similar to the field of a pure heat conduction object. 1,100 1,000
900
6Z aoo 700
2L CI
600 500 400 300
0
1
2
3
4
5
time ( S )
Fig. 2 Particle temperature vs. time (2mm particle). 1095
6
The heat from the environment is transferred into the particle layer by layer and become the sensible heat of the cellulose substance. This situation remains until the outside layer of particle reaches an enough high temperature of about 660K,at which, the chemical reactions start to come into effect. Due to the reaction heat needed for pyrolysis, the temperature increase is suppressed significantly, and is turned to the chemical reaction stage of rather flat temperature profiles. The temperature is held between 650 K to 750 K by endothermic pyrolytic reactions, which is much lower than the environmental temperature of 1073K. The different stages of temperature build-up can clearly be seen to be corresponding to different zones in Fig. 4 where the temperature is plotted against the space variable. Zone A, drawn with the dashed lines, corresponds to the first stage of a pure heating of an object. Zone B, drawn with the solid lines, represents the endothermic chemical reaction stage. Zone C represents the fast temperature increasing stage (see Fig. 2) after the pyrolytic reaction is complete at the time of 5.8s. In this zone, only a char particle of developed porous structure is left. The temperature is uniform over the entire particle space and increases rapidly up to the environmental temperature due to a good access to the heat radiation and cease of endothermic reactions. It should be emphasized that the endothermic reaction of decomposition plays the central role in determine the particle temperature profile. When pyrolysis begins to proceed intensively (after 1.6s in Fig 4), the reactions take place in out layer, absorb almost all the heat transferred from outer layers and cut down the energy flux to internal layers. This is the reason why the layers situated inside the reacting layers have an uniform and steady temperature variation over the time (see Fig. 2). Consequently, make the pyrolysis process can only proceed layer by layer. Figure 3 shows the temperature profile of 10 mm size particle, similar to that of the 2 mm size particle. The shapes of these curves are very similar to those of small particle. However, the variation of temperature increasing rate for every layer is much more distinct than the case of 2 mm particle. This is because the high char yield of the larger particle reduces the radiative fraction of thermal conductivity, and increases the resistance of heat transfer.
1096
I100 I
I
1000 900 n
2
800
$ 700 &
+
600
500 _.___._.___ - ..-..---
400
300
.. , . . . .. .
1
A
_ _ _ _ ...
2
.
.
~ ~ . I - - - . ~-~~ - ~ I- - - .I
3
4
5
-.-. ~ . ~ .
I
0.4s
I
I
1
I
6
7
8
9
10
position alone the radius Fig. 4 Particle temperature vs. position (2-
particle)
PRESSURE INSIDE THE PARTICLE
Pressure is an important parameter of biomass particle pyrolysis, which governs the product gas flow through the particle, and consequently influences the mass transfer and heat transfer inside the particle. The results calculated by the model are shown in Figs 5 and 6 for particles of 2 mm and 10 mm in diameter. In Fig. 5, the numbers of ‘1’ to ‘10’ mark each layer from particle center to surface. Three peaks appear in the pressure profile over time, marked with ‘A’, ‘B’ and ‘C’, The peak ‘A’ appears before chemical reactions, and is caused by expansion of inert gas in the pores when the temperature rises. The particle internal pressure rises fast to the highest value over the entire pyrolysis process due to high heating rate and comparably compact porous structure in the very beginning. The peak ‘C’ is caused by biomass devolatilization that proceeds layer by layer into the particle center and produces a large quantity of gaseous products. The small peak ‘B’ for the inside Iayers is induced from rapid devolatilization of the surface layer starting at about 0.8s. After peak ‘C’, the out layers become more porous char structure whose permeability is several order of magnitudes larger than raw cellulose, and gaseous products can flow out of particle with little resistance. Thus, the local pressure level off down to ambient pressure. The general trend in the pressure profile for the small particle remains the same for the large particle of 10 mm as shown in Fig. 6. But some differences between them can be seen clearly. The peak ‘B’ disappears, the in-particle pressure has reached a much higher level and the pressure level difference between the peaks ‘A’ and ‘C’ is enlarged dramatically. These can be attributed to the low permeability of both reacting layers and char layers, leading to higher resistance inside the large particle against gas flow through the particle. Unlike coal, the internal pressure of a biomass pyrolytic particle under fluidized bed condition is very low due to the high permeability of cellulose and char structure.
1097
Even for the large particle of lOmm, the maximum internal relative pressure inside the particle is only a little more than 70Pa. Under such a low pressure, the physically cracking process caused by internal pressure can not be the case for biomass particles as it is for coal particles.
25 20 h
k
'-3 15
v
v)
!i
.E 2
10
Y
(d
5
0
0
1
2
3
4
5
time(S) Fig. 5 Internal relative pressure vs. time (2mm particle) 80
-k
70 60
v
8 &
50
v) m
40
0
>
-
'2 30
d"
20 10 0
6
pressure gradient through the particle and is governed by many factors such as temperature, reaction rate, gas properties and porous structure of the particles. 0.12 h
0.1
3,2 0.08
.co.
% 0.06 (cl
0
x
-
50 0.04 9 0.02
0
2
4
6
8
10
position along radius
Fig. 7 Gas velocity vs. position (2mm particle) The gas phase velocity distribution over the particle space is shown in Fig. 7. This figure consists of a serial of velocity distributions along the radius at different moments. Those moments are chosen for the reason that the gas velocities of local layers reach their maximum value at the moments. It should be mentioned that those moments just match the time when the active cellulose decomposition rates for corresponding layers reach the maximum values. It suggests that the gaseous product from active cellulose primary decomposition contributes the main part of the gas phase flow. The maximum gas velocity at the moment of 5.787s is higher than others. This can be explained by particle size effect. A small particle has a shorter reaction period and a high reaction rate, which leads to a higher gas velocity. LIQUID ACTIVE CELLULOSE
Concerning active cellulose, the most attention is given to its phase state. As recently proposed by Lede [9][3], the temperature of this intermediate substance is a key parameter. Active cellulose is believed to undergo a phase change at temperature around 740K. In slow pyrolysis as used for thermogravimetric analysis, active cellulose tends to be consumed in solid phase before it reaches the phase-change temperature, while in flash pyrolysis, visible liquid intermediate can be found at the high temperature frontier of reacting zone. The flow of liquid active cellulose m a y have a significant influence on pyrolysis if it exists to a certain extent. The case of middle intensive pyrolysis is identified as shown in Figs 8 and 9. The existing zone of liquid active cellulose, in which the temperature is above 740K,can be found at the right-down comer of Fig. 8 for the small particle of 2 mm. The liquid active cellulose only appear in the internal layers of a limited volume at the very end of reaction period with quite low densities. The amount of liquid active cellulose is likely negligible. It is estimated that the mass of the active cellulose that flows out of the particle is less than 0.005% of origial cellulose mass. For the large cellulose particle of 10 111111, no liquid active cellulose at all can be 1099
found from the model calculation as shown in Fig. 9. Therefore, it can be concluded that the effect of active cellulose is negligible for the pyrolysis under middle intensive heating condition. For the case of flash pyrolysis, intensive heating supplies endothermic reactions with sufficent heat in time and raises the temperature of the solid active cellulose product above 740K immediately so that a large quantity of liquid active cellulose may emerge from the particle surface. 400
;;;350 <
5
300
v
.$’ 250 8 ‘O
200
8 2 150 0
0
.-*i? 100 0
50 0
0
1
2
3
4
5
6
time (S)
Fig. 8 Active cellulose density vs. time (2mm particle)
350 iT 300 E
2.-b 250 5 2 3- 150 a 200 d
8 .-p 100
‘ +a
50
0 0
40 50 60 Time@) Fig.9 Active cellulose density vs. time ( l o r n particle) 10
20
30
1100
70
CELLULOSE DE VOLATILIZATION Only cellulose is the feedstock of pyrolysis process treated in the model. Cellulose mass consumption represents the degree of the pyrolytic conversion. Figures 10 and 11 show the mass losing curves as pyrolysis proceeds for 2 mm and 10 mm particles respectively. As can be seen that cellulose density decreases all the way down to zero. The density in the outside surface layer decreases much faster than the center. It takes about 5s for the 2 mm cellulose particle to be devolatilized completely, and 60s for the 1Omm particle, It also shown that devolatilization of cellulose particles proceeds layer by layer, which is more obvious for the outside layers and for the large particle.
0
1
2
3
4
5
time (S) Fig. 10 Cellulose density vs. time (2 mm particle)
Tine(s)
Fig. 11 Cellulose density vs. time (10 mm particle)
1101
6
PRODUCT CHAR Char is the only solid product of cellulose pyrolysis and constitutes the reactionfinished layer behind the reacting frontier. Figs 12 and 13 show the char density variation in solid curves and the density change rate in dashed curves (differential curves) as a function of time, i.e. char accumulationand production rate in each particle layer. The numbers of ‘1’ to ‘10’ again represent the particle layer fiom the center to the surface. Production of char can be divided into two stages. The first one is characterized by the peaks marked by numbers on the dashed curves, which correspond to the existing periods of active cellulose for every layer. Therefore, these peaks of char production are the results of primary decomposition.The second stage takes place after the end of pyrolysis reaction, which is marked by overlapping peak ‘S’for every layer. This is because the remaining volatile in porous char structure stays still, and undergoes the secondary reaction at a high temperature. Before this stage, the secondary reaction is less important. The two cases of different particle sue show similarity of general variation trend and also some differences. The char yield for the large particle is higher due to a low heating rate during the pyrolysis process. The rate of increase in char density of the outer layers during the late stage of pyrolysis accelerates, which is associated to the secondary reaction of the primary volatiles.
2
6
0
1
3
2
4
time (S) Pig. I2 Char density vs. time (2mm Particle)
1102
5
6
25 20 n
m
E
15 2.
.r. v)
d
4 I0 3
8
5 0 0
10
20
30
40
50
60
70
Time (S) Fig.13 Char density vs. time (lOmm particle)
PRODUCT VOLATILES The volatile concentration is determined by the rate of volatile generation, the gaseous phase flow and cracking of the secondary reaction. The spatial distribution of volatile concentration over a 2 mm particle is shown in Fig. 14. The curves labeled with ‘1’ to ‘10’ correspond to the curves from 3.192s to 5.797s in Zone B of Fig. 4. It can be seen from Figs 4 and 14 that the volatile concentration in the reacting outer layers increases rapidly in the beginning of pyrolysis (at the moment of 0.6s). This trend holds until the surface layer is consumed at the moment of 3.192s (curve ‘1’). Afterwards, the improved permeability of outer layer enhances the outward product flow and suppresses the increasing of volatile concentration. As the pyrolysis frontier steps into the internal small layers, the quantity of produced volatile decreases. Consequently, the volatile concentration in outer layers begins to decrease from curve ‘ 5 ’ . After curve ‘10’ (moment 5.797s), the primary reaction stops and the outward flow becomes negligible. At this stage, the volatile inside particle char matrix is almost stagnant with an uniform distribution of the concentration. Figure 15 shows the spatial distribution of volatile flux in a 2 mm particle. All the solid curves for time 3.92s to time 5.797s give a flat profile on the right hand side. For the moment specified by each curve, the primary decomposition has finished locally in the outer layers. This is also the case for a larger particle of 10 mm diameter. Thus, it is suggested that the secondary reaction plays a negligible role during this stage. In other word, volatile interior-particle crack during cellulose pyrolysis is not important under the fluidized-bed condition. From the model prediction, 0.088% of volatile undergoes the secondary crack inside a 2 mm particle, and about 2% for a 10 mm particle.
1103
1.2 n
Q%m
l
0 0.8
.3
0
8
.0
0.6
c . +-I
f
0.4
0.2 0
1
4
3
2
5
7
6
9
8
10
positions along radius Fig. 14 Volatile concentration vs. position (2mm particle)
4.5
#J
4
d
3
n
3.5
2 G
0
CI .CI
4 c
>
2.5
2 1.5
1 0.5 5 78s 5.795
0 0
1
2
3
4
5
6
7
8
9
10
positions along radius
Fig. 25 Volatile flux vs. position (2 mm particle) CONCLUSION A comprehensive mathematic model describing the pyrolysis process of single cellulose particle under a fluidized bed condition is built up by formulating the
1104
transport phenomena and chemical reactions involved. The model addresses the most important parameters such as temperature, pressure, velocity and concentration over time and space, and gives reasonable predictions. The computation results show that the pyrolytic reactions take place in the form of layer by layer inside the particle. The reaction temperature is much lower than the environmental temperature. The endothermic primary decomposition reaction against a slow heat transfer through the particle plays the most important role in governing the biomass pyrolysis process. The particle permeability is the key parameter for the in-particle pressure and the heat and mass transfer through the particle. The appearance of liquid active cellulose is nearly impossible under the simulated pyrolytic condition even for small particles (2 mm). An extremely small part of volatile can be consumed by internal-particle secondary cracking. ACKNOWLEDGEMENTS
The authors would like to acknowledge the financial support of EU under Project Contract: JOR3CT980306, for the realization of this work. The co-operation of the partners involved in the EU-project is also greatly appreciated. Thanks should also go to Swedish National Energy Administration for a part of the financial support
NOMENCLATURE
dynamic viscosity density P Stefan-Bolzman constant 0 Ki (i=1,2,3,4) Arrhenius rate Ei(i= 1,2,3,4) activation energy Ai(i=l ,2,3,4) pre-exponential factor velocity U pressure P universal gas constant R molecular weight M heat capacity C temperature T AHi(i=l ,2,3,4) heat of reaction effective thermal conductivity Icff porosity E pore diameter d* time 7 level of conversion (<pa>+)/' rl
P
Subscripts W
C
a V
t2 g
cellulose char active cellulose volatile secondary tar permanent gas 1105
inert gas total gas phase (v+t2+g+i)
i f
Superscript 0 Oa
Of
ow
oc
initial state variable based on volume occupied by active cellulose variable based on volume occupied by total gas phase variable based on volume occupied by cellulose variable based on volume occupied by char
REFERENCES 1.
2. 3. 4. 5.
6. 7 8.
9.
Jalan, R.K. and V.K. Srivastava, (1999) Studies on pyrolysis of a single biomass cylindricalpellet - kinetic and heat transfer effects. Energy Conversion & Management, Vo1.40,467-494 Di Blasi, C., (1996) Heat, momentum and mass transport through a shrinking biomass particle exposed to thermal radiation, Chemical Engineering Science, VO1.5, No.7, 1121-1132 Lede, J., (1994) Reaction temperature of solid particles undergoing an endothermal volatilization. Application to the fast pyrolysis of biomass. Biomass and Bioenergy, Vo1.7, No. 1,49-60 Bradbury, A. G. W. et. al. (1979) A kinetic model for pyrolysis of cellulose. Journal of appliedpolymer science, Vo1.23: 3271-3280 Varhegyi G.P. et. al. (1994) Kinetics of the thermal decomposition of cellulose under the experimental conditions of thermal analysis. Theoretical extrapolationsto higher heating rates. Biomass and Bioenergy, Vo1.7 No. 1,69-74 Liden, A.G. et. al. (1988) A kinetic model for the production of liquid from the flash pyrolysis of biomass, Chem. Eng. Comm. Vo1.65,207-221 S i Blasi, C.m (1996) Kinetic and heat transfer control in the slow and flash pyrolysis of solid. Ind. Eng. Chem. Res, Vo1.35,37-46 Boutin, 0. et. al. (1998) Radiant flash pyrolysis of cellulose -evidence for the formation of short life time intermediateliquid species. Journal of Analytical and applied pyrolysis, Vo1.47, 13-31 Lede, J. et al., (1996) The nature and properties of intermediate and unvaporized biomass pyrolysis martierial, Development in thennochemical biomass conversion, IEA Bioenrgy, Editor: A.V. Bridgwater, D.G.B. Boocock, p. 27-4 1
1106
Modeling Potassium Release In Biomass Pyrolysis Chunjiang Yu and Wennan Zhang a ' Department of Applied Science, Mid Sweden UniversiQ, 871 88 Harnosand, Sweden Institute for Thermal Power Engineering, Zhejiang University, 31 0027 Hangzhou, I?R.China
ABSTRACT: Alkali compound emission during biomass thermochemical conversion gives rise to a number of problems such as agglomeration, slag, fouling and metal corrosion in the conversion process system. Understanding the behaviour of alkali emission from biomass fuels is important to solve these problems. In this work, a study focused on the potassium compound release in biomass pyrolysis is carried out. The transformation of the element, K, is interpreted based on literature and the present study. A mathematic model of the potassium compounds emitting from biomass during pyrolysis is proposed. The different existing forms of potassium in biomass, the chemical equilibrium of the compounds at the pyrolytic environment and the potassium compound release due to vaporization at h g h temperature are taken into account in the model. The result of the mathematic model provides an understanding of the alkali compound release in biomass pyrolysis and the effects of the pyrolytic temperature, the fuel composition and residence time on the release process.
INTRODUCTION Biomass particles undergo a rapid heating once fed into a reactor of thermal conversions. The thermal degradation occurs in the first stage, which destroys the matrix structure of the main constituents in biomass: cellulose, hemicellulose and lignin. In the same time, the trace inorganic elements that are bonded in biomass fuel in different forms are released and transformed into various compounds. Some of them may form volatile compounds and are vaporized easily during the process, while others reside in char matrix as a part of ash. The alkali compounds released give rise to a number of problems in thermochemical conversion systems, such as bed material agglomeration and sintering of fluidized beds, slagging on refractory-lined walls in the h a c e , fouling on heat transfer tubes, deposit-related corrosion of superheaters, filters, gas turbines and engines [ 1][2]. Alkalis enriched micron-size char particles in pyrolytic bio-oil present a potential problem to the combustion in engines. Pyrolysis is the first and essential step in biomass thermochemical conversions, which is supposed to determine, basically, the quantity and forms in which alkalis are distributed into gas, liquid and solid phases. In the applications, for instance, the fuel
1107
gas from pyrolysis or gasification can be a good quality fuel with less alkali for reburning, while the remaining char and ash are alkalis-enriched and can be considered for fertilizer. In order to find solutions to the alkalis-related problems, the pyrolysis of biomass particle with respect to the alkali compound release needs to be studied in detail. Potassium in biomass exists as dissolved salts in the moisture or as a cation attached to carboxylic and other functional groups, which is highly mobile. It is most possibly bonded to chlorine as KCI if chlorine is available in biomass fuels [3]. The most possible potassium compounds released from biofuels are KCI, K2SO4and KOH [l]. A measurement by Surface Ionization Mass Spectrometry technique [4] showed that potassium release to the vapour phase in pyrolysis takes place in two stages. The first stage is in the temperature range of 200 C" to 500 C". The potassium release can be well coupled to the biomass devolatilization, and insensitive to the chlorine content in the fuel. In the second stage of temperature above 500 C", the potassium release is at its full speed, depending on the chlorine content in the fuel. Based on reported experimental studies, it can reasonably be assumed that the potassium release completes in two stages of low and high temperatures. In the first stage of low temperature, the organic bonded potassium releases following devolatilization of biomass constituents, cellulose, hemicellulose and lignin. In the second stage of high temperature, the inorganic compounds of potassium such as KCI and KOH are vaporized into the vapour phase. The present study will give a quantitative description of potassium release in the two stages, respectively.
EXPERIMENTAL Alkalis contained in biofuels vary in a wide range depending on biomass sources. Agricultural residues such as straw are on the top in the alkalis content, followed by grass, short rotation coppices and bark of trees. Woody biomass has the least content of alkalis. In the present work, a straw is used as a typical fuel for potassium release study. The elementary analysis of the straw is given in Table 1.
Table I Composition of the straw used in the experiment (w% on dry basis). C
H
0
N
S
c1
Si
Ca
K
others
43.2
4.96
37.7
1.22
0.05
0.56
7.61
0.29
2.22
2.19
In the experiment, the straw was heated up in a lab-scale pyrolyzer with external heating element. The temperature in the pyrolyzer was increased gradually and kept at about 800 C" for a period of time. The temperature of the fuel was registered as a function of time. The temperature history of the straw pyrolysis is shown in Fig. 1. The residual char was weighted and analyzed for determination of inorganic content. 0.39 kg char per 1 kg straw was obtained at the end of the test. The inorganic content in the residual char is given in Table 2.
1108
t
1,000
900
p
800
?
700
5 i
600
500
!
400
!
300
1 Temp. of straw (experiment)
= --_ 0
-'d " " ~ ~ " 0
20
40
60
"
~
80
Temp. of furnace Regression data ~
~
'
100
~
120
~
~
140
'
~
"
160
~
180
'
~
~
200
~
"
~
220
Time(min)
Fig. 1 Temperature history during the experimental test of biomass pyrolysis.
Table 2 Inorganic content in the residual char after pyrolysis (%w) L O
Me0
CaO
P7Oc
4.44
2.98
2.5
0.5
Thus, the potassium release to the vapor phase in the test can be calculated from the data of Table 2 as expressed below: Potassium release = (potassium in fuel - potassium in residual char)/potassium in fuel = 35.3 %
MATHEMATIC MODEL
The potassium release is divided into two parts, the organic bonded potassium release and the inorganic potassium compound vaporization. The amount of the organic potassium in biomass can be determined experimentally with the chemical fractionation method. The chemical fractionation techmque distinguishes different types of inorganic material according to their solubility in a series of increasingly aggressive solvents. Those materials soluble in water are soluble salts and loosely bonded material. Those soluble in ammonium acetate solvent are the organic bonded materials (i.e. ion exchangeable by ammonium acetate). Those soluble in hydrochloric acid are typically carbonates or sulfates. Those not soluble in any of these solvents are commonly in the form of oxides, silicates or sulfides. The literature data for straw are shown in Table 3, which gives the organic bonded potassium of 13%.
1109
~
~
~
'
~
Table 3 Chemical fractionation results of straw (from [ 5 ] )
Element
Water Soluble
Ion Exchangeable
Acid Soluble
%
%
%
%
0 93 87 16 100
0 7 13 0 0
0 0 1 0 0
100 0 0 84 0
Si Na K S
c1
Residual
ORGANIC BONDED POTASSIUM RELEASE The modeling of the organic potassium release is based on the following assumptions: There is no big difference in distribution of organic potassium between three main biomass components of cellulose, hemicellulose and lignin. , The mfluence of other inorganic materials and moisture content in biomass on organic bonded potassium release is negligible due to slow reactions at low temperature and short devolatilization time of biomass. . 3) The organic potassium release follows devolatilization of cellulose, hemicellulose and lignin, respectively. 4) All organic potassium is released to the vapor phase through cany-over of sub-micro particles with gas flow. The kinetic model of biomass devolatilization is employed to describe the potassium release process as given in Eq. 1. I da. = Ai ex(
dt
- A)fi- a i )
6 = 1,2,3)
RT
where i=l, 2, 3 denote cellulose, hemicellulose and lignin. a i is the fraction of potassium released, Ai and Ei are the pre-exponential factor and the activation energy. The data are introduced from literature as given in Table 4.
Table 4 Parameters of potassium release model Parameters Activation energy ElL6' Activation energy EJ6] Activation energy E3r61 Preexponential factor log ~ ~ [ Preexponential factor log A J ~ ] Preexponential factor log A ~ [ ~ I Content of hernicell~lose[~~ Content of cellulose L71 Content of lignin L71 Constant in Antoine Eq. for KClL8] Constant in Antoine Eq. for KOH[']
1110
Value, unit 154000 kjlmol 199000 kj/mol 34000 kj/mol 13.9 min-' ~ l 16.6 min-' 1.7 min-' 18% 36% 16% A4.782 B=7440.691 C=-122.709 A4.028 B=5854.906 C=-144.113
Incorporating the relationship of temperature verse time used in the experimental test, the organic potassium release is calculated as shown in Fig. 2. It can been seen from the figure that the organic potassium release starts at 180 C" as lignin devolatilizes first. The rapid release appears at temperature about 300 C" up to 400 C" due to fast devolatilization of cellulose and hemicellulose. Afterwards, the potassium release is slowed down by slow devolatilization of lignin. Up to the temperature of 500 C", the organic bonded potassium is released completely. 1
0.9 0.8
5.-
5
P
0.7 0.6
$ 0.5 2
r.
0.4
'E 0.3 w
0.2 0.1 0
0
500 1,000 1,500 2,000 2,500 3,000 3,500 4,000 4,500 Time@)
Fig. 2 Organic potassium release during biomass pyrolysis. INORGANIC POTASSIUM COMPOUND VAPORIZATION The modeling of inorganic potassium release is based on the following assumptions:
1) 30% of potassium bonded in KC1 reacts with silica oxide to form the eutectic compund into solid phase at a low temperature [3]. 2) In addition to KCl and K2S02,KOH is regarded as the only potassium compound produced from the reactions of potassium salts and moisture. The reactions are complete before vaporization. 3) KC1 and KOH become condensed small particles in biomass after drymg, and vaporize into the gas phase by molecule dispersion. 4) The vaporization takes place under a stagnant environment, on resistance-free particle surface and at zero partial pressure of the potassium compounds far from the particles. 5) K2S04 is supposed to remain in the solid phase in comparison with the volatilities of KCI and KOH. The above assumptions define the potassium release process as a mass transfer from the particle surface into the vapor phase. The vapor partial pressure of KC1 and KOH vary with temperature, which can be calculated by the well-known Antoine equation:
1111
log,, (P)= A - (B/(T + c)) in which, P is the saturation vapor pressure (bar), T the compound temperature and A, B, C are constants. The vapor dispersion process can be described with Shewood numbel;
i.e.
Sh = k,D I rA
(3)
where kA is the mass transfer coefficient of A in a medium ( d s ) , D the particle diameter and rAthe mass diffisivity of A in a medium (m2/s). Assumption 4. leads to is Sh = 2. Thus, from Eq. 3, the release rate of A on the particle surface, (DmolOA(mol/s), expressed as the following equation, ( D m o l ~ ~kA = A c,4=2 r A / D
lTD2 P/RT
(4)
where A is the particle surface area (m’), CA mole concentration of A on the particle surface (moVm3) and R the ideal gas constant. The KCl and KOH diffisivities can be estimated with Hirschfelder’s equation [9]:
The above equation expresses the diffusion of vapor A through B medium, in which, MA and MB are the molecular weight (kdkmol), Po the absolute pressure, csAB and RD are the molecule collision diameter and the molecule collision integral which are functions of substance properties and can be determined by empirical equations [lo]. Assume QA (kg) to be the total A vapor mass released, the following relation between QA and D holds true:
where Do is the initial diameter of the particles and pthe density of A substance. Putting Eq. 6 into Eq. 4 leads to the expression of A substance release rate by the vaporization of the substance:
Thus, KCl and KOH releases can be calculated from Integration of Eq. 7 by incorporating the fuel temperature history measured in the experimental test. The results are shown in Fig. 3.
1112
0.35
2
0.3 Period of organic
$0.25 .-C .- 0.2
K emission
Period of inorganic K emission
5
v) u)
g0.15
$
0.1
.t:
60.05 0 0
2,000
4,000
6,000
8,000
10,000 12,000
Time(S)
Fig. 3 Potassium release during biomass pyrolysis. DISCUSSION The potassium releases from straw pyrolysis in different ways are shown in Fig. 3 for comparison. The sum of the organic potassium release and the inorganic potassium compound vaporization gives the overall potassium release. At low temperature of 200 to 500 C" indicated on the left of the dashed line, the inorganic compound vapor pressures are kept at a too low level to give any contribution to the total potassium release. In this stage of pyrolysis, the potassium release is entirely attributed to the organic bonded potassium release, following the biomass devolatilization process. At the temperature of 500 C", the all organic potassium in the fuel is released (13% of total K in the present work). After ths, the release process is switched to the inorganic potassium release indicated on the right of the dashed line. The KOH vaporization starts at a temperature of about 800 C", and hold a high vaporization rate as the temperature is kept about 900 C". The KCl vaporization, on the other hand, starts at a higher temperature of about 850 C" and hold a low vaporization rate under the same condition. From the model computation, 15.89% of total fuel potassium is released in the form of KOH, While 1.29% of total fuel potassium is released in the form of KCI. The total potassium release (plus the organic potassium release) is calculated from the model to be 30.18%. T h s value is fairly in agreement with the experimental measurement (35.3%) in the present work. After vaporization, KOH may be converted into KCl by reaction with HCl depending on the chlorine content in the fuel. Chlorine contained in the fuel is released already, to a great degree, as HC1 in the biomass devolatilization stage. In summary, a small amount of potassium is released in the biomass devolatilization stage due to the limited content of organic bonded potassium in the fuel. The main potassium release occurs at a high temperature by means of vaporization of potassium compounds. The pyrolytic temperature, fuel properties and the residence time have an significant influence on the potassium release to the gas phase. In the case of a lower temperature of 800 C" being maintained instead of 9OOC" in the present study, the inorganic potassium release will be reduced to 6.26% as calculated from the 1113
model. Thus, 80.7% of potassium can be kept in the solid residue of pyrolysis. Therefore, low temperature pyrolysis and gasification show a great advantage when a hgh alkali content biofiel is treated. CONCLUSION Alkalis-related problem in the thermochemical conversion of biofuels is one of the main limits to the increasingly utilization of biomass. A theoretical study focusing on the potassium release from biomass pyrolysis is carried out. The mathematic model built up in the study gives computation of potassium release in biomass pyrolysis in a good agreement with the experimental measurement. According to the model, the potassium release in the biomass pyrolysis can reasonably be explained as two parts, organic bonded potassium release and inorganic potassium compound vaporization. The latter part plays a more important role in determining the amount of alkalis released to the vapor phase. The alkalis release depends on fuel properties, pyrolytic residence time and the pyrolytic temperature. Keeping a lower temperature from 900C" to 800 C" will capture alkalis in the solid residue by more than 80% under the conditions of the present study. ACKNOWLEDGEMENTS The authors would like to acknowledge the financial support of EU under Project Contract: JOR3CT980306, for the realization of t h s work. The co-operation of the partners involved in the EU-project is also greatly appreciated. Thanks should also go to Swedish National Energy Administration for a part of the financial support. REFERENCES 1. Bryers, R. W. (1996) Fireside Slagging, Fouling, and High-temperature Corrosion of Heat-transfer Surface Due to Impurities in Steam-raising Fuels. Prog. Energy Combust. Sci. V01.22 22-120. 2. Valmari, T., Lind, T. M., Kauppinen, E.I., Sfiris, G, Nilsson, K. and Maenhaut, W. (1999) Field study on ash behavior during Circulating Fluidized-Bed Combustion of Biomass.2.Ash Deposition and Alkali Vapor Condensation. Energy & Fuels, 13, 390-395. 3. Bjorkman, E. and Stromberg, B. (1997) Release of Chlorine from Biomass at Pyrolysis and Gasification Conditions,Energy & Fuel, 11,1026-1032. 4. Olsson, J. G., Jaglid, U. and Pettersson, J. (1997) Alkali Metal Emission during Pyrolysis of Biomass. Energy & Fuel, 11,779-784. 5. (1996) Alkali Deposits Found in Biomass Boilers: The behavior of inorganic material in biomass-frred power boilers - field and laboratory experiences, NREL/TP-433-8 142, SAND96-8225 VII. 6. Hsisheng, T., and Wei, Y.C. (1998) Thermogravimetric Study on the Kinetics of Rice Hull Pyrolysis and Influence of Water Treatment. Ind. Eng. Chem. Res., 37,3806-3811 7. Orfao, J. J. M., F.J.A. Antunes and J.L. Figueiredo (1999) Pyrolysis kinetics of lignocellulosic material- three independent reactions model. Fuel, 78,349-358 1114
Stull, D.R. (1947) Vapor Pressure of Pure Substances Organic Compounds. Znd. Eng. Chem., 39, 517-540 9. Hirschfelder, J. O., R. B. Bird and E. L. Spotz (1949) Chem. Rev. 44. 205 8.
10. Brokaw, R.S. (1969) Znd. Eng. Chem. Proc.Des.Develop,8,240
1115
Comparative study on char properties and pyrolysis kinetics of different lignocellulosic wastes P.R. Bonelli, P.A. Della Rocca, G.E. Cerrella, A.L. Cukierman* Programa de Investigacion y Desarrollo de Fuentes Alternativas de Materias Primas y Energia (PINMATE) - Departamento de Industrias Facultad de Ciencias Exactas y Naturales. Universidad de Buenos Aires. Intendente Guiraldes 2620. Ciudad Universitaria - (1 428) Buenos Aires, Argentina.
ABSTRACT: Chemical composition, heating values and surface properties of the char samples generated fkom slow pyrolysis of different abundant lignocellulosic wastes under identical operating conditions were examined to assess comparatively their suitability for potential use as biofuel andor for further processing to produce activated carbons. Sawdust from two different wood species (Aspidosperma australe, Populus deltoide), peanut husks, and olive stones were selected. Proximate and ultimate analyses as well as physical adsorption measurements of nitrogen and carbon dioxide were performed. Heating values were estimated from a multivariate correlation in terms of samples' elemental composition. All the chars obtained are found potentially useful as relatively pollution-free solid biofuels, the char arising from sawdust of Aspidosperma wood having the greatest potential. Comparison of surface properties with those determined for the olive stones-char, taken as a reference, suggests that the chars from peanut husks and from sawdust of Populus wood could be potential adequate candidates for further conversion into activated carbons. Kinetic measurements of the wastes pyrolysis were additionally conducted by non-isothermal thermogravimetry over the temperature range 25"C-9OO0C. Kinetic parameters were estimated by applying a published model, which assumes a steadily increasing variation in the activation energy with the course of pyrolysis and allows effective representation of experimental data for all the selected wastes over the whole range of temperatures. Appreciable differences in the estimated kinetic parameters are found.
INTRODUCTION
Large amounts of lignocellulosic wastes generated in processing of wood and agricultural products constitute attractive potential raw materials for production of charcoal and activated carbons, especially in those countries where economies are strongly based on agriculture and forestry. Conversion may additionally contribute to an overall waste reduction strategy in order to diminish environmental pollution arising from accumulation, landfilling and/or open field burning (1 -3). * Corresponding author. Phone: 54-1 1-45763383. Fax: 54-1 145763366, E-mail: [email protected] 1116
Slow pyrolysis has been recognised as a convenient process to attain high yields of the solid product (char). The char has potential as a solid fuel. In addition, activation of the char to produce activated carbons may represent another economically interesting option. Commercially available activated carbons are highly porous carbonaceous materials with high BET surface areas typically in the range fiom 400 to 1500 m2/g. These properties are responsible for the large adsorption capacities characterising activated carbons which are commonly used to remove pollutants from gaseous and liquid streams. Char properties are important to assess its potential use for specific applications. They depend strongly on the raw material employed as feedstock and on pyrolysis conditions. Despite the significant progress that has been achieved, it is still difficult to predict with confidence char properties for a given lignocellulosic waste from available information in the literature. In particular, char surface properties are relevant to activated carbon production by the "physical" activation two-stage method. This method involves pyrolysis of the precursor and subsequent activation, i.e partial gasification, of the resulting char employing steam, C02 or flue gases as the most common activating agents (4). On the other hand, efficient conversion of lignocellulosic wastes into char requires knowledge of pyrolysis kinetics of each waste. Regardless of the final char application, pyrolysis kinetics plays a key role for the design of commercial reactors, depending on the reactivity of each waste (5-6). Pyrolysis of lignocellulosic wastes is complex due to the different degradation behaviour of major constituents, i.e. cellulose, hemicellulose and lignin, interactions between constituents, and effects of minute amounts of mineral matter naturally present in whole biomass samples (7-8). As a consequence, a generalized knowledge of pyrolysis kinetics of any lignocellulosicmaterial has, not been yet achieved. In this context, the present work aims at examining comparatively: (1) chemical compositions, heating values and surface properties of char samples obtained from slow pyrolysis of different lignocellulosic wastes, in relation to their potential as biofuel andor for further processing to produce activated carbons; (2) pyrolysis kinetics of the selected wastes, necessary for the design of the reactors for chzir production. EXPERIMENTAL SECTION
Different lignocellulosic wastes extensively generated in wood and agricultural products processing in Argentina were selected. The wastes employed were sawdust from two different wood species, Guatambti (Aspidosperma australe), a native species, and Poplar (Populus delfoide), husks of peanut (Arachis hypogaea), and olive (Olea L. europaea) stones free from pulp. They are respectively denoted in the text as AA, PD, PH, and 0s. The wastes were milled and screen-sieved. Wastes composition was determined in terms of the major constituent biopolymers, holocellulose (cellulose + hemicellulose) and lignin, and extractive components that are soluble in ethanol-benzene. Lignin and extractives were isolated according to TAPPI standard methods. Holocellulose was obtained following the procedure described elsewhere (9). For charcoal production, fiactions of 1.3 x 10'3m (1.3 mm) average particle diameter were used. The wastes were thermally treated under flowing nitrogen in a fixed bed reactor heated by an electrical furnace. The samples were subjected to a heating rate
1117
of 0.25 "C s" up to 850°C, and held at this temperature for 1 hour. Afterwards, the resulting chars were cooled under nitrogen to room temperature, weighted, and carefblly stored for characterisation. Typical yields of the chars ranged between 15 and 40%, depending on each waste. Proximate analyses of the wastes and char samples were conducted following the ASTM standard methods. An elemental analyser (Carlo Erba EA 1108) was used to determine elemental compositions. Surface properties of all the samples were determined fiom physical adsorption measurements by volumetric techniques, using N2 at -196 "C and COz at 25 OC. The complementary use of these adsorbates allows characterisation of complex networks of pores of different sizes (5). A Micromeritics Gemini 2360 surface area analyser was employed for N2adsorption experiments, whereas adsorption measurements performed with C02were carried out in a Micromeritics Accusorb 2lOOE sorption instrument. All the samples were outgassed overnight at 120°C at a final pressure of 1.33 x 1O4 Pa (1 0" mm Hg) prior to adsorption measurements. Kinetic measurements of pyrolysis of the different wastes were performed by nonisothermal thermogravimetric (TG) analysis, fiom ambient temperature up to 900°C.A thermogravimetric balance Netzsch STA 409 coupled with a nitrogen flow device and a data acquisition system was used for the kinetics measurements. From a preliminary set of experiments, sample mass, particle diameter, and gas flow rate, for which diffusional effects are negligible, were determined. Nitrogen flow rate of 2.5 x lo5 L s-l, sample masses of 10&0.2 mg, and 37-44 pn particle diameter were employed for all the experimental runs. Experiments at different heating rates, in the range 0.17-1.7"Cs-' were also carried out. No appreciable differences in mass loss against temperature curves were found for the heating rates examined, pointing to negligible heat transfer effects (10).
RESULTS AND DISCUSSION CHEMICAL COMPOSITION,HE4 TING VALUES AND SURFACE PROPERTIES OF THE SAMPLES
Contents of holocellulose and lignin, five fiom extractives, are shown in Table 1. Table I Content of main constituentsof the lignocellulosic wastes.
Raw Material
% Holocellulose
% Lignh
AA sawdust
73
PD sawdust Peanut husks Olive stones
65 74
26.6 25.4 28.8 44.3
-
'
#Expressed on an extractive-free basis
As already mentioned, holocellulose and lignin percentages, reported in Table 1, were
determined independently by different experimental methods. Thus,values may not add
1118
to lOO?!. Olive stones possess a considerably larger percentage of lignin compared to the other selected wastes. Tables 2 and 3 report proximate and ultimate analyses of the wastes and derived char samples, respectively. Heating values (HV) of the wastes and chars were estimated by an own multivariate correlation in tenns of samples' elemental compositions, previously reported (5):Estimated heating values on a dry-ashfiee basis are listed in Table 3.
Table 2 Proximate analyses (dry basis) of the lignocellulosic wastes and char samples. Sample
Volatile Matter
Ash
Fixed Carbon*
(%I
(W
(%I
72.7 8.1 71.4 7.4 84.9 38.8 74.8 22.1
I .o 7.2 0.5 4.6 1.7 4.2 0.4 1.6
26.3 84.7 28.1 88.0 13.4 57.0 24.8 76.3
AA sawdust AA-Cha PD sawdust PD-Char Peanut husks PH-char Olive stones OS-char *Estimated by difference,
Table 3 Ultimate analyses and estimated heating values (HV) of the lignocellulosic wastes and char samples.
'
Sample
%c
%H
%N
%0*
AA sawdust AA-Char PD sawdust PD-char Peanut husks PH-char Olive stones os-char
48.2 92.8 48.3 89.6 47.4 84.6 46.6 94.4
5.6 1.3 5.8 0.8 6.1 0.8 6.0 0.5
0.9 0.8 0.8 0.3 2.1 0.9
45.3 5.1 45.1 9.3 44.4 13.7 47.4 4.3
0.8
w(MJ/kg) 19.01 29.65 19.14 27.40 19.39 25.80 18.8 1 27.67
# All the results are expressed on a dry-ash free basis. *Estimated by difference.
Pyrolysis of the wastes promotes pronounced changes in chemical compositions, as reflected in the results obtained for the char samples that are strongly dependent on each waste. Release of volatile matter leads to enhance the inorganic material of the samples, the char from AA sawdust showing the highest ash content (Table 2). All the wastes are enriched as fuels during pyrolysis. Increases in %C and decreases in o/oo may be 1119
appreciated in Table 3. Variations in chemical compositionsof the selected lignocellulosic wastes due to pyrolysis are in accordance with well-known general trends (5, 11). Char samples possess high carbon contents (84 - 95%), negligible contents of sulphur, and nitrogen percents lower than 1%. Similar heating values are found for the different wastes. As expected, all char samples show higher heating values compared to the parent wastes owing to the enhanced %C, the differences in HV among the chars becoming larger. Heating values are comprised between 25 MJkg and 30 MJkg. Char fiom AA sawdust shows the highest HV, followed by the char samples produced fiom olive stones and PD sawdust. The PHchar presents the lowest HV. Heating values of the wastes and chars are within the range reported in the literature for this kind of biomass and derived chars (1 1-13). Surface properties of the samples were examined from nitrogen and carbon dioxide adsorption measurements. Figure 1 illustrates typical N2 adsorption isotherms determined on the char samples produced fiom the different wastes, providing information about samples larger pores, mainly macropores, mesopores and larger micropores. Nitrogen adsorbed volumes expressed in standard conditions of temperature and pressure (STP) per sample mass unit, V,, as a h c t i o n of the relative pressure @/po) are shown in the figure.
6o 50
I 1
/'
40
+AA-char 30 /
+P +P
D -char
+O
S-char
H -char
20
10
0
0
0 2
0 4
PIP,
0 6
0 8
I
Figure J Nitrogen adsorption isotherms for char samples derived fiom the different lignocellulosic wastes.
Differences in the shapes of the isotherms are found (Figure 1). The 0s-char shows a flat plateau which characterises type I-isotherms according to IUPAC classification, pointing to a predominantly microporous structure. N2 volumes adsorbed on the AA-char are considerably larger compared to those determined on the other wastes. Results indicate
1120
that char samples possess different pore networks, that seem to depend on the original botanical structure inherent to the type of waste used for char production (1 0). The BET model was applied to fit N2 isotherms and evaluate the specific surface area of the samples ( S B ~ )following , the conventional procedure (14). Estimated BET surface areas are detailed in Table 4. Table 4 Surface properties of the lignocellulosicwastes and derived chars, as estimated from N2 and Cot adsorption data applying BET and Dubinin-Radushkevich equations.
,x
"-..
...
..I.I.............. ..........
""
AA sawdust AA-char PD sawdust PD-char Peanut husks PH-char Olive stones OS-chU
.
I .6 100.0 1.3 21.0 1 .o 13.0 0.6 17.0
..
.
..
..
.....
~
.............,._
87.5 3.2 63.1 23.3 107.0 41.1 88.3 42.2
140 320 82 490 107 534 53 717
Micropores in the lignocellulosic wastes and resulting chars were analysed f?om C 0 2 adsorption isotherms by applying the Dubinin-Radushkevich(DR) equation (14). DR plots log
(4)
0 s
~
?.-
---t A A - c h a r
+PD-char
+P H - c h a r +O -0 5
--
-I 5
--
-2
S - c h ar
-I
0
2
4
log'(p/p,)
8
10
I t
14
Figure 2 Dubinin-Radushkevich(DR) plots for char samples resulting h m the different lignocellulosicwastes, as calculated fiom COzadsorption data.
1121
for the char samples are illustrated in Figure 2. In this figure, the amounts of C02adsorbed (4) expressed as a liquid volume per sample mass unit are plotted against l o g @/PO). As can be seen fiom Figure 2, COz adsorption data for all the samples are well described by the DR- equation. Consequently, the total micropore volume (+J for each sample was obtained from the ordinate at the origin of the DR plots. I$, values were then used to calculate the apparent surface areas (Sco2) according ti, the evaluation procedure described in detail elsewhere (14). The estimated COz surface area for each sample is listed in Table 4. The ratios between the nitrogen and carbon dioxide surface areas (SCO&ET)are also included in the same table, for the sake of comparison. ~ SCOZareas than the wastes. Release of volatile Char samples show larger S Band matter during pyrolysis of the wastes, as a result of the complex network of reactions that takes place, promotes the development of new pores and/or favours widening of preexistent pores. Some restructuring in the char matrix may cause additional increases in microporosity leading to larger SCOZvalues (10). All the samples possess BET surface areas considerably smaller than those calculated from COz adsorption data, pointing to very narrow micropores or obstructions of pore entrances that restrict N2 diffusion.Thus, differences between values of SBET and ScOZ suggest prevalence of micropores over meso and macropores in the lignocellulosic wastes and chars. The relative contribution of micropores to the samples’ structures may be additionally inferred fiom the !!&/SBET ratio. Char samples present smaller ratios than the wastes, indicating decreases in the relative proportions of micropores owing to the thermal treatment. Results indicate that the AA-char has the largest BET surface area and smallest amount of micropores compared to the char samples ftom the other wastes, pointing to a more open structure. The SCOZ/SBET ratio for this char is one order of magnitude smaller, On the contrary, the char from olive stones shows the largest amount of micropores and ratio suggesting a relatively close, compact structure. Similar, the highest Smz/SBET intermediate proportions of micropores, as judged h m Sco2 areas, characterise the chars produced fkom peanut husks and PD sawdust, although the S C O ~ / Sratio B ~ for the latter is smaller. The OS-char obtained under similar pyrolysis conditions to those employed in the present study has proved suitable for production of high-quality activated carbons, leading to BET surface areas larger than 1100 m2/gby subsequent partial gasification using carbon dioxide as activating agent (15). Activated carbons of microporous character were obtained. Hence, properties of the char samples from the other selected wastes are compared with those exhibited by the OS-char in an attempt to predict potential conversion into activated carbons of reasonably good quality. Although SBET and Sm areas for the OS-char differ from those determined for the rest of the char samples, the PH-char shows the closest SCO~ISBET ratio to that of the OSchar, pointing to a similar relative contribution of pores of different sizes to the structure, followed by the PD-char. The AA-char possesses an interesting large BET surface area, ratio is considerably smaller than that corresponding to the OS-char. but the Scoz/SBET Large BET surface areas and sorption capacities of activated carbons arise from the presence of significant amounts of micropores (4). Thus, apparently, the AA-char should not lead to a pronounced development of porosity through further activation due to its relatively more open structure. Heschel and Klose (16) reported that coarse-cellular structures, such as those characterising high porosity woods, are disadvantageous as
1122
activated carbon feedstock. Nevertheless, the AA-char might be usefbl for preparation of activated carbons presumably with a smaller .BET surface area but with a larger average pore size that could be appropriate for the removal of large size compounds in liquidphase. The ash content of char samples is another feature that has to be taken into account since activated carbons with low ash contents are normally desired. The low ash content of the olive stone-char should be emphasised. From this viewpoint, the relatively higher ash content of the AA-char compared to char samples from peanut husks and PD sawdust is disadvantageous.
KINETICS OF THE WASTES PYROLYSIS Figures 3 and 4 illustrate typical thermogravimetric (TG) curves obtained for pyrolysis of the different lignocellulosic wastes at a heating rate of 0.83 "C 6'. The TG curves show the recorded mass losses, expressed as weight fraction (w), against the temperature (T). Weight fractions are given by w=m/n, where m, mo are the instantaneous and initial mass samples, respectively. Pyrolysis rates (-dw/dt) for the different wastes were estimated by differentiation of the weight fraction-time curves. In Figure 5 , the pyrolysis rates are represented as a k c t i o n of temperature. w
1
0.9
0.8
0 7
0 6
0
A A sawdust
n
PD
-M
snwdusl
odcl prcdictionr
0 5
0 4
0 3
5 0
100
150 200
250 300
350
400
450
SO0
550
600
650 700
750
800
8SO
900
T ['Cl
Figure 3 Thermogravimetric curves of pyrolysis of sawdust fiom both wood species, Aspidosperma australe (AA) and Populus deltoide (PD): weight fraction (w) against temperature (T). Comparison between experimental data (symbols) and model predictions (solid lines).
1123
W
...
"
I
I
I
0 9
0 8
0 7 I
0 6
A 0
-M
0 5
Peanut husks Olive slones ode1 predictions
0 4
0 3
0 1 -~
! 0 I
i
--
I
-dw/dt [seg"] 0.012
0.01
I
1
i
!
ni \
sawdust
----AA
0 008
--
- .- P D
-Peanut
-------
sawdust husk
0 l i v e sto n e s
0.006
0 004
0.002
0
50
1 0 0 I50 2 0 0 2 5 0 3 0 0 3 5 0 4 0 0 4 5 0 5 0 0 5 5 0 6 0 0 6 5 0 7 0 0 7 5 0 8 0 0 8 5 0 9 0 0
T ["CI
Figure 5 Influence of the temperature on the pyrolysis rates for the different lignocellulosic wastes.
1124
Pyrolysis of the wastes becomes substantial above 150°C. Maximum pyrolysis rates for the selected wastes are attained in an interval of temperatures ranging between 290°C and 320°C. At higher temperatures, pyrolysis slows down. This behaviour may be attributed to the thermal degradation of major biopolymers constituting the lignocellulosic wastes. At low temperatures, degradation of cellulose and hemicellulose having a polysaccharide structure takes place. As temperature increases, degradation of lignin, which has a predominantly aromatic structure and is known to show more resistance to pyrolysis than holocellulose, seems to become progressively more prevalent. Mineral matter composing the wastes also appears to play a role affecting the pyrolytic behaviour (8, 17). The highest maximum rate is found for pyrolysis of sawdust from Poplar wood, while lower values are determined for the other wastes (Figure 5). A model published in the literature (1 8) was used in order to fit experimental data and evaluate kinetic parameters over the whole range of degradation temperatures investigated, up to 900°C. The model assumes that pyrolysis of the wastes occurs as a frst-order overall decomposition. It considers that the physical and chemical changes which take place within the solid as pyrolysis proceeds cause solid deactivation. The reaction rate is given by: (-dw/dt) = kapp(W- w,) Changes within the solid are assumed to affect the reaction rate constant (kapp)and are taken into account through an increase of the activation energy with the temperature and the solid conversion according to: kapp= ko exp [-EAO( 1+ p T F ) / RT] where X is the normalised fiactional conversion, given by: X=(l-w)/(l-w,) w, w,, are the instantaneous and residual weight fractions, respectively. ko is the preexponential factor, EAO, the initial activation energy, for X=O, and p, y, fitting parameters (p, the deactivation rate, y, the order with respect to X). R, T, are the universal gas constant and the absolute temperature, respectively. For model application, fitting of TG curves (w vs T) was carried out taking into account the measured linear relationship between temperature and time. As mentioned in the experimental section, 60m preliminary experiments it was verified that different heating rates in the range 0.17-1.7"C s-l led to similar w against T curves. Model parameters were estimated by non-linear regression analysis. The sum of squares between the experimental and calculated instantaneous weight fractions was minimised. Estimated values of the model parameters together with standard deviations (SD) are detailed in Table 5. Model fit to experimental results for sawdust from both wood species and agricultural wastes is shown in Figures 3 and 4, respectively. As can be seen in these Figures, the model appropriately describes experimental data over the wide range of temperatures for all the selected wastes. Other models were also applied. However, they
1125
were unable to describe experimental results for pyrolysis of all the wastes, led to larger SD, or adequately represented the data only for restricted ranges of temperature (1 0). Table 5 Estimated characteristic parameters of the kinetic model.
AA sawdust
Poplar sawdust Peanut husks Olive stones
59.4 59.1
3.70
35.0 67.0
0.17 340
12.4
8.0 1.30 0.0023 8.4 1.28 0.00089 5.0 0.58 0.00095 7.0 0.56
0.0022
Pyrolysis of sawdust 60m the two wood species is characterised by similar values of initial activation energy. In contrast, larger differences in the estimated EAO values for pyrolysis of the agricultural wastes are found. The degradation of olive stones shows a relatively higher value of initial activation energy, presumably due to its larger lignin content (Table 1).
CONCLUSIONS Characterisation of chars produced from selected lignocellulosic wastes by slow pyrolysis under identical operating conditions is carried out. Distinctive compositions, heating values and surface properties of the chars are determined. All the chars obtained are found potentially useful as relatively pollution-6ee solid biofuels. Char samples possess high carbon contents (84-95%), negligible contents of sulphur, and nitrogen percents lower than 1%. Ash contents are also sufficiently low, with the exception of the AA-char which has a relatively higher ash percent (~7%).Heating values lay between 25 MJkg and 30 MJkg. The char fiom sawdust of the native wood species (Aspidosperma australe) shows the greatest potential as biofuel, followed in a decreasing order by the chars arising from olive stones, Poplar wood sawdust, and peanut husks. The higher ash content of the AA-char might be an inconvenient for this application. Comparison of surface properties of char samples with those characterising the OS-char, which has proved suitable for preparing high quality activated carbons through subsequent activation, suggests that chars from peanut husks and Poplar sawdust could be adequate candidates for further conversion into activated carbons with well-developed surface areas. Nevertheless, other factors that are relevant to manufacture and applications of activated carbons should be further examined. Pyrolysis kinetics of all the selected lignocellulosic wastes is properly described over the wide thermal degradation range 25"C-9OO0C by a model that considers an increasing dependence of the activation energy on the temperature and waste conversion with the process course. Appreciable differences in the estimated kinetic parameters are found.
1126
ACKNOWLEDCMENTS The authors gratefully acknowledge Consejo Nacional de Investigaciones Cientificas y TBcnicas (CONICET), Universidad de Buenos Aires (UBA) and Agencia Nacional de Promoci6n Cientifica y Tecnol6gica (ANPCYT) from Argentina for financial support. REFERENCES 1. Bridgwater, A.V., Bridge, S.A. (1991). A review of biomass pyrolysis and
2.
3. 4.
5.
pyrolysis technologies. In: Biomass Pyrolysis Liquids Upgrading and Utilisation. (Ed. by A. V. Bridgwater & G. Grassi), pp. 11-92. Elsevier Applied Science, London and New York. Rashid Khan, M., Gorsuch, C. (1996). Utilisation of waste materials: examples of business successes. In: Conversion and Utilisation of Waste Materials. (Ed. by M. Rashid Khan.), pp 3-13.Taylor and Francis, Washington DC. Karaosmanoglu, F., Isigigur-Eradenler, A., Sever, A. (2000). Biochar fiom the straw- stalk of rapeseed plant. Energy and Fuels 14,336-339. Patrick, J.W. -(Ed.) (1995). Porosity in Carbons: Characterisation and Applications. Halsted Press, New York and Toronto. Cukierman, A.L., Della Rocca, P.A., Bonelli, P.R., Cassanello, M.C. (1996). On the study of thermochemical biomass conversion. Trenak in Chemical Engineering
3, 129-144. 6. Della Rocca, P.A., Cerrella, E.G., Bonelli, P.R., Cukierman, A.L. (1999). Pyrolysis of hardwoods residues: on kinetics and chars characterisation. Biomass and Bioenergy 16,79-88. 7. Antal Jr, M.J., Varhegyi, G. (1995). Cellulose pyrolysis kinetics: the current state of knowledge. Ind. Eng. Chem. Res. 34,703-717. 8. Caballero, J.A., Conesa, J.A., Font, R., Marcilla, A. (1997). Pyrolysis kinetics of almond shells and olive stones considering their organic fractions. J Anal. Appl Pyrolysis 42, 159-175. 9. Browning, B. I. (1970). Methods of Wood Chemistry, Volume 2, Wiley, J. (Ed) New York (Chapter 19). 10. Della Rocca, P.A. (1998) Study on biomass thermal conversion processes. DPhil
thesis, Universidad de Buenos Aires. 11. Minkova, V., Marinov, S.P.,Zanzi, R., BjUmbom, E., Budinova, T., Stefanova, M., Lakov, L. (2000). Thermochemical treatment of biomass in a flow of steam or in a mixture of steam and carbon dioxide. Fuel Processing Technology62,45-52. 12. Raveendran, K., Ganesh, A. (1996). Heating value of biomass and biomass pyrolysis products. Fuel 75, 1715-1720. 13. Di Blassi, C., Signorelli, G., Di Russo, C., Rea, G. (1999). Product distribution fiom pyrolysis of wood and agricultural residues. Ind. Eng. Chem. Res. 38, 22 162224. 14. Gregg, S.J., Sing, K.S.W. (1982). Adsorption, Su@uce Area and Porosi@.
Academic Press Inc., London. 15. Horowitz, G.I., Lombardi, G., Cassanello, M.C., Cukierman, A.L. (1997).
Gasification of an agricultural by-product: a study towards activated carbons
1127
preparation. In: Proc. 41h Asian Pacijic International Symp. on Combustion and Energy Utilisation. Chulalongkorn University, Bangkok, pp 757-76 1. 16. Heschel, W., Klose, E. (1995). On the suitability of agricultural by-products for the manufacture of granular activated carbon. Fuel 74, 1786-1791. 17. Raveendran, K., Ganesh, A., Khilar, K, C.(1995). Influence of mineral matter on biomass pyrolysis characteristics. Fuel 74, 18 12-1822. 18. Balci, S., Dogu, T., Yilcel, H. (1993). Pyrolysis kinetics of lignocellulosic materials. Ind. Eng. Chem. Res. 32,2573-2579.
1128
The Pyrolysis Kinetics of a Single Wood Particle Davidsson, K. 0.' Pettersson, J. B. C, Bellais, M. ', Liliedahl, T. and SjCistrom, K. I Department of Chemistry, Physical Chemistry, Goteborg University, SE-412 96 Goteborg, Sweden Department of Chemical Engineering and Technologv, Royal Institute of Technology, SE-10044 Stockholm, Sweden
'
ABSTRACT: Experimental results from birchwood and pinewood pyrolysis in a new single particle reactor are presented. Apparent lunetic parameters for the mass-loss of wood particles (5-800 mg)at temperatures from 300 to 860 'C are determined. Kinetic parameters for the evolution of CO, COZ, H20, H2 and CH, are also established. The drylng process was examined and it was found that drying and pyrolysis increasingly overlap in time as temperature rises and that the overlap is substantial above 450 OC.
INTRODUCTION Increasing environmental concern, as to the global warming, has made renewable energy sources for combustion and gasification an advantageous option. If the use of coal were shifled towards increased use of biomass the COZ-emission would decrease. To perfonn such a shift, existing and future power plants have to be (re)constructed for biomass conversion. Although biomass was mans first fuel the large-scale energy production from biomass is rather a new issue, especially compared to coal, where one can rely on decades of experience. Hence there is a need for knowledge about biomass behaviour in large-scale conversion. In the present paper wood pyrolysis will be considered, since wood is the most lmportant biomass in many areas. The first step in the combustion process is pyrolysis, which is the production of combustible volatiles as the temperature of the solid fuel increases. Wood pyrolysis is a complex phenomenon due to the composition and structure of the fuel. Despite this, many models have been proposed for the pyrolysis of a single wood particle. Bamford [l] coupled the kinetics of the pyrolysis with the heat conduction within the wood. Kung [2] included internal transfer by heat convection of the volatiles and reaction endothermicity. Kansa et al. [3] included Darcy's law. Chan et al. [4], Saastamoinen and Richard [5], Bilbao et al. [6] and Melaaen and Gr0di [7] have considered the drymg process. Other effects such as shrinking [8,9] and exterior conditions [ 101 have been considered. Di Blasi [ 1 13 made a two-dimensional model of the wood pyrolysis. Although these models may be regarded crude on a molecular level they may be too complicated to be used for modelling of a power plant reactor. Wood
1129
pyrolysis experiments have been performed with TGA (therinogravimetric analysis) applying temperature ramps 112-141. An exception is Petek et al. [15] who pyrolysed wood particles by injection feeding.to a thermobalance preheated to 500-950 OC.This way of heating resembles the heating of a wood chip in large-scale conversion. A few experiments have been made where the time-resolved gas evolution is treated [12,14,15, 17-19]. It is not evident to us that the behaviour of a thexmally thick wood particle, upon entering a hot reactor, is predictable from standard TGA-studies because of the differences in heating processes. In a TGA-study one typically finds that a heating rate of a few to a few hundred degrees per minute has been applied. This is to be compared with the heating of a particle rapidly moved from 20 OC to a reactor of 400 or even up to 800 "C. We have chosen wood, rather than cellulose or lignin, because it is unclear whether wood can be described by simple combination of the pyrolysis behavioui. of its components. Antal[16] clearly states that this is not the case. To contribute to the knowledge of the behaviour of a single wood particle upon entering and pyrolysing in a large-scale reactor, experiments in a new single particle reactor were performed. The wood particle (0.6-800 mg, cubic or cylindrical) falls through a cooled tube from room temperature onto a plate inside the reactor surrounded with gases of 300-850 "C. The mass is recorded and the light gases emitted are measured with mass spectrometry yielding time resolved data. The single particle pyrolysis in a preheated environment resembles the situation for a particle in a largescale reactor, even though the heat transfer and attrition due to interaction with other particles and bed material are not accounted for. The purpose of the present paper is hence to obtain apparent kinetic parameters for mass-loss and gas emission for a wood particle under a range of conditions during pyrolysis.
EXPERIMENTAL The experimental system consists of a single particle reactor, a mass spectrometer (Baltzers QMG 421C) and a balance (Sartorius BP 21 1 D) and is shown schematically in Fig. 1 . In the middle of the reactor is an A1203-plate on which the sample is pyrolysed. It is surrounded by five U-shaped heaters, which can heat the reactor up to 1000 "C.The temperature of the reactor is measured with thermocouples (TC type K) close to the plate, which is connected to the balance. Above the plate is a cooled insertion tube, through which the sample is brought. Level with and a few mm from the plate is situated the capillary, through which pyrolysis gases are sampled to the mass spectrometer. The capillary consists of an A1203-tube,which is inside the reactor, and a heated deactivated silica tube (diameter = 0.23 mm) between the reactor and the mass spectrometer. Visual observation and other measuring techniques are also possible. Preheated inert gas at atmospheric pressure enters through holes in the bottom of the reactor. Balance- and mass spectrometer data are recorded and represent time resolved mass and emission of light gases respectively. In the present case tars were not collected. The A1203-plate is heated from the inert gas flow and no large temperature gradients occur over the distance between the plate and the TC. Therefore it is reasonable to assume that the gas and the plate are of the same temperature as the TC. At higher temperatures radiation becomes increasingly important so that the TC and the plate may become warmer than the gas but at these temperatures radiation is also more important than convection and conduction for the pyrolysis.
1130
vertical view
horizontal view
Fig. I Horizontal and vertical views of the experimental set-up. 1:mass spectrometrycapillary, 2: to the balance, 3: insertion tube, 4: heater, 5: pre-heated gas, 6: A1203-plate (for the sample); The figure is not entirely to scale.
At 400 "C heat transfer to the particle was calculated. The energy required to evaporate 85% of the water in moist particles corresponds to an average heat transfer of 69 kWm-'. Birchwood (Betula verrucosa) was mainly used in the experiments but a few experiments were performed with pinewood (Pinus silvestrii). Both these trees are common in northern Europe. Samples of cubic and cylindrical forms and masses from 0.6 to 800 mg were used. The masses have been divided into four groups and the shape of the particles in each group is given in Table 1. Table I Description of sample sizes and shapes mass
shape
1 mg cubic lmm
5 mg cubic 2mm
60-100mg cylindrical 5x5mm and cubic 5mm
600-800 mg cylindrical 4 10 mmxl5 mm
The samples were usually dried but a series of experiments was carried out with moist samples. Elemental analysis showed that both birch and pine contained 50% C, 6% H, 44% 0 and 0.1% N with respect to dry mass. A typical experiment was carried out in the following manner. The sample was cut, weighed and dried at 105 OC until it had constant mass. It was brought onto the plate by falling through the insertion tube. The reactor was held at a constant temperature and had been evacuated with respect to oxygen (
1131
Mass spectra were analysed and it was found that no mass to charge ratio ( d z ) above 84 was registered. Further analysis also showed that m/z = 2, 15, 18, 28 and 44, to a major part, correspond to Hz,CH4,H20, CO and C02respectively. To examine the effect of the sample plate on the results, other plates than the A1203-platewere used. A three mm thick Cu-plate with relatively large heat capacity, maximising the heat transfer between the plate and the sample, and a plate consisting of a Pt-net suspended on thin Alz03-supports, so that the sample rested on the net with close to free convection around, were used. No significant effect of the sample holder was found. To examine the effect of the ambient atmosphere a few experiments were carried out in N2. No difference compared with experiments in Ar could be found.
RESULTS The results were obtained under Ar-atmosphere with dried birchwood samples, if not otherwise stated. The evaluation of data can be subdivided into that concerning sample mass and that concerning gas emission.
SAMPLE MASS The pyrolysis of a wood particle entering a hot zone from room temperature can roughly be said to proceed in two phases. During the first phase, the heating phase, the particle is heated so that water, if any, evaporates and the temperature rises to about 200 OC at which devolatilisation of wood is known to start. Above this temperature, the mass-loss phase starts, where devolatilisation and consequently mass-loss occur. The rate of the heating phase is governed by external and internal heat transfer. The massloss phase is kinetically or heat transfer controlled. It should be noted that the temperature at which pyrolysis occurs may differ fiom the ambient. Both phases show strong temperature dependence, which can clearly be seen in Fig. 2, where mass as a function of time at different temperatures is shown. The initial mass of the birchwood sample was 80-100 mg cubic. The temperature dependence of the pyrolysis process is evident from the leftmost diagram, where it can be seen that the pyrolysis decreases fiom hours to less than ten minutes as the temperature increases from 315 to 400 "C.
1
1 .o-
0.8:
U
-.-8 0.6f 0.4s 0.28
8
'
500 OC
0.0 0
200
400
600 800 Time (9)
1000
1200
lo
20
j,
40
M,
I
3
Time ( 8 )
Fig. 2 Mass as a function of time for different temperatures during pyrolysis of an 80100 mg cubic dried birchwood sample.
1132
1.0-
In
2
Q
= 0.9,5.2,96.5 and 633.3mg
= 6.5,99.7 and 598.6 mg
?% I,-
0.8-
U
-.z0
0.60.40.2-
0.07 0
100
io
300
200
$0
40
Time (5)
60
Ibo
1 0
Time (s)
Fig. 3 Mass as a function of time for different initial sample masses at 400 and 600 "C during pyrolysis of a cubic dried birchwood sample. In the leftmost diagram the curve for mQ = 0.9 mg has been smoothed. Fig. 3 shows mass as a h c t i o n of time for different initial sample masses. A small sample is heated faster than a large one, implying a shorter heating phase but the mass-loss phase is also shorter. The char yield for different samples is shown in Fig. 4. The yield ranges from almost 40% to less than 10%. An increase in reactor temperature (in this temperature range) obviously causes a decrease in char yield. A hgher reactor temperature means a higher heating rate and a higher final temperature. Both these parameters have been shown to have a reducing effect on the char yield [17,20,21], with final temperature apparently being the most important. Also the size of the sample affects the char yield so that a smaller size yields less char. Two reasons may be given for this. Smaller samples experience larger heating rates and therefore also higher average temperatures. The other explanation would be that in a small particle, tars flow through a thin char layer compared with a large particle, reducing the importance of secondary reactions [22], i. e. tar cracking into volatiles and char re-building. The effect of particle size on char yield is less pronounced at hgher temperature.
-
50o dried birch 60-110 rng
: o 5 . 0 .z* 30'g%& :A;% c" 20: L.
0
0
40-
. 10-
A
x
dried birch 600-800 rng dried birch 5-7 rng dried pine 55-65 rng
8 A A
O0
4? A o 3
& @ 8 & Ak 0 %
0 7 " " 1 " " 1 " " 1 " ' .
I
"
"
I
"
"
Temperature (C") Fig. 4 Char yield vs. temperature for different sues of dried birchwood samples and one size of dried pinewood samples.
1133
In order to investigate the mass-loss process of the particle, apparent rate coefficients from a first-order expression were obtained. The following expression was applied: m
- mJ e-" + m,,
= (mo
(1)
where m, mo and m, are mass, initial mass and char mass (mass at infinite time) respectively and t is time.These parameters were obtained from experiments, while the rate coefficient, k, was obtained by fitting of Eq. 1 to experimental data.Assuming that the rate coefficient can be described by the Arrhenius expression
where A , E and T are pre-exponential factor, activation energy and temperature respectively, the activation energy can be found as can be seen in Fig. 5 . Due to the heating phase preceding the mass-loss, a kinetic first order expression cannot comprise the complete process. Despite this, the value of the rate coefficient at different temperatures gives an idea of the rate-limiting step and the time it will take to pyrolyse the particle at different temperatures. The activation energy below 370 OC is in agreement with those found for the pyrolysis of the constituents of wood, i. e. cellulose, hemicellulose and lignin [20,23-261 indicating that the pyrolysis kinetics are rate determining. The explanation for this is that, at low temperatures, the heating phase is very fast compared to the mass-loss phase, which may take hours. Above 370 OC, the activation energy is about 30 Hmol-' for all sample sizes except the smallest ones. This low activation energy cannot correspond to a chemical reaction. We contribute this to external heat transfer limitation. The activation energy in this temperature range can be said to reflect the temperature dependence of the heat transfer rate. 1 -1
V
1
V
x -3-
C
-5-7-
dried birch 600-800 mg 0 dried birch 60-1 10 mg A dried birch 5-7mg 0
v dried birch e 1 mg x dried pine 5585 mg
-9
I
I
Fig. 5 Arrhenius plot according to a first-order mass-loss expression for pyrolysis of dried birchwood particles. Least square fits to dried birchwood samples of 60-1 10 mg are indicated in the figure. Above 370 O C , E = 3 1 kJmor' and A = 3 s-'. Below 370 "C, E = 177 kJmol-' andA = 1.9.1Ol2s-'.
1134
Figure 5 indicates that in the range 5-800 mg the size of the particle does not affect the apparent activation energies, implying that the temperature dependence of the limiting process is the same within these ranges of temperature and size. A series of experiments was made with pinewood, and a comparison of the pyrolysis of pine and birch is made in Fig. 6. The qualitative differences of the massloss are small but reproducible. Birch loses weight a little faster during the heating phase. This may be attributed to higher content of bound water or some other easily gasified substance like hemicellulose. The compositions of birch and pine are listed in Table 2.
Table 2 Characteristicsof birch and pine by mass-% [26]. substance cellulose hemicellulose lignin
Scandinavianbirch
Scandinavianpine
40
44
39 21
26 29
Hemicellulose is a thermally unstable compound and is known to start to pyrolyse at 200°C [23,27,28]. During a short period of the mass-loss phase the mass of birch is higher than that of phe but the char yield of pine is slightly higher. The differences are small and may be attributed to different heat transfer but because of the independence of temperature (in the range 400-700 "C) it is more likely that different structure and composition explains the difference in mass history. The evaluation of the kinetic parameters, conducted as above, shows no difference fkom birch (see Fig. 5).
-birch ---pine
0
50
100
150
200
I
2 i0
T i m (s) Fig. 6 Mass vs. time during pyrolysis of dried pine- and buchwood samples of equal sizes (55-100 mg) at two different temperatures.
GAS EMISSION. The gas emission during the pyrolysis of birchwood was analysed by use of mass spectrometry. Mass-loss and the emission of selected gases (mass spectrometer ion current) during pyrolysis at 500 OC and 800 OC are shown in Fig. 7. The gas emission 1135
curves correlate in time, as expected, with the mass-loss curve. This observation is evidence of the short response-time for the experimental set-up. As the temperature increases, more CO and H2 are evolved compared with the other gases in accord with Beaumont and Schwob [17]. The reason for this may be a change in pyrolysis chemistry inside the particle with temperature. This agrees with the observation that the char yield decreases with temperature [17,20,21]. Antal[16] presented a scheme for the thermal decomposition of lignocellulosic materials. This scheme included a high activation energy fragmentation of the solid and it would occur at very high heating rates and yield CO and H2.In the present work at high reactor temperatures at least the outer layers of the particle will be heated very fast and may undergo this reaction. Another explanation would be that reactions outside the particle change with temperature. For example H20 and CO may react to give H2 and C02 and its equilibrium is shifted towards higher concentrations of H2 and C02 as temperature rises. This can though not fully explain our observations since we do not observe an increase in CO2. However, the reaction between CO and H2to produce CH, and H20 is shifted towards CO and H2 with increasing temperature. This could partly explain the observations even though we see no decrease in C&-concentration. A first-order expression was applied to the gas emission &ta: g i = l - eR ,
(3)
where giis the relative amount of gas i released at the time t (amount of gas released at time t divided by amount of gas released at infiite time). Rate coefficients were obtained by fitting of Eq.3 to experimental data. Arrhenius diagrams for the emission of all the selected gases showed that the activation energy is around 30 kJmol-', i. e. close to the one obtained for the mass-loss above 370 "C. The Arrhenius-diagram for CO is given as an example in Fig. 8. The low activation energy, similar to the activation energy of the mass-loss, also implies that the measured gas emission is a product from within the particle and not the result of gas phase reactions outside. This means that the gas phase reactions discussed above do not have a major effect on the observations.
1136
0.100 4.0~10-~O 0.075
3.0~ 10 -lo
3 P)
u) (I)
0.050
2.0x10-'0
S Y
0.025
1.ox10-10
.I25
n
4.0~10~~
5
s
U
3.0~10~
t
"Ca
-
2.0x104B
0
1.0~1049
0.0x10-
lime (s) Fig. 7 Mass and MS signals as functions of time during pyrolysis of dried birchwood particles at 400 and 800 O C .
E, = 33 kJmol-' A = 3 s-'
-71 0.0008
0.0010
0.0012
0.0014
0.0016
0. 318
l/T (I/K) Fig. 8 Arrhenius-plot of CO emission during pyrolysis of cubic dried birchwood particles of approximately 100 mg.The line corresponds to a least square fit. Kinetic parameters are given in the figure. 1137
Drying Experiments were performed where the birchwood particle was dried and weighed and thereafter immersed in water until it had a moisture content of 50-70 % on dry basis. It was then pyrolysed as described above. A comparison between pyrolysis at 400 OC of a moist and a dried birchwood sample is given in Fig. 9, where the mass has been normalised to initial dry mass. This figure illustrates the main features of pyrolysis of moist wood. It is clear that the conversion takes longer time due to the longer drying period. The mass-loss phase (due to pyrolysis) is also somewhat slower for the moist particle but the char yield is very similar in the two cases. At this temperature and below, a distinct drying phase is observed but at higher ambient temperatures the w i g and pyrolysis tend to overlap as illustrated in Fig. 10, where the mass spectrometer signal for H20 is given for the pyrolysis of samples at different temperatures. Much of these fmdings agree well with Di Blasi et a1 [29]. At 400 "C one clearly sees two peaks at about 50 seconds and 130 seconds. These peaks correspond to drying and devolatilisation respectively. At 500 OC one can see a small peak at 85 seconds corresponding to pyrolysis but this peak is strongly overlapping with the drying peak. At 600 "C there is practically only one peak.
400 OC n
s
1.5~10
-1.5
z 4
zg
I .
1.ox10
-1.0
5.OX10
-0.5
u)
3
c 0
I
o.ox10
0
50
100
150
200
250
300
350
u) u)
400
Time (s) Fig.9 Comparison of mass as a function of time for a dried and a moist birchwood particle (mo(dry) = 63 and 96 mg) during pyrolysis at 400°C. The MS signal from water is given for the moist particle. Initial dry mass = 1.
1138
Ir !
-400
'I I$
' ' (I
3.0~10-'~- 1 5
OC
.._.._ 500 O c
\
Fig. 10 MS-ion current as a function of time for H20at different temperatures during pyrolysis of moist birch particles of 60-100 mg initial dry weight.
CONCLUSIONS AND FUTURE WORK The main results and conclusions from this work can be surmnarised as follows: (1) Apparent kinetic parameters have been obtained for mass-loss and noncondensable gas evolution. The activation energy, above an ambient temperature of 370 "C,was found to be about 30 Idmol-' under the present circumstances. The activation energy for the mass-loss below 370 "C agrees with those found in the literature for the kinetics of the main wood constituents. (2) Dried Birch and pine behaves somewhat differently and the difference is qualitatively independent of ambient temperature. It is likely that differences in structure or composition, rather than hfferences in heat transfer, explains this. (3) The evolution of CO and H2 is enhanced by increased temperature and may be attributed to the fragmentation pathway of lignocellulosic materials mentioned by Antal [16]. The trends for the other gases are not as clear but evolution of C& shows an increase whereas a slight decrease can be observed for COzand water. (4) Water is emitted in two phases. We conclude that the first phase is evaporation of free water and that the second phase is water produced in the pyrolysis reactions. These phases are fairly separated up to 400 OC, above, which they increasingly overlap. Increasing the temperature increases the importance of the internal heat transfer limitation of the processes. Thus, at high temperatures, a drying wave precedes a pyrolysis wave. If the particle is large, both phenomena occur at the same time. This, is important from a simulation point of view: for high temperature, the overall pyrolysis cannot be simulated by a two-step model, namely a drymg model and a pyrolysis model. A global model featuring both sets of equations has to be solved.
1139
As shown in the preceding parts, kinetic parameters cannot be directly calculated when internal heat transfer limits pyrolysis. A model taking into account both kinetic scheme and heat- mass transfers becomes necessary. A one-dimension model has already been implemented and solved. It features a classical set of equations for heat and mass transfers in porous media, i.e. heat transfer through convection, conduction, radiation and mass transfer due to pressure gradient (Darcy’s law) and binary diffusion. Different kinetic schemes from the literature are and will be tested mass-loss as lumped first order reaction, formation of volatiles, tars and char from decomposition of cellulose, hemicellulose and lignin [26],the Broido-Shafuadeh model [30]and the one proposed by Di Blasi [3 11. None of them can describe the composition of the volatiles and supplementary schemes have to be found. Drying could be simulated in a separated model. However, since it was shown that drying and pyrolysis occur at the same time for large particles and high temperatures, the model should solve simultaneously an equation set for drying and one for pyrolysis. The model also needs to be extended to two dimensions due to the wood structure and the particle foxm. ACKNOWLEDGMENT We would like to thank Mr. Benny Liinn for support and construction work. This project was financed by The Swedish National Energy Administration, project No P 10444. REFERENCES
1. 2. 3.
4.
5. 6. 7.
8. 9.
Barnford C. H., Crank J. and Malan D. H. (1946)The combustion of wood. Proc. Camb. Phil. SOC. 42, 166-82. Kung H. C. (1972) A mathematical model of wood pyrolysis. Combustion and Flame 18, 185-95 Kansa E. J., Perlee H. E. and Chaiken R. F. (1977)Mathematical model of wood pyrolysis including internal forced convection. Combustion and Flame 29,3 1 1-24 Chan W.-C. R., Kelbon M. and Krieger-Brocket B. (1985) Modelling and experimental verification of physical and chemical processes during pyrolysis of a large biomass particle. Fuel, 64,1505-13 Saastamoinen J. and Richard J. R.(1996) Simultaneous drymg and pyrolysis of solid fuel particles. Combuston and Flame, 106,288-300 Bilbao R. Mastral J. F. Ceamanos J and Aldea M. E. (1996) Modelling of the pyrolysis of wet wood. J. Anal. Appl. pVrolysis, 36, 81-97 Melaaen M.C and Grsnli M. G. (1997)Modelling and simulation of moist wood drying and pyrolysis. proc. Dev. Thermochem. Biomass Conversion, 1, 132-46 Villermaw J., Antoine B., Lede J. and Soulignac F. (1986) A new model for thermal volatilization of solid particles undergoing fast pyrolysis. Chemical Engineering Science, 40,No. 1, 15 1-7 Di Blasi C. (1996) Heat momentum and mass transport through a shrinking biomass particle exposed to thermal radiation. Chemical Engineering Science, 51, NO.7 , 1121-32
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10. Miller R. S. and Bellan J. (1996) Analysis of reaction products and conversion time in the pyrolysis of cellulose and wood particles. Combust. Sci. and Tech., 119,331-73 11. Di Blasi C. (1997) A transient, two-dimensional model of biomass pyrolysis. Proc. Dev. in ThermochemicalBiomass Conversion 1, 147-60 12. Bilbao R., Arauzo J., Murillo M. B. and Salvador M. L. (1997) Gas formation in
13. 14.
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20. 21.
the thermal decomposition of large spherical wood particles. J. Anal. Appl. Pyrolysis, 43,27-39 Varhegyi G . and Antal M. J., Jr. (1989) Kinetics of the thermal decomposition of cellulose, hemicellulose, and sugar cane bagasse. Energy & Fuels, 3,329-35 Szab6 P., Vhrhegyi G., Till F. and Faix 0. (1996) Thermogravimetric/mass spectrometric characterization of two energy crops, Arundo donax and Miscanthus sinensis. J. Anal. Appl. Pyrolysis, 36, 179-90 Petek J., Rummer B., Seebauer V., Steiner G. and Staudinger G. (1999) Thermal decomposition of large wood particles in inert and reactive atmosphere: mathematical simulation and experiments with rapid heat up. 6th international conference on circulating fluidized beds. In: Circulating Fluidized Bed Technology VI, (Ed. by J. Werther), pp. 469-474, Dechema e. V. Antal M. J. (1985) Biomass pyrolysis: A review of the literature, part 2lignocellulosic pyrolysis. in Advances in Solar Energy (Ed. by Boer K. W. and DuMie J. A) 2, 175-255. Plenum Press, New York Beaumont 0. and Schwob Y. (1984) Influence of physical and chemical parameters on wood pyrolysis. Ind. Eng. Chem. Process Des. Dev., 23, 637-41 Ohlemiller T. J., Kashiwagi T. and Werner K. (1987) Wood gasification at fire level heat fluxes. Combustion and Flame, 69, 155-70 Varhegyi G., Antal M. J. Jr., Szekely T., Till F., Jakab E. and Szab6 P. Simultaneous Thennogravimetric-mass spectrometric studies of the thermal decomposition of biopolymers. 2. sugar cane bagasse in the presence and absence of catalysts. Energy &Fuels, 2,273-7 Min K. (1977) Vapor-phase thermal analysis of pyrolysis products fiom cellulosic materials. Combustion and Flame, 30,285-94 Antal J. A. Jr., Mok W. S. L., Varhegyi G. and Szekely T. (1990) Review of methods for improving the yield of charcoal fiom biomass. Energy & Fuels, 4,
NO.3,221-5 22. Connor M. A. and Salazar C. M. (1988) Factors influencing the decomposition processes in wood particles during low temperature pyrolysis. In: Research in Thermochemical Biomass Conversion, pp. 164-78 (Ed. by A. V. Bridgwater and J. L. Kuester). Elsevier, London. 23. Bar-Gadda R. (1980) The kinetics of xylan pyrolysis. Thermochimica Acta, 42, 153-63 24. Nguyen T., Zavarin E. and E. M. Barall (1981) Thermal analysis of lignocellulosic materals. J.Macromol. Sci.-Rev. Macromol. Chem. C20, 1-65 25. Kanury M. and Blackshear Jr. P. L. (1970) Some considerations pertaining to the problem of wood-burning.Combustion Science and Technology, 1,339-55 24. 26. Grranli M. (1996) A theoretical and experimental study of ,the thermal degradation
of biomass. PhD Thesis, Norges teknisk-naturvitenskapligeuniversitet Tronheim 27. Rajeswara T. R. and Sharrna A. (1998) Pyrolysis rates of biomass materials. Energy, 23, No. 11,973-8 28. Roberts A. F. (1970) A review of kinetics data for the pyrolysis of wood and related substances. Combustion and Flame, 14,261-72
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29. Di Blasi C., Hernandez E. G . and Santoro A. (2000) Radiative pyrolysis of single moist wood particles. Znd. Eng. Chem. Res. 39,873-882 30. Shafizadeh F. (1982) Chemistry of pyrolysis and combustion of wood. Progress in Biomass Conversion, 3,5 1-76 3 1. Di Blasi, C. and Russo, G . (1994), Modeling of transport phenomena and kinetics of biomass pyrolysis. In.Advances in ThermochemicalBiomass Conversion. (Ed. by A.V. Bridgwater) pp. 906-921. Blackie Academic & Professional, London
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Dynamics and products of wood pyrolysis Colomba Di Blasi, Carmen Branca, Antonio Santoro, Elier Gonzalez Hernandez and Raul Albert0 Perez Bermudez Dipartimento di Ingegneria Chimica, Universitii degli Studi di NapoIi "Federico12" - P.le V: Tecchio, 60125 Napoli - Italy
ABSTRACTThe radiative pyrolysis of wood (thick cylinders and chip beds) has been investigated experimentally for external radiative heat fluxes in the range 28-80kW/m2, resulting in maximum sample temperatures of 600-950K.Radial temperature profiles, product yields and composition, and devolatilization rates have been measured. The influences of wood variety (hardwoods and softwoods) on the pyrolysis characteristics are discussed and comparisons are made with biomass (agricultural residues).
INTRODUCTION A high number of studies is available on the pyrolysis behaviour of wood particles 1.11 .For large part of the investigations, measurements are concerned with the time history of temperature at different radii or stations along the axis of the sample, while in a few cases, mass loss curves, profiles of solid density and gas overpressure have also been measured. For complete process characterization all these measurements are necessary and, for applications in thermochemical conversion, especially yields (and composition) of the main product classes, i.e. gases, char and liquids. However, these have been measured only in a very few cases 7 ~ 1 0 ~ 1.1In addition, the measurements produced by the different laboratories are not directly comparable because of the differences in reactor configurations (heat transfer mechanism from reactor to particle), wood variety (different physical properties and chemical kinetics) and sample characteristics (size, shape and moisture content). Therefore, fhdamental information is not currently available for optimal feedstock pretreatment and selection, in order to guarantee efficient operation of conversion units and to reduce the emission of noxious substances. Also, the applicability of mathematical models Iv3 for process design and development is highly dependent on the validation stage, which requires sets of data measured over wide temperature intervals. In this study results of an extensive experimental activity are presented on the pyrolysis of wood (chips and thick cylinders) for a wide range of external heating conditions and several varieties, namely, two hard-woods (beech, chestnut) and three soft-woods (Douglas fir, redwood and phe).
1143
EXPERIMENTAL SYSTEM The experimental system applied for the pyrolysis of wood cylinders is the same presented in [ 1 1,15,16] and a schematic representation is reported in Fig. 1, which also includes details of the heater and the reactor. The hmace, manufactured by Research Inc., presents four tubular quartz infrared lamps with a tungsten wire filament that emits radiant energy in proportion to the applied voltage. Elliptical, polished aluminium, water-cooled reflectors focus the high-density infrared energy, emitted by lamps, onto a cylindrical shaped target area (diameter 6 . 5 ~ 1 0 'm). ~ The furnace, equipped with a PID controller and a transducer (SCR),allows a constant radiative heat flux to be emitted. To avoid interaction between the volatile pyrolysis products and the lamps, a quartz tube (ID 6x10" m), transparent to infrared radiation, is located inside the furnace and used as a reaction chamber. A wood cylinder is vertically positioned in the uniformly heated zone of the reactor, through a suspension system, which is connected to a precision balance. The sample is exposed to the same radiative heat flux along the lateral surface. For each chosen radiation intensity, steady temperatures of the radiant heater are achieved within a couple of minutes (maximum heating rates of about 750Ws)but, given the thick sample, pyrolysis takes place under heat transfer control. A nitrogen flow is applied at the top of the quartz reactor in order to reduce the extraparticle residence time of volatile pyrolysis products and to establish the proper reaction environment. A nominal (ambient conditions and absence of sample) gas velocity of 0.0125ds has been employed for all the tests, which corresponds to volatile residence times along the heated section of about 6s. Two separate tests are made, one for the continuous recording of the sample weight loss and another for temperature measurement, product collection and gas analysis. Temperatures along the particle radius, r, at the median section are continuously monitored (0.5~10"m bead chromel-alumel thermocouples) at five positions starting fiom the center. The composition of the gas is analyzed, at selected times, through a gas-chromatograph equipped with a TCD and a packed column. These measurements are also applied to evaluate the yields of non condensable gaseous components (indicated as "gas"), through integration of the concentration of each species over the time of the experiments. After complete conversion, the power is turned off and the solid residue is left under a nitrogen flow until its temperature lowers to ambient values. This residue is indicated as "char". The liquids are collected through a condenser train, consisting of water-cooled traps, cotton wool demisters and a silica gel bed. All the condensable products collected and weighed from the traps (organic compounds (tars) and product water formed) are indicated as "liquids".
RESULTS The tests have been carried out for two hardwoods, beech (Fagus sylvatica) and chestnut (Castanea sativa), and three softwoods, Douglas fir (Pseudotsuga menziesii), redwood (Sequoia sempervirens) and pine (Pinus pinea). Chemical analyses are reported in Table 1. Extractives are determined by means of the extractor Soxhtec HT2. Acetone (60ml) is used as solvent ( l g of biomass) with residence times for the boiling and rising stages equal to 90 and 20min, respectively (temperature equal to 363K). Lignin is determined according to the Klason method and holocellulose is obtained by difference.
1144
The chemical composition of Douglas fir and redwood is about the same (29-33% lignin, 6556% holocellulose and 6- 1 1% extractives, respectively), but differences are larger with pine wood, which presents much lower lignin contents (24%) with 69% holocellulose and 7% extractives. Differences exist also between the two hardwoods, with the chemical composition of beech consisting of 20% lignin, 78% holocellulose and 2% extractives. Chestnut is characterized by a high content of extractives (1 6%) associated with 18% lignin and 66% holocellulose. Differences are also observed in the initial wood density. In a first set of experiments, pyrolysis has been carried out of wood cylinders (grain direction is parallel to the longitudinal axis) with 4cm diameter and length, after drying at 373K for 8-10h. The applied radiative heat flux has been varied in the range 2880kW/m2,so as to achieve steady surface temperatures, T,, varying fiom about 600K to 950K. Measurements have been made of radial profiles of temperature, yields of the main product classes and composition of the gas. For beech, chestnut and pine, weight loss curves have also been recorded.
t [SI Fig.2A Solid mass fraction and time derivative of the solid mass fraction for a heat fluOf Q=3 lkW/m2 and different wood chips varieties (beds with 4cm diameter).
-
1.o
0.9 0.8
*
0.7
0.6 ---.--_ ..-
t
[SI
Fig.2B - Solid mass fraction and time derivative of the solid mass fraction for a heat flux of Q=49kW/mzand different wood chips varieties (beds with 4cm diameter).
1145
Beech Chestnut Douglas Redwood fir 2 16 6 11 extractives % 20 18 29 33 lignin % 78 66 65 56 holocellulose %
Variety
Pine 7 24 69
Table I Chemical composition of wood samples.
The reproducibility of the tests is good, as maximum variations in the stead Y temperatures of 10-20K are observed only for applied heat fluxes above 4OkW/m , probably as a consequence of structural failures and volume variations, while variations in the weight loss dynamics are always negligible. Weight loss curves and temperatures profiles have also been measured for wood chips beds (after drying) for low (31kW/m2) and high (80kW/m2) applied radiation intensities. In order to privilege the influence of the chemical composition, similarity in the physical behaviour of the packed beds has been accomplished by working with small particles (0.5-lmm) and the same bed density (0.212glcm3) . The sample consists of about 8g arranged to give a bed with 4cm diameter and 3cm height. It is worth noting that, as found in previous studies (for instance 14), for a chosen bed density, particle sizes do not influence the pyrolysis characteristics, indicating that the heat transfer resistance across the bed predominates over that across the particles. WEIGTH LOSS DYNAMICS OF WOOD CHIPS BEDS
From the qualitative point of view, the behaviour of the different wood chip beds is the same, but large quantitative differences exist, as shown by comparing weight loss dynamics in Figs. 2A-2B. For low heat fluxes (Fig2A), maximum devolatilization rates are observed for the two hardwoods, whose degradation process also terminates earlier. More precisely, for times shorter than 300s, the degradation rate of chestnut is faster, but for longer times it is beech wood which attains the fastest rate. The degradation rates are comparable for all the softwoods. For intermediate (Fig.2B) and high heat fluxes, apart from very short times, when the degradation dynamics of the different varieties are very similar, it is still beech wood which attains the highest rates. The most interesting result is the similar behaviour of redwood and Douglas fir, which degrade with much slower rates than the hardwoods, while the degradation behavior of pine wood is, in some way, comprised between that of beech wood and chestnut. In other words, close similarity exists between beech and pine, on one side (fast devolatilization rates), and redwood and Douglas fir, on the other (slow devolatilizationrates). Figures 2A-2B also show that the lowest residues are detected for beech, whereas the highest solid residues are obtained for the other hardwood variety, chestnut. Though at low temperatures the solid residue is high for all the softwoods, as soon as temperatures become sufficiently high, pine wood residues are significantly lower (and practically coincide with those of beech wood for very high temperatures) than those of the other two softwoods. Finally, the conversion times, defined as the time when the devolatilizationrate lowers to 1/10 of the maximum value, show the shortest values for the two hardwoods. Times are significantly longer for all the softwoods. Hence, contrary to the trends shown by the solid residues, close similarity exists between the hardwoods, on one side, and the softwoods, on the other.
1146
900
flfifil
049
8001 700
Q=28kW/m
1600t 500
//M
4( -\
C)
3 0 O 0 T ’
400
-n
ni
1--v.v I I11
I
I
1
I
I
800
1200
1600
2000
1 2400
t bl Fig. 3 - Time-temperature profiles at several radial positions for beech wood cylinders exposed to Q=28 and 49kWlm’ .
snn -~
r
800 700 600 500
400
300 0
400
800
1200
1600
2000
t [SI Fig.4 - Time-temperature profiles at the sample centerline for chestnut (dashed lines) and beech (solid lines) wood cylinders exposed to several radiation intensities. 900
1
500 -
Beech
-.- -- - - - Redwood I
0
400
1200
800
t
1600
2000
[sl
Fig.5 - Time-temperature profiles at the sample centerline for redwood (dashed lines) and beech (solid lines) wood cylinders exposed to several radiation intensities.
Given the same heating conditions established for the different varieties of wood chips, differences in the degradation dynamics are mainly due to the variable amounts (Table 1) and nature of the main components, though other factors,such as interactions
1147
between the component fractions, catalytic actions by the mineral matter content, may play an important role. Beech, on one side, and Douglas fir and redwood, on the other, can be considered as classical examples of hardwoods (low lignin contents) and softwoods (high lignin contents), respectively. The high extractive content of chestnut and the relatively low lignin content of pine introduce significant deviations with respect to the basic characteristics of the two wood categories. Compared with beech wood, chestnut degradation starts at shorter times, global devolatilization rates are lower and solid residues significantly higher. These features can be, for large part, attributed to extractive degradation lo, whereas the higher temperatures needed by the softwood categories to attain fast de radation rates can be due to the higher degradation temperature of their lignin fraction 19 . HEATING AND WEIGHT LOSS DYNAMICS OF WOOD CYLINDERS The macroscopic behavior of degrading wood cylinders is again qualitatively similar for the different varieties and examples are given for beech wood at low (28kW/m2) and high (49kW/m2) radiative heat fluxes (Fig.3). The temperature profiles show the existence of two main regions: the most external is essentially determined by heat transfer effects and the second by reaction energetics. For the first region, heating rates become slower as the distance from the external surface increases, because of the increasing internal heat transfer resistance, but the degradation of the main wood components takes place at about the same time. For the inner core of the sample, due to the slower heating rates of the final conversion stage, the endothermic degradation of the low-temperature components (extractives and holocellulose) is completed before the high-temperature and exothermic degradation of lignin 15*16. Indeed, there is an almost flat region which corresponds to the first endothermic stage. The subsequent exothermic lignin degradation gives rise to local maxima in the time derivative of temperature. The temperature dynamics remain qualitatively the same but the size of the second region becomes successively smaller as the externally applied radiation intensity is increased. In other words, as the spatial temperature gradients increase, endothermic and exothermic processes become more coupled, given that they take place at the same time for a successively larger portion of the sample. The simultaneous degradation of all the wood components results in a local process nearly isothermal, dominated by transport phenomena. Given the high contribution of the low temperature components in the wood compositions (67-82%, Table l), their initial degradation temperature, as reported by the centerline thermocouple, can be considered as representative of the pyrolysis process, T, . On the other hand, lignin degradation takes place over a wide range of thermal conditions ' ' ~ 3 and its exothermicity does not allow a clear and univocal definition of a reaction temperature. The time needed to attain the temperature T, can be considered as the characteristic heating time, th, and can be used to evaluate the average heating rate as P=(Tp:TO)/th, where To is the initial temperature. Though, from the qualitative point of view, the dynamics of sample heating are the same for the different wood varieties, significant differences exist for the pyrolysis temperature and the average heating rate and, consequently, for the minimum heat flux needed to achieve significant wood degradation (i.e. volatile yields at least equal to 50% of the initial dry solid mass). Figures 4-5 compare examples of the time histories of temperature at the sample centerline, while Figure 6 reports T, and p as hnctions of the applied radiation intensity. For both temperature dynamics and parameter values,
1148
heat fluxes of about 40kW/m2separate a region of strong dependence from a region of weak dependence. On the whole, it can be observed that the average heating rates increase with the external radiative heat flux and roughly reproduce those of thermal analysis. The average heating rates are, for large part, determined by pyrolysis temperatures which, after a rapid increase (Q=40kW/m2), approach constant values. With reference to maxima, the lowest T, values (about 600K) are observed for chestnut, as a consequence of the high content of extractives and hemicellulose, which degrade at temperatures lower than cellulose. The highest Tp values are measured for beech wood (about 650K). Softwoods, in particular redwood and Douglas fir, present pyrolysis temperatures slightly lower than beech wood. The higher levels of extractives and the complete degradation of their hemicellulose content within a narrow temperature range lo may be the factors responsible for these features. 660
1.1
0.9
620
0.7
'e33
0.5
cn,-
580 Y
F
bn 540 500 25
0.3 35
45
55
65
75
0.1
85
Q [kW/rn2] Fig. 6 Pyrolysis temperature, T,, and average heating rate, 0, as functions of the applied radiative heat flux for different wood varieties (cylinders).
t
bl
Fig. 7 Time profiles of the solid mass fractions and time derivatives of the solid mass fraction for chestnut (dashed lines) and beech (solid lines) wood cylinders exposed to several radiation intensities.
1149
* t [SI Fig. 8 Time profiles of the solid mass fractions and time derivatives of the solid mass fraction for pine (dashed lines) and beech (solid lines) wood cylinders exposed to several radiation intensities.
Differences in the temperature dynamics among varieties are high especially for Q=28kW/m2, when heating rates are slow and thermal conditions barely adequate for degradation. The effects of variable contribution andor reactivity of the components are enhanced also in relation to process energetics (exothermic lignin degradation). Thus large differences appear between the two hardwoods (Fig. 4) and between hardwoods and softwoods (Fig. 5). The very low pyrolysis temperature of chestnut causes that, for a given heat flux, the region controlled by reaction energetics is always larger than for the other wood varieties. In particular, for the lowest heat flux investigated, it regards the whole sample radius, as the degradation of extractive and holocellulose components, on one side, and lignin, on the other, take place into two separate stages. Therefore, despite of the low lignin content, its separate degradation is able to cause a significant temperature increase above the steady (final) value (Fig. 4), with effects on both product distribution and weight loss dynamics. From Fig. 5, where a classical hardwood (beech) and a classical softwood (redwood) are compared, it appears that the second exothermic stage is completely absent for the latter when Q=28kW/m2 . In accordance with component analysis lo, this is due to higher degradation temperatures of softwood lignin. As a consequence, higher applied radiation intensities are needed in order to cause a softwood conversion above 50%. A comparison between the weight loss dynamics for the two hardwoods considered in this study can be made through Fig.7. For beech wood, the apparent kinetics of weight loss can be described as a one-stage process, for Q=40kW/m2 . Indeed, the temperature gradients are small, the size of the degrading region is large and maximum degradation rates are attained for heat-up conditions. Devolatilization characteristics are, on the whole, qualitatively similar to those discussed for woodchips beds. However, the complete or significant separation between the endothermic and exothermic degradation stages results in two different regions in the devolatilization rate curve. Indeed, a first wide region of low values (low-temperature component degradation) is followed by a narrow second zone of higher values (lignin degradation). Due to the significant exothermicity of the second stage, large part of the devolatilization is completed at times shorter than for beech.
1150
For Q> 40kWlm’ and both hardwood varieties, the process tends to become globally two stage, with the first due to surface degradation and the second to sample heat up. Indeed, the fast rise of the surface temperature causes the prompt devolatilization of a thin layer, below the irradiated surface, while the increase in the local heating rates and the spatial temperature gradients cause the simultaneous degradation of the different wood components. Again, chestnut degradation starts at shorter times (so the appearance of a first peak associated with surface degradation is displaced towards higher Q values), global devolatilization rates are significantly lower, but conversion time is about the same as for beech wood. That is, degradation dynamics of the two hardwoods tend to become the same as the reaction conditions are made severe. Figure 8 compares the global degradation characteristics of beech and pine woods. On the whole, it appears that pine devolatilization starts at shorter times and ends at times longer than beech (again with the same trends as for wood chips). The higher content of extractives, on one side, and the lower reactivity of softwood lignin, on the other, may account for these features (including longer conversion times). Devolatilization rates are, on the average, lower and the process is characterized by slightly higher solid residues. Also, given the lower density of pine wood, spatial temperature gradients are lower and the devolatilization rates remain relatively constant for large p h of the degradation process. Again, differences between the two different wood varieties are reduced for high external radiation intensities. An evaluation of the importance of internal heat conduction time versus external heat transfer time, based on the external radiation intensity, can be made by means of the Biot number, Bi, and the use of average medium properties. Bi values are comprised between about 10 and 40 1 1 ~ 1 5for the range of heating conditions considered, indicating that internal heat transfer is controlling and it also becomes successively more important as Q is increased (higher values for softwoods). The Biot number is defined as Bi=(aQR)/(AAT) ,where average values are used for the medium properties (a=0.8, h=1.4x10-’ W/mK (hardwoods 15) and h=0.93x10-’ W/mK (softwoods Is) and temperature difference, AT , referred to the pyrolysis temperature). Also, the characteristic time associated with internal heat transfer can be compared with the characteristic time of chemical reaction, through the thermal Thiele number, Th, and average devolatilization rates 1 1 , 1 5 and medium properties. The thermal Thiele number is defined as Th=(kp,c,RZ)lh) (again average values for the medium properties are used with p,=450kg/m3 for hardwoods and ps=310kg/m3for softwoods, c,=1.25 kJkg K and, k ,the average devolatilization rate). Computed values of Th vary.from 0.4 to 2.3 and, for all the woods, applied heat fluxes of about 40kW/m2 roughly separate a region where chemical reaction times are slow (sample temperatures are hardly sufficient for wood degradation) from another where internal heat transfer becomes relatively more important. Heat and mass transfer processes are affected by volume variation and structural changes undergone by wood while degrading, though these effects on the conversion characteristics are difficult to quantify. Sample shrinkage, mainly along the radial direction, appears as a reduction in the sample diameter by factors of 15% for beech wood and Q<40kW/m2 . For higher heat fluxes, structural failure occurs at a large extent and precise evaluations are not possible. A small reduction in the sample diameter (about 2-3%) is also observed for the other varieties and Q=40kW/m2 . For higher values, successively larger swelling occurs. For chestnut, it corresponds to an increase in the sample diameter by factors of 3-9% and it is also accompanied by
1151
formation of cracks. These are completely absent for sofiwoods, which present swelling factors of 340% (pine and Douglas fir) and 10-20% (redwood).
PYROLYSIS PRODUCTS Product (char, liquids and gas) yields from wood cylinder pyrolysis are reported, as percent of the initial dry solid mass (%, wt), as functions of the radiation intensity in Figs. 9-11. The mass closure is'good (95.5-99%) and, from the qualitative point of view, all the wood varieties show the same trend. Again two regions are shown with distributions highly dependent on the heating conditions for external heat fluxes below 40kW/m2 (sample temperatures below 750K).For this region, as the external heat flux is increased, the final char yield decreases whereas volatile (gas and liquid) yields increase, as a result of a competition between charring and devolatilization reactions, which become successively more favored 18,19. As secondary reaction activity is considered to be negligible for temperatures below 750K ('I* *'), product distribution is essentially the result of the selectivity of primary wood decomposition reactions.
25
35
45
55
65
5
75
Q [kW/m2]
Fig.9 - Char yields fiom wood cylinders pyrolysis as functions of the radiative heat flux.
redwood
0
rn -
chestnut
f 302Ls
I
I
1
I
1
35
45
55
65
75
Q [kW/m2]
85
Fig.10 - Liquid yields fiom wood cylinders pyrolysis as functions of the radiative heat flux.
1152
I
L
25
55
40
70
85
Q [kW/m2] Fig. I I - Gas yields from wood pyrolysis cylinders as functions of the radiative heat flux. For radiative heat fluxes above 40kW/m2, liquid yields present a barely visible maximum, due to secondary tar reactions as, for the experimental conditions of this study, these effects are small also for temperatures above 750K (small activity only within the pores of the degrading solid). Indeed, intra- and extra-particle residence times of volatile specie are short, as a consequence of the relatively fast devolatilization rates and the negligible intra-particle resistance to mass transfer, on one hand, and to the forced nitrogen flow, on the other. In the second place, the volatile temperatures are much lower (values always below 550K ' I ) than those reported by thermocouples in strict contact with the solid phase. On the whole, all the wood varieties show the same qualitative dependence on the heating conditions, which, as pointed out in 'I, is also in agreement with previous experimental and modelling analyses. As for the solid residue (char), it should be observed that the high yields reported at low heat fluxes also include some unreacted material. In particular, sample temperatures are not sufficiently high to cause complete degradation of the ligninic components, especially for softwoods. For the whole range of thermal conditions investigated, the highest char yields are measured for chestnut and the lowest for beech wood, whereas softwoods attain intermediate values. For the most severe conditions, when the volatile content of the char is at a minimum 20, yields vary from about 21% (beech and pine) to 27% (redwood and Douglas fir) and 33% (chestnut). Consequently wood variety causes variations in the char yield up to 12% with respect to the initial dry wood mass. An inspection of the liquid yields as functions of the radiative heat flux reveals that the higher the char yields are the lower the liquids. The maximum liquid yields, attained for applied heat fluxes of 49-69kW/m2, vary from about 56-57% (beech, Douglas fir and pine) to 52% (redwood) and 47% for chestnut. Hence, the wood variety may cause variations in the optimal liquid yields up to 10% of the initial dry wood mass. Variations in the gas yields with the wood variety are lower than for the other two product classes. Indeed, for the highest heat flux (and maximum yields) they are about 5% of the initial dry solid wood. The highest yields are measured for beech, pine and chestnut (21-1 8.5%), with lower values for redwood and Douglas fir (17-16%). The local maximum observed for chestnut and Q=28kW/m2 is due to the separate degradation of the low-temperature components and lignin for the whole sample,
1153
resulting in a second high exothermic stage with maximum temperatures well above the final steady value (Fig. 4), causing significant solid devolatilization.
w.w
650
700
750
850
800
900
950
1000
1050
T, [KI Fig.12 - Ratio YL/(YC+YG)(ration of liquid yields to char and gas yields) as a function of the surface temperature in the pyrolysis of wood cylinders and packed biomass beds (for both cases a diameter of 4cm) exposed to radiative heating. 9.0 Redwood 0
7.0
8 2, iz
Graperesidue
v Douglas fir 0 Olive husks
:::
A
Chestnut
A
Ricehusks
1.0
d
-1.0 -3.0
6.5
6.6
6.1
6.8
6.9
7.0
(T,) Fig.13 - Logarithms of the ratio CO/CO2, CO/CH4and CO/C2 yields as functions of the logarithm of T, (steady surface temperature) for different wood varieties (cylinders) and packed biomass beds (for both cases a diameter of 4cm) exposed to radiative heating.
Chemical composition may explain, for large part, the differences in product distribution among the different wood varieties. Previous studies 21 report that extractive fiee wood produces up to 10% higher liquid yields than untreated wood. It was postulated 22 that the layer of extractives which surrounds the wood fibers acts as a barrier in the mass transfer of volatile pyrolysis products, out fiom the reacting sample. The increase in the residence time of tar vapors inside the high temperature pores of the degrading wood was thus considered responsible for an enhanced activity of secondary tar degradation reactions, with higher char or gas yields 21. The results of this study are in qualitative agreement with such findings given that the lowest liquid yields are observed for the varieties with the highest extractive content (chestnut and redwood).
1154
Also this feature is associated with high char (chestnut) or both char and gas (redwood with respect to Douglas fir, both belonging to the same variety and with very similar composition) yields. However, the explanation based on secondary reaction activity could be consistent only for high temperatures, while the above characteristics are observed for the whole range of heat fluxes. Hence, it is plausible that the products from the degradation of extractives, especially if these are present in high amounts, may contribute significantly to product distribution, that is, to the formation of char andor gas, at the expense of liquid products. Specific features may depend on the chemical pattern of extractives which vary greatly between the different wood species and also between different parts of the same tree. Finally, in accordance with thermogravimetric results (for instance lo), ligniri is the main responsible for char formation, as redwood and Douglas fir produce higher yields of this product than pine and beech. Figure 12 shows that, for reaction temperature above 650K,the ratio between the liquid yields and the yields of the other two product classes slightly increases with temperature (woods) or remains roughly constant for biomass. Indeed, for comparison purposes, results previously obtained for several agricultural residues I 4 (packed beds of the same size as wood cylinders) have been examined. Hence, once the process becomes controlled by internal heat transfer, increasing the external temperature has a certain effect only on the relative amounts of gases and char (lignin degradation is more favored). Ratio values become successively lower with holocellulose content, with significant differences mainly between chestnut wood and the other woods, and also between the different biomass species. The pyrolysis gas consists mainly of C02, CO, CH4 and lower amounts of H2 and C2 hydrocarbons (details reported in l 4 for biomass and in l5 for wood). At low temperatures, the evolution of CO and C 0 2 (and water vapor) is due mainly to the degradation of extractives, hemicellulose and the activity of the frst path in cellulose degradation, leading to gas and char formation. As temperature increases, the formation of tar vapor from cellulose becomes predominant, while lignin degradation also attains fast rates. The last process is largely responsible for char formation and also for the evolution of C02, CO, CH4 and H2 . From the quantitative point of view, the two hardwoods and pine (all including low lignin contents) are characterized by C 0 2 yields higher than redwood and Douglas fir. Yields of CO and C& are comparable for the different woods, apart from the slightly lower CH4 and C2 yields for chestnut, caused by the low lignin content. Previous pyrolysis studies, for instance 20, report very general, approximate relationships for the ratios of the major gas phase products, CO, CH4, C02 and C2 hydrocarbons, over wide ranges of experimental conditions (temperatures). These are due to comparable values of the activation energies for the formation of the different species, with reaction mechanisms representative of apparent kinetics, taking into account both primary and secondary reactions. As apparent kinetics are also important in relation to large wood samples, a plot is constructed of the logarithms of the ratio of yields of major gas-phase species on the logarithm of the steady surface temperature, T,, and comparisons are again made with agricultural residues 14 . Examples for CO/C02, CO/CH4 and CO/C2 are shown in Fig.13. It can be seen that a linear dependence is roughly established, confirming that major gas species present comparable kinetic constants (in particular activation energies) for both the different wood and biomass varieties.
1155
CONCLUSIONS
The influence of wood variety (beech, chestnut, Douglas fir, redwood and pine) on the conventional pyrolysis characteristics has been investigated for a wide range of applied radiation intensities (from 28 to 80kW/m2), which correspond to final sample temperatures of 600-950K. The first difference between the pyrolytic behaviour of hardwoods and softwoods is the higher initial (minimum) heat flux required by the latter category for achieving conversion levels above 50% (about 33kW/m2 against 28kW/m2). This is a consequence of the higher degradation temperature of softwood lignin and the failed contribution of its reaction exothermicity for thermal enhancement. The lower reactivity of this component also results in slightly longer conversion times for softwoods. For all the wood varieties, the temperature field shows an outer and an inner region, along the cylinder radius, determined by the simultaneous and the sequential degradation of the components, respectively. Though the details of process dynamics depend on the external heat flux and the wood category, the differences tend to disappear for heat fluxes above 40kW/m2, when a shift from chemical reaction to internal heat transfer control occurs. The wood components, independently of their nature, attain maximum degradation rates in a relatively narrow temperature range, i. e. 500-650K.Consequently, global parameters, such as pyrolysis temperature, average devolatilization rate, product yields and gas composition are affected mainly by the relative amounts rather than by the different reactivity (hardwood or softwood category) of the components. In general, despite of the same variety (hardwoods), the differences in product yields between beech and chestnut are very large. On the other hand, the differences are significantly lower between beech and pine (softwood), given the comparable contributions of the main components. Finally, softwoods with nearly equal chemical composition (redwood and Douglas fir) also present nearly equal product distribution. Thus, in terms of product distribution, the percentage of the different components in the chemical composition of wood is more important than the wood variety (or the different reactivity of the main components). In particular, the share between solid and liquid products from the degradation process is highly dependent on the holocellulose contribution in the wood chemical composition, with successively higher solid residues as lignin and/or extractive contents increase. Differences between pyrolysis products from wood and biomass are high for liquid and char yields, whereas the gas yields are comparable and so are the apparent kinetics of single species formation. Finally, this study provides an extensive set of data on thick wood pyrolysis which can be better interpreted and generalized by the use of mathematical models taking into account the effects of transport phenomena and chemical reactions. Models including such features are already available in the literature (for instance, see References 23,24) and have proven to give quantitative predictions of temperature dynamics, but product yield predictions are still unacceptable, mainly because of unreliable kinetic constants. Therefore, this issue deserves further investigation before extensive computer simulation and/or development of more advanced physical models of thick wood pyrolysis are proposed.
1156
RFERENCES 1. Roberts A. F., Clough G., Ninth Symp. (Int.) on Combustion, The Combustion Institute, Pittsburgh, 1963, pp. 158-66. 2. Tinney, E. R., Tenth Symposium (Int.) on Combustion, The Combustion Institute, Pittsburgh, 1 965, pp.925-30. 3. Kanury, M. A., Combustion and Flame, 18,75-83 (1972). 4. Lee C. K., Chaiken R. F., Singer J. M., Sixteenth Symposium (Int.) on Combustion, The Combustion Institute, Pittsburgh, 1976, 1459-70. 5. Pyle D. L., Zaror C. A., Chemical Engineering Science 19, 147-158 (1984). 6. Ohlemiller T. J., Kashiwagi T., Werner K., Combustion and Flame, 69, 155-170 (1 987). 7. Chan W. R., Kelbon. M., Krieger-Brockett B., Ind. Eng. Chem. Res., 27:2261-75 (1988). 8. Fredlund B., Fire Safety Journal, 20,39-69 (1993). 9. Bilbao R., Millera A., Murillo M. B., Ind. Eng. Chem. Res., 32, 181 1-17 (1993). 10. Gronli M. G.,A theoretical and experimental study of the thermal degradation of biomass, PhD. Thesis, NTNU, Trondheim, Norway, 1996. 1 1 . Di Blasi C., Gonzalez Hernandez E., Santoro A., Ind. Eng. Chem. Res., 39, 87382, (2000). 12. Di Blasi C., Chemical Engineering Science, 55,2931-44, (2000). 13. Di Blasi C., Modelling the fast pyrolysis of cellulosic particles in fluidized bed reactors, Chemical Engineering Science, in press, 2000. 14. Di Blasi C., Signorelli G.,Portoricco G.,Ind. Eng. Chem. Res., 38: 2571-81, 1999. 15. Di Blasi C., Branca C., Santoro A., Gonzalez Hernandez E., Pyrolytic behaviour and products of some wood varieties, Combustion and Flame, in press 2000. 16. Di Blasi C., Branca C., Santoro A., Perez Bermudez R. A., Weight loss dynamics of wood chips under fast radiative heating, J. of Analytical and Applied Pyrolysis, in press 2000. 17. Antal M. J., Varhegyi G.,Ind. Eng. Chem. Res. 34,703-17 (1995). 18. Shafizadeh F., in Fundamentals of Biomass Thermochemical Conversion, (R. P. Overend, T. A. Milne, L. K. Mudge, Eds.), Elsevier, London,, pp. 183-218 (1985). 19. Di Blasi C,, Progress in Energy and Combustion Science, 19,7 1 104 (1993). 20. Scott D. S., Piskorz J., Radlein D., Ind. Eng. Chem. Res. 24,581-88 (1985). 21. Roy C., Pakdel H., Brouillard D., Journal of Applied Polymer Science 41, 337-48 (1990). 22. Ahmed A., Pakdel H., Roy C., Kaliaguine S., Journal of Analytical and Applied Pyrolysis 14,281-94 (1989). 23. Di Blasi C., Chemical Engineering Science, 5 1 , 1 121 1 132 (1 996). 24. Di Blasi C., Int. J. of Heat and Mass Transfer, 41,4139-4150 (1998).
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-
1157
The Mathematical Modeling of Biomass Pyrolysis in a Fixed Bed and Experimental Verification G. Chen', J. Andnes' and D. Y. C . Leung2 Section of Thermal Power Engineering, Department of Mechanical
'
Engineering and Marine Technology, Delft University of Technology, Mekelweg 2, 2628 CD Dew, the Netherlands Department of Mechanical Engineering, The University of Hong Kong, Pok$ulam Road, Hong Kong, China
ABSTRACT: A modelling study of biomass pyrolysis in a fmed bed has been investigated using a novel kinetic scheme, which comprises all the basic Phenomena occurring at biomass pyrolysis. The model considers transient behaviour of biomass particle's physical parameters which are a h c t i o n of the conversion rate. Mass transfer covering the term of convective flow, heat transfer through conduction, convection and radiation, as well as the change in the structure occurring in the biomass particle are accounted for. The prediction based on this mathematical model shows good consistence when compared with those experimental data on conversion, gas and tar yield, gas and tar evolution, and temperature over a wide range of operating conditions executed in the fured bed. The model can be also applicable to other cases like biomass or coal gasification.
INTRODUCTION Biomass is a renewable resource of plenty of storage covering all the world land surface particularly in rural districts, its conversion to synthetic fuels and chemical products appears attractive, especially in developing countries where the imported fuel and chemical products are much unbearable in price. In addition, biomass can be utilized together with coal in varied proportion with an additional advantage of synergistic effects with regard to char reactivity, tar formation and emission of harmful components*. Currently, among all kinds of conversion routes, thermochemical conversion, especially pyrolysis is being paid more attention in consideration of economy combined with technology. The pyrolysis process can be camed out very effectively in a fixed bed in conventional way with additional advantages of relatively easy design, operation and optimization. However, determining a lunetic approach which could describe well pyrolysis mechanism involving in this case is much difficult due to the varying composition of biomass 1158
materials used and the complexity of the chemical reactions involved in the pyrolysis. Furthermore, the mathematic modelling coupling knowledge of kinetics with heat and mass transfer description appears much more difficult, but h s task is really mandatory to be undertaken in that case when precise prediction by model are expected. Since the beginning several studies have been carried out. Bamford et al. proposed a model by the equation of heat transfer coupled with the equation for heat changes due to chemical reactions2. Afterwards, this model was adopted and adjusted by some researchers in order to incorporate the effects of internal convection and variable transport proper tie^^,^. Using dimensionless groups to Bamford’s model, Pyle and Zaror defined the relative importance of internal and external heat transfer and of the intrinsic pyrolysis kinetics5. A general pyrolysis model was specified by Kothari and Antal with a coupled set of two ordinary and three partial differential equations, as well as approximate boundary conditions6. This model concerned three levels of problems covering chemical reaction, external heat transfer, internal heat transfer and internal mass transfer. Based on the assumption of pseudo-steady-state, a transport model applicable for radiative heat transfer was presented by Curtis and Miller7. This model concerns only convection flow and reaction terms but neglects diffusion and pressure-driven flow. Recently, Hastaoglu and Hassam applied a general gas-solid reaction to flash pyrolysis of wood, where convective mass transfer and diffusion through the porous medium were considered but pyrolysis mechanism was depicted using simple two-parallel primary reactions followed by a secondary cracking reaction8. More recently, Batra and Rao introduced into the model a heat transfer equation combined with the pyrolysis reaction rate for biomass pyrolyzed in an annular finned pyrolyser, still using simple two order decomposition model9. The following basic steps can be used to describe in mechanism biomass pyrolysis process: step 1, the degradation of the virgin biomass materials into primary products (tar, gas and semi-char); step 2, the decomposition of the primary tar into secondary products (tar, gas and char), and step 3, the accompanied reaction between primary gas and semi-char. Unfortunately, almost those above-mentioned mathematical models on the one hand did not reflect step 2 or step 3, even if step 2 and 3 together, on the other hand majority of them did not incorporate convective terms of heat and mass transfer by volatile products inside reactor. Although under certain experimental condition, step 2 or step 3 may be ignored due to its non-conspicuous influence to the final pyrolysis product distribution, however, a good mathematical model for biomass pyrolysis should be versatile applicable to other pyrolysis conditions, and thus it should be involved the abovementioned three steps of process, of course heat and mass transfer equations should be included also. This paper presents this kind of mathematical model. Although the model is constructed based on sawdust pyrolysis, it is quite straightforward to apply the same approach to other cases such as straw and municipal solid waste pyrolysis even if to biomass or coal gasification or metal ore reduction.
1159
MODELLING
KINETIC MODELLING The basic principle of the overall kinetic pyrolysis model (OKP model) proposed by us is shown in Fig. I. Under our experimental conditions aimed at the maximum production of medium calorific gas, step 2 and step 3 exert significant influences to final products yielded, and thus attention should be given to them. Our kinetic scheme obviously differs ftom those mechanisms used previously5, The OKP model reveals the significant secondary cracking of the primary pyrolysis tar product and in particular interactions between semi-char and gaseous phase which always exist irrespective of reaction temperature, residence time of the volatile phase. Those important phenomena of adsorption, deposition, char-gas reaction, cracking and repolymerization reactions are all incorporated in this model. Of course, the OKP model can be simplified to those existed in particular case where a certain secondary reaction occurs only to a minimum degree and thus may be neglected.
Biomass -b
Gas
Tar Fig. 1. Kinetic scheme proposed by us for biomass pyrolysis.
Kinetic scheme is typically described by the following partial differential equations, which are in closure physically.
+
-awd - -(kl k2 + k3)Wd at
ac
-L= k,W;' at
- k4C,
1160
aT
--" - k,W:' - (k, + k, + k, )Ta, at
Here k j = Aj exp(-
'/
LT ) ; the initial conditions is as follows:
t=o:
=&, C' = C2' = c,, = T,,= 0
where,
w d
is degradable portion of biomass weight at time t; Aj is pre-exponential
factor at reaction j;
c, and c21,c22 are primary and secondary char weight,
E j is apparent activation energy at reaction j; T,,is primary tar weight; k, - k, are kinetic rate constants; T is temperature and L is gas constant.
respectively;
MA TERCQL CONSERVATION
For tar and total volatile phases, the following two equations are applicable.
1
where, r is radial distance of reaction front; E is void fraction; 8 is deposition coefficient; U is radial diffusion velocity; pT , pd , Pv , PTaland PGIare concentrations of total tar, degradable biomass, total volatile phase, primary tar and primary gas, respectively. This material balance equations contain terms only for convective flow and reaction, neglecting both diffusion and pressure-dnven flow. This is due to the fact that early simulation showed diffusion and pressure-driven flow in the material balance several orders of magnitude smaller than the convection and reaction In our material conservation equations, a term with respect to the transient term is introduced. A pesudo-steady-state assumption is not made for tar and total volatile phase flow because residence time of volatile phase in the reactor is comparable in magnitude to the time necessary for temperature profile to change significantly. ENERGY CONSERVATION
The energy balance in the fixed bed includes convective, conductive and radiative heat transfer and heats of reactions terms. The equation is described below:
1161
Where
K, is the effective thermal conductivity, Cv,p,Cs,p are specific heat of
total volatile phase and solid phase, respectively; p, is the concentration of solid phase and Ah is reaction heat. The first term of the left-hand side in the equation is the generation term related to internal enthalpy of solid and volatile phases. The second term is the convective flow of the volatile components. The first term of the right-hand side is the effective conductive flow and the second term is the reaction heat with regard to chemical reactions and phase changes. INITIAL AND BOUNDARY CONDITIONS In order to solve mass and energy balance equations, initial and boundary conditions should be given. Initially, the biomass material is in a quiescent environment at ambient conditions and thus is specified as uniform temperature and solid compositions. For P O , the spatial conditions at the centerline ( ~ 0 are ) specified by symmetry, the two sides (FR) are exposed to radiation by a constant temperature of heater. At these two sides, radiant heat transfer from the surface takes places. Further conditions of constant ambient pressure and zero gradient of tar and gas mass at centerline are also used.
T r = R : K,-=eea(T, r
4
-T4);U = O
Where: e is emissivity of pyrolyzing biomass particle; coefficient; T, is pre-set temperature.
CT is
Stefan-Boltzmann
SOLUTIONS OF THE MODELLING EQUATIONS Assumptions (1)
The radiative resistance and conductive resistance outside inner surface of the reactor are advertently omitted and thus the pre-fixed temperature of the heater can be assumed as the initial value of inner wall of the reactor. (2) Biomass particle is taken as non-shrinkable during pyrolysis process and its influence may be accounted for by the change in void fraction of the bed. (3) The temperature is symmetric along the centerline and thus only radial temperature profile exists as the ratio of height of reactor to its diameter is larger
than 5 .
1162
Parameters and physical properties Table I presents the values of physical properties and simulation parameters used, which were derived fiom data given in our experimental results and literature. These values are regarded as overall values with regard to not only the corresponding values of the biomass material itself but also of its derivatives produced. It is considered that the local values of these properties in particular heat conduction coefficient K , and specific heat of biomass.
cs,p are hnctions of the pyrolysis temperature and the conversion rate K , , cs,p and B (permeability) were estimated as the s u m of the value
corresponding to the virgin biomass material and of the char at a certain temperature. Each of the values of three parameters contributes to the s u m in a way proportional to the conversion rate. The expression is as follows:
Table 1. Parameters and properties used. Parameters and properties
Source Chan
Ah, = Ah2 = Ah3 = -41 8kY I kg Ah, =Ah, = 4 2 M l k g Ah, = -492kYl kg Ah, = 4 2 k J l k g 6 0 = 166.5kgIm’
Koufopanous2 estimated estimated Chen14
C, = 2 . 3 M l k g . K
Kung4
C, = 1.1WI k g - K
Kansa
d =4x10%~ Kb =O.llWlm*K
chanl5 Chen14
,~=3xlO-~kglrn.s e = 0.95 E, = 0.70
Kansa
B, = 1.O x 1O-’ Darcys B, = 0.32Darcys
Di Blasil
pyle5 Chen14
estimated
Where X is to represent K,, , Cs,pand B; 77 =
w%o
(w~ -
, is the conversion
rate of biomass material in the bed. Advertently, the contribution of conduction of
1163
volatile phase is not accounted for by K, due to its value more than one order in magnitude smaller than the values of virgin biomass or of char. Dummy thermal conductivity of K,, (also named as equivalent thermal conductivity) was calculated as follows:
ad K,, = 13.5T’ -e E
The effective thermal conductivity K,, defined by the following equation, which accounts for the combined contribution of conductive and radiative heat transfer can be thus applicable.
The specific heat of total volatile phase, C v , p ,is taken as the molar-weighted average of tar-phase and gas-phase, and the tar as well as gas phases specific heats are also taken as the molar-weighted average of all its components, respectively. The values of the specific heats of tar and gas phases are define by equations below:
%J
T
= -286.721 + 1549.749-
T O
CG,p=41.934+1.572-
T
T O
Radial velocity U can be derived from the assumption of ideal gas and the void fraction of the bed can be calculated below: E = 1 -(1-
w
PS,,
wo
Ps
Eo)---
Where the density of solid in the bed, p expression with a modified coefficient,
,can be deduced ffom general pyrolysis rate
%--E)-
Mathematical manipulation Considering the following work is to be carried out in the fixed bed, where reaction temperature is not constant whilst the constant heating rate (a)may be assumed within accuracy, therefore here only a certain heating rate is concerned. In this case, lunetic equations should be re-constructed by expression of T = cDt To( To is
-
+
initial temperature). The best values of k, k, were obtained, using the least square analysis where the s u m of square of deviation between theoretically calculated
1164
residual mass values and experimentally measured results by thermogravimetry. Detailed method for obtaining kinetic parameters is available elsewhere14 . The method of zonation was applied to the energy and material conservation equations. Based on centered finite difference approximations, this method can transform three partial differential equations in radial distance and time to ordinary differential equations in time only. Following this, the ordinary differential equations were solved by using Crank-Nicholson algorithm. On the basis of this, the volumetric fluxes of those tar-phase and total volatile phase components were integrated with time by using improved Euler method to evaluate overall pyrolysis product yields, and afterwards the gas yeld can be deduced.
EXPERIMENTS The verified experiments were carried out in the fixed bed, which was set up by us in particular for the purpose of investigating pyrolysis behaviors of biomass materials. The detailed description of that apparatus is shown elsewhere14. The following parameters and operating conditions were used: for fixed bed, 80mm i. d., 450mm in height, bed temperature 500-90Ooc; for pine sawdust particle: 0-lmm in diameter, apparent density 166.5 kg/m3, initial porosity 0.70, energy density 3007.7 kJ/ m3, apparent char density 95 kg/m3; initial sawdust weight 0.375kg. In addition, the following estimated properties were also used: initial sawdust permeability 1.OE-5 darcys, char permeability 0.2 darcys; molar molecular weight of tar, gas, initial sawdust and char 145,31, 145 and 12, respectively. Those kinetic experiments were canied out using a thermogravimetry. The biomass sample fnst was kept in a differential compartment and then introduced into the reaction zone by an inert gas flow of nitrogen. The temperature and the sample weight are continuously and automatically monitored. The experiments were based on constant heating rate of 50 " C / min while not on other heating rates in order to keep close simulation of those following experiments executed in fixed bed.
RESULTS AND DISCUSSIONS The relation of weight-loss and temperature variation at pine sawdust pyrolysis by thermogravimetry is shown in Fig. 2, where the continuous line is predicted result by model and the symbols are the experimental values. As constant heating rate is considered, the correlation of weight loss and temperature variation represents weight-loss evolution history. As can be seen from this figure, the kinetic model predxts well the weight-loss and temperature variation. Fig. 3 shows temperature histories at the wall surface and center of the reactor, respectively. In general, the simulated results agree well with those experimentally monitored data. The divergence appearing in the figure may be contributed to the negligible moisture evaporation terms in the model. The evaporation of moisture contained in the virgin biomass particles is an endothermic process, and thus monitored temperatures are lower than those predicted at the initial stage of pyrolysis. Another possible explanation to that divergence is due to term Ah4incorporated in
1165
energy balance equation. Up to now, there is still no accurate description to the reaction scheme related to k, . Fig. 4. shows the evolution history of volumetric flow of gas-phase yielded. The experimental dots if regressed by a curve whether in shape or magnitude is well similar to that theoretically predicted. The maximum gas-phase flow is achieved at time of 240s or so. Theoretically, this point of time is exactly corresponding to the beginning of fast decomposition of cellulose, main component of biomass. Also the phenomena represented by this figure reflects the same trend by Fig. 3, where the predicted temperature is normally a little higher than of experimentally measured. In this figure, it is easily to be observed that the predicted volumetric flow curve is also put forward a little. The area beneath the predicted volumetric flow curve is corresponding to the simulated non-condensable gas phase yield, 100
r-0
70
exp.
-model
0' 0
I
200
400
600
800
1000
Temperature ("C)
Fig. 2. Weight-loss versus temperature variations for pine sawdust pyrolysis at constant heating rate of 50 "c/ min . Dots,experimentaldata; solid line, model predictions.
1166
I
loo0
800
G
5 H
600
8
E
400
f -model
center
200
0 0
200
400
600
800
1000
1200
lime (s)
Fig. 3. Temperature profile at the center-line and inner wall surface of the reactor versus time during pyrolysis at temperature of 800 "c, Symbols, experimental data; line, model predictions.
500 A
J!! E400
B .- 300 U L
Q)
-5 200
8
100
0
0
0
200
400
600
800
1000
1200
Time (s) Fig. 4. The evolution hstory of volumetric flow of non-condensable gas-phase yielded during pyrolysis at temperature of 900 "C .Dots,experimental data; solid line, model prediction.
1167
The final product yields at different temperatures are presented in Fig. 5. It can be seen an increase in temperature results in increasing value of gas yield defined by kg of per kg of biomass feed, while reducing value of tar phase yield. In the latter case, the decrease observed could be related to secondary reactions (tar cracking reactions) where tar components of heavy molecular mass decompose into that of light molecular mass and accompanied gases. Also, when the reaction temperature increases further from 700 "c , the decrease observed in the value of tar yield is becoming weaker. This may be attributed to the fact that tar cracking is easy to process at temperature of 500 "cabove and furthermore approximately accesses final point as temperature reaches 700°C. It can be also observed that the increase of gas yield is sufficient to compensate for loss of tar yield. Char-gas interactions occuning at temperature of 700 "c higher probably contribute to this observation. When compared with the data obtained experimentally, the values of gas and tar yields estimated by the model are shown to be ideal, though an over-estimation of gas yield and an under-estimation of tar yield are presented. The over-estimation may be due to the fact that char-gas reactions of char and H20,char and C02 incorporated in the model are not significantly influential to products yields under our experimental conditions where heating-time is comparable to devolatilization-time of biomass particles. The under-estimation of tar yield, of course, is accompanied by an assumption of secondary reactions occurring in all yielded tars.
25
-L z 20
ii
500
+
-
600
800
700
900
1000
Temperature (OC)
exp. gas -model tar
exp. tar model gas
-
Fig. 5. The final gas and tar yields at different temperatures. Symbols, experimental data; solid line, predicted gas yield; dash line, predicted tar yield.
Fig. 6 presents the evolution of tar and gas yields measured by weight percent. Also the figure shows the corresponding simulated results. The profile is sharp at initial stage of the process (Wminutes), followed by a relatively sharp stage (4-7minutes) and then by a flat period (7minutes-end). The gas and tar yields at the time period 0-7 minutes account for 90% and 95% of the total yields, respectively. Such considerations are important when the economics of the process are to be taken into account. The simulated results are well corresponding to the measured data, considering the complexity of sawdust pyrolysis process. From this figure, conversion
1168
rate 7 can be obtained, however, h s part of work will be discussed in the near future publication. 50
r
*g25 2
CI
20 g 15
-0
-mn 10 C
L ' 5
0 0
2
4
8
6
10
12
14
16
Time (minute)
+
-
exp. tar -model tar
0
exp.gas model gas
-
"c
Fig. 6. Evolution of gas and tar velds during pyrolysis at temperature of 900 . Symbols, experimental data; solid line, predicted gas yield; dash line, predicted tar yield. CONCLUSIONS It has proven that our kinetic pyrolysis scheme covering three basic steps can be well applied to the analysis of sawdust pyrolysis in a fxed bed. Further to this, the selfconstructed mathematical model incorporating heat and mass transfer equations combined with our OKP model can be used to successfully predict the behaviour of biomass particles pyrolyzing in the fixed bed and the comparison between experimentally obtained values and predicted results by model shows good agreement. The modelling also considers the structural changes of biomass particles inside the bed through the equation of void fraction. In addition, the model incorporates convective flow terms in material conservation and radiative heat transfer in energy conservation. With the aid of the model, it is completely possible to optimize various parameters for any required yield. REFERENCES 1.
2.
Andries J. & Hoppesteyn P. D. J. (1997) Pressurized combustion of biomassderived low calorific value, fuel gas. In: Biomass, Gasification and Pyrolysis, (Ed. by M. Kaltschrmtt & A. V. Bridgwater), pp. 1282-291. Bamford C. H. & Crank J. et al. (1946) The combustion of wood. Proceed. Cambr. Phil SOC.,42, 166-82.
1169
3.
4.
5. 6. 7. 8.
Capart R. & Fagbemi L. et a1 (1986) Wood pyrolysis: a model including thermal effects on the reaction, energy from biomass. (Ed. By W. Palz), pp. 842-846. Elsevier Applied Science Publication. Kung K. C. (1972) A mathematical model of wood pyrolysis. Combustion and Flame, 18, 185-95. Pyle D. L. & Zaror C. A. (1984) Heat transfer and kinetics in the lower temperature pyrolysis of solids. Chem. Eng. Sci., 39, 147-58. Kothari V. & Antal M. J. (1985) Numerical studies of the flash pyrolysis of cellulose. Fuel, 64, 1487-494. Curtis L. J. & Miller D. J. (1988) Transport model with radiative heat transfer for rapid cellulose pyrolysis. Ind. Eng. Chem. Res., 27, 1775-783. Hastaoglu M. A. & Hassam M. S. (1995) Application of a general gas-solid reaction model to flash pyrolysis of wood in a circulating fluidized bed. Fuel, 74, 697-703.
9. 10.
11.
12. 13.
14. 15.
Batra D. & Rao T. R. (2000) Analysis of an annular finned pyrolyzer-11. Energy Convers. Manage, 41,573-83. Koufopanos C. A. & Papayanakos N. (1991) Modelling of the pyrolysis of biomass particles: studies on kinetics, thermal and heat transfer effects. Can. J. Chem. Eng., 69, 907-15. Ward S. M. & Braslaw J. (1985) Experimental weight loss kinetics of wood pyrolysis under vacuum. Combustion and Flame, 61,261-69. Di Blasi C. (1996) Kinetic and heat transfer control in the slow and flash pyrolysis of solids. Ind. Eng. ChemRes., 35, 27-46. Di Blasi C. (1996) Heat, momentum and mass transport through a shrinkmg biomass particle exposed to thermal radiation. Chem. Eng. Sci., 51, 1121-132. Chen G. (1998) Study on mechanism for biomass pyrolysis and its experimental verification. DPhil thesis, Zhejiang University. Chan W. R. & Krieger B. B. (1985) Modelling and experimental verification of physical and chemical process during pyrolysis of a large biomass particle. Fuel,
64, 1505-513. 16. Kansa E. J. & Perlee H. E. (1977) Mathematical model of wood pyrolysis including internal forced convection. Combustion and Flame, 29,3 11-24. 17. Lee C. K. & Chaiken R. F. (1976) Charring pyrolysis of wood in fire by laser
simulation. Sixteenth Symposium on Combustion, the Combustion Institute, pp. 1459-470.
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Origin and Nature of Paramagnetic Moieties in Pyrolysis Oils T. Dizhbite, G. Dobele, N. Mironova, G. Telysheva, Latvian State Institute of Wood Chemistry, 27 Dzerbenes St. , L V-1006, Riga, Latvia D. Meier, 0.Faix Federal Research Centre for Forestry and Forest Products, Institute for Wood Chemisty and Chemical Technology of Wood,Lewchnerstr. 91, D-21031 Hamburg, Germany
ABSTRACT: Paramagnetic properties of pyrolysis bio-oils produced on different fast pyrolysis pilot plants have been studied using ESR spectroscopy. A narrow (0.3-0.4 mT) ESR singlet detected in the oil samples was assigned to stable paramagnetic centers of expanded n-conjugation systems, which involve 6 - 12 benzene rings. The content of paramagnetic species varied significantly (3 - 4000 ppm) for oil samples from different plants. It has been shown that paramagnetic domains in oils are related to both char microparticles admixtures and polyarylene structures present in the tar (pitch) fraction of oil, rich in “pyrolytic lignin”. Based on ESR and polarization microscopy data it is assumed that formation of anisotropy associates of n-conjugation paramagnetic domains, forming an ordered mesophase, took place. The possibilities of ESR for monitoring phase separation of pyrolysis liquids during storage is discussed. INTRODUCTION Proposed application of bio-oils as fuel for gas turbines or diesel engine generators requires stability of their properties in conformity with the standard specifications. However, it has been recognized for a long time that bio-oils properties tend to change during storage resulting in an increase in molecular weight and viscosity, phase separation and fine char sedimentation [l, 21. The aging of bio-oils depends on a combination of physico-chemical processes and chemical reactions including association and agglomeration, polymerization and condensation as well as redox processes. Bio-oils can be considered a multicomponent system, which often contain impurities such as char fines and ash elements which have paramagnetic properties [ 3 ] . However, there are no published data about paramagnetic properties of bio-oils obtained by fast pyrolysis of wood, although many research groups have reported paramagnetic properties of charcoal and petroleum pitch. It is well known that the various paramagnetic impurities can play a key role in promoting ageing as they act as catalysts for hydrolytic andor oxidative reactions. Therefore, study of paramagnetic properties, their nature and origin seems to be necessary to forecast the stability of biooils, and to assess of the role of the different paramagnetic centers in the ageing process of bio-oils. Hence, this study aimed at monitoring pyrolysis liquids by ESR and analyzing the results.
1171
MATERIALS AND METHODS
Bio-oils were produced in 1995-1998 from various wood species on fast pyrolysis pilot plants, which differed by their construction features: NREL (ablative pyrolysis with hot gas filter, poplar wood, 1996), Aston FB (fluidized bed, poplar wood, 1995), IWC, TP 29 (fluidized bed, bamboo, beginning of 1998) BTG (rotating cone, pine wood, 1995), VTT (circulating fluidized bed, pine wood, 1997), IWC, E-filter (electrostatic precipitator, beech wood, end of 1998) and IWC Kl-FlO from different coolers (beech wood, end of 1998). ESR spectra at X-band were measured with a RE-1006 spectrometer operating at 9.6 GHz (empty cavity at ambient temperature) and 50-kHz magnetic field modulation. Spectra were recorded at a microwave power of 1 mW and a modulation amplitude 0.02 mT to avoid line shape distortions, which could arise from experimental conditions such as microwave saturation and ovennodulation. Scan range was 150420 mT. Calibration of g-values is based on diphenylpicrylhydracyl (DPPH, g=2.0036) and Cr” (g=1.9796) standards. The amounts of the paramagnetic species (PMS) were calculated by double integration of the resonance line areas. Oil samples for ESR measurements were introduced in the ESR sample tube for ESR measurements in a layer of 5 mm thickness. Typically ESR measurements were carried out on samples in open tubes. To avoid oxygen influence air from the tube was carefully removed by repeated thawing and freezing under high vacuum (ca. Tom) and then sealed off. All measurements were carried out at ambient temperature. An automatic microscopic image analyzer “Morphoquant” (Karl Zeiss) with a high resolution optical microscope was applied to measure the sizes of fibres. The accuracy of measuring was 0.1 pm. About 1000 microparticles were analyzed in each oil sample. Ageing of oil samples was carried out under the following conditions: 80°C for 24 hours [2]. Oil viscosity was measured using a capillary viscometer prior and after ageing. RESULTS AND DISCUSSION ESR measurements have shown that all bio-oils studied contain measurable quantities of paramagnetic species, excluding the sample from NREL (Table 1). For the latter only trace amounts of PMS, < 0.1 ppm, was observed. The ESR spectra of the oils exhibit a symmetric singlet signal at g-value about 2.0025 without hyperfine structure. Line width values are within a nmow range of 0.30.4 mT. On the contrary, PMS content varies considearbly from sample to sample, being 3 ppm and 4000 ppm for oils from VTT and BTG plants, respectively (Table 1). The ESR signals from ashes were not detected,
1172
Table I Parameters of ESR spectra of bio-oils. Oil sample
ESR line
g-value**
width* (mT)
Concentration of paramagnetic species: spinslg ppm < 1.1015
NREL poplar 3.4 2.0027 Aston FB poplar TP 29 2.9 2.0024 (5.1ko.5) bamboos Twente V 3.7 2.0028 (1.4f0.2) * l o ” BTG pine 4.1 2.0028 (1.5f0.4) VTT ITP pine E-Filter LDR (1.9k0.1) *loL9 3.1 2.0024 K1-F 10-LDR (6.6k0.5) *lo” 3.4 2.0025 * The accuracy of line width measuring was 50.1 **The accuracy of g-value measuring was kO.0001
100
30 3
4000 130
In the case of relatively fresh bio-oil samples, IWC E-Filter LDR and K1-F10-LDR, produced a couple of weeks before ESR monitoring started, a prominent tendency of a decrease in PMS amount within the first three months of storage was observed. After that the PMS content was stable (Table 2). Table 2 Dynamics of PMS content in the bio-oils during storage in open ESR measuring tubes at ambient temperature. Content of paramagnet1 c species, f10-17 spins/g, after storage for: Initial sample 7 days 14 days 21 days 30 days 45 days 60 days 80 days 200 days 300 days
Oil sample BTG VTTpine pine
Aston FB poplar
IWC bamboos
6.5f0.5
5.1f0.5
1.4k0.2
6.6f0.4 6.2f0.6 6.4f0.5 6.5k0.5 6.5k0.5 6.9-t-0.6 6.3k0.5 6.1f0.5 6.2k0.3
4.6f0.5 4.8f0.5 4.5f0.5 5.2k0.5 4.8f0.3 5.3f0.6 4.6f0.5 5.OkO.5 5.2+0.2
1.21t0.1 1.4k0.2 1.2f0.2 1.6k0.3 1.4k0.2 1.4k0.1 1.4k0.3 1.5f0.2 1.2k0.3
IWC beech E-Filter
IWC beech K 1 -F10
0.15f0.04
190+10
6.6k0.5
0.19k0.05 0.18f0.04 0.17f0.04 0.19f0.05 0.1650.03 0.15+0.03 016250.04 0.15f0.04 0.15k0.02
180+15 160k10 140+10 13057 130f10 130+10 120k10 120+15 120k10
6.4f0.5 5.7k0.5 5.3k0.2 4.8k0.5 4.8k0.3 4.6f0.5 4.4k0.2 4.4f0.1 4.4f0.3
Paramagnetic centers of all other oil samples exhibit a high stability: for. most samples the PMS content and other ESR spectral parameters practically did not change during storage neither under vacuum nor in the presence of air oxygen at ambient temperature (Table 2). Analysis of the detected parameters reveals that ESR spectra of the bio-oils are characterized by paramagnetic centers, whose unpaired electron is delocalized in the n-
1173
conjugation system (n-polyconjugation paramagnetic domains). The presence of expanded n-conjugation systems, bearing paramagnetic species, is a characteristic feature for carbonized materials, i.e. chars, coals and coal pitches and conducting synthetic polymers, “synthetic metals”. Up to now the origin of paramagnetic centers in polyconjugated systems is under discussion. Nowadays, the most accepted model is the concept of solitons for polyacetylenes and polarons for polyaromatic polymers [4, 51.1. According to the ESR theory polyconjugated systems are characterized by the degree of spin density delocalization, which can be estimated on the basis of the ESR line width [ 6 ] . The parameter n corresponding to the number of equivalent protons interacting with an unpaired electron, can be estimated from the following expression:
n 2 1 + (AH~AH,~)’, where AH,=2.25 mT is a splitting corresponding to interaction of a p-electron with a proton on a CH-fragment and AHeff is the experimentally observed ESR spectrum line width. For the oils studied n is in the range 25-30. According to the well-known characteristics of thermal biomass conversion, the backbone of n-conjugation systems of bio-oil components consists of aromatic rings. So the delocalisation region of unpaired electron can derive kom 6 to 12 benzene rings, which corresponds to the paramagnetic domains size of 2 30-50 A. The above-mentioned decrease in the PMS content for the fresh samples from the BTG plant could be due to the decay of unstable radicals formed by breaking of chemical bonds during pyrolysis. However, a persistence of the ESR spectrum line width and shape suggests that the paring of a part of PMS through the n-conjugation system or by intermolecular interaction may also take place [7]. Based on the concepts proposed in literature of structural peculiarities of biomass derived pyrolysis oils and pitches [8,9], which have about 50% of aromatic and olefinic carbons, these matenah can contain high-developed n-conjugation systems as structural constituents. At the same time, ESR characteristic observed for the samples studied are close to the same for charcoals obtained from pyrolysis at temperature about 500°C [lo]. It remains uncertain, whether the paramagnetic centres in the oils are associated with n-conjugation systems coming from high molecular weight lignin fragments, or whether they are associated with fine char impurities. Taking into account the absence of detectable amounts of paramagnetic centers in the NREL oil (the plant is equipped with a hot vapour filter), it seems that the ESR spectra of all other oils are due mainly to the presence of char microparticles. However, careful study of the spectra prior and after oil samples microfiltration revealed the presence of one or more polyconjugated species depending on pyrolysis plant construction features. It has been shown that microfiltration (pores size 2 pm) of oil samples decreased sufficiently both the quantity of microparticles and their size. In the case of the pyrolysis oil obtained from IWC plant, an ESR signal was not detected after filtration, showing that fine char particles caused the detection of paramagnetic moieties in this sample. However, the PMS content in oil samples from Aston FB, BTG and VTT plants remained constant after filtration.(Table 3).
1174
Table 3 Changes in parameters of ESR spectra of bio-oils after microfiltration and ageing.
Oil sample
NREL poplar Aston FB poplar IWC bamboos BTG pine VTT pine IWCE-Filter beech IWC K1-F10 beech
*
**
Concentration of paramagnetic species, spinslg Prior filtration After filtration After ageing of filtrated samples 110’5 11015 110i5 (7.7+0.8)* 017 (6.5+O.5)*1Ot7 (7.2+0.4)*lOI7 (1.6+0.3)* lo1’* (46_+5)* 1017** ( 5.1+0.5)* 10” 0.00 0.00 (1 .4f0.2)*10I7
(1.5+0.2)* 017
(1.6f0.2)*1017
(1.5+0.4)* 0l6
(1.6+0.4)*1016
(1.5+0.2)* 10lb
(1.9f0.1)* 019
(0.8+0. 1)*1019
(1.o+o. 1)* l o i 9
(6.6f0.5)* 017
(6.8+0.5)*1017
(6.9+0.5)*1017 (0.33+0.05)* (290*10)* loi7** The probe was not stored in an ESR measuring tube but was taken from the top of the vessel with the bulk oil sample The probe was taken from the bottom of the vessel with the bulk oil sample
Filtration of the IWC E-Filter sample led to decreasing average effective diameter of microparticles from 4.5 pm to 2.5 pm (Fig. 1). Simultaneously PMS concentration decreased about two times (Table 3). For this sample narrowing of the ESR spectrum line was observed after ageing from 30.32 rnT to 2.5 mT accompanied with a slight increase in the PMS content (Table 3). In accordance with expression (1) narrowing of the ESR line reflects an increase in delocalization region of unpaired electron. The cause of this phenomenon could be chemical and physico-chemical processes. The former is connected with condensatiodpolymerization reactions and the latter with densification of intermolecular packing of coplanar aromatic cycles resulting in increasing electron exchange interaction between them [ 111. An increase in the PMS content after ageing confirms that the main reason of the ESR line narrowing is physico-chemical intermolecular interaction resulting in formation of more dense microstructures, e.g. aggregates consisting of x-conjugation domains. Owing to co-planarity of sr-conjugation regions, such aggregates have more ordered structures than bulk oil and to some extent can be considered as ordered mesophase. In accordance with the assumption of the presence of ordered entities, which rotate a plane of light, polarization was observed by polarization microscopy in the IWC E-Filter oil sample after ageing. Ageing of this sample at higher temperature (105OC) led to anisotropy in its ESR spectrum with an amplitude of anisotropy of 0.9 mT [12]. The average g-value of the anisotropy spectrum is equal to that of isotropic spectrum. Anisotropy observation under conditions of intermolecular exchange of unpaired electrons is possible only in the case of mutual orientation of the main molecular axes of paramagnetic domains located in high ordered mesophase structures [13]. This leads to a preferential orientation of PMS with respect to the magnetic fields.
1175
Prior Filtration 16
-. 14 'u .-8 12 5 lo n
r
0
8
After Filtration ..
... .I.I "
"
".."""IIx_xI,.-" ..... .
....."_ .. . ..
~
..
Fig. / Histograms of the distribution of char particles before and after microfiltration of BTG oil.
Results obtained so far allow to consider paramagnetic species of bio-oils as natural spin probes, which can provide an information about changes in the oil's microstructure. With this respect, testing of bio-oils phase separation during ageing could be an important application of the ESR method. Comparison of the results of viscosity measurements (using a capillary viscometer) before and after ageing (Table 4) showed that for a series of bio-oils viscosity decreased or practically did not change after ageing.
1176
Table 4 Changes in viscosity of bio-oils after microfiltration and ageing. before filtration Oil sample
NREL poplar Aston FB poplar IWC bamboos BTG pine VTT pine IWC E-Filter beech IWC K1-F10 beech
vzOQc
after filtration
after ageing of filtered samples
"SOT
"20°C
"50°C
V2O"C
VSOT
(CSt) 370.53
(CSt) 41.46
(CSt) 339.89
(CSt) 43.07
(CSt) 222.44
(CSt) 29.44
690.61
64.68
726.43
66.43
752.08
63.86
5.72
2.3 1
6.24
2.40
5.72
2.22
1758.14
105.19
1728.42
107.75
133.25
1258.21
89.54
1256.06
88.17
265.73
31.60
2321.7 8 1416.8 4 372.77
30.38
11.11
14.93
5.02
97.93 42.65
Obviously a viscosity index recommended for characterization of bio-oil stability [2], "using a viscometer type available in a participating laboratory", cannot be correctly applied due to tendency of bio-oils, non-Newtonian liquids, to phase separation into thin oil, thick tar and solid admixtures [ 11. However, the ESR method allows to reveal considerable difference between the properties of the samples prior and after ageing. It was shown (Table 3) that the concentrations of paramagnetic centers in samples taken after ageing from the top and the bottom of oil storage vessel differed significantly, e.g. in the cases of oils from IWC and Aston installations, by 10 and 800 times, respectively.
CONCLUSIONS ESR measurements indicated the presence of highly stable paramagnetic species in biooils, which are associated with paramagnetic centers (PMS), whose unpaired electron is delocalised in the sr-conjugation system (n-conjugation paramagnetic domains). The delocalisation region of unpaired electron comprises 6 to 12 benzene rings that corresponds to domains size of 30-50 A. Partially these PMS can be assigned to the presence of fine char impurities in biooils. An increase in the PMS content after ageing is connected with physico-chemical association of coplanar aromatic rings resulting in the formation of more ordered microstructures with a higher density. ESR could be considered as a monitoring procedure of oil's susceptibility to the phase separation during the storage.
1177
REFERENCES 1. Diebold J. P., Milne T. A,, Czernik S., Oasmaa A., Bridgwater A.V., Cuevas A., Gust S., Huffman D., Piskorz J. (1997) Proposed specification for various grades of pyrolysis oils. In: Developments in Thermochemical Biomass Conversion (Ed. by A.V.Bridgwater, D.G.B.Boocock), 1, pp. 433-447. Blackie Academic and Professional. 2. Czernik S. (1999) Workshop on stabilisation /upgrading of bio-oils. Pyrolysis Network. Vacuum Pyrolysis, 8, 3. 3. Scahill J., Diebold J.P., Feik C. Removal of residual char fines from pyrolysis vapors by hot gas filtration ( 1997) Proposed specification for various grades of pyrolysis oils. In: Developments in Thermochemical Biomass Conversion (Ed. A.V.Bridgwater, D.G.B.Boocock) 1, pp. 253-266. Blackie Academic and Professional. 4. Kume K. and Mizoguchi, K. (1993) Magnetic resonance and resistivity studies on polyacetylene. Progr. Theor. Phys. Suppl., 113,9 1-96. 5 . MacDiamid A.G., Epstein A.J. (1991) “Synthetic metals”: a novel role for organic polymers. Makromol. Chem., Macromol. Symp., 51, 11-28. 6. Blumenfeld L., Voevodsky V., Semenov A. (1962) Application of ESR in Chemistry. Novosibirsk. 7. Hayashi H. and Yamamoto T. (1998) New n-conjugated polymers derived from a benzimidazole unit. Preparation, solvatochromism, and oxidation of poly( aryleneethyn1ene)s composed of 2-(3,5-di-tert-butyl-4hydroxypheny1)benzimidazole bearing a hindered phenolic substituent. Macromolecules, 31, 6063-6070. 8. Scott D.S., Legge R.L., Piskorz J., Majerski F., Radlein D. (1997) Fast pyrolysis of biomass for recovery of specialty chemicals. In: Developments in Therinochemicnl Biomass Conversion (Ed. by A.V.Bridgwater, D.G.B.Boocock), 1, pp. 523-535. Blackie Academic and Professional. 9. Pasa V.M.D., Carazza F., Otani C. (1997) Wood tar pitch: analysis and conceptual model of its structure. In: Developments in Thermochemical Biomass Conversion (Ed. by A.V.Bridgwater, D.G.B.Boocock), 1, pp. 448-4446 17. Blackie Academic and Professional. 10. Mrozovski S., Gutsze A. (1977) From char to carbon: a study of the electron spin resonance. Carbon, 15,335-342. 1 1. Kovarski A. L., Lopatin S.I., Kasparov V.V., Tikhonov N. A. and Oleinik E.F. ( 1999) Visokomolekularnye soedineniya, A 40 , 1803- 1808. 12. Barta P. S., Nniziol M., Pron A. and Pacyna A. ( 1993) Low temperature magnetic properties for poly(3-alkylthiophenes) and poly(4,4 ’-dialkyL2,2’-bithiophenes). Synthetic Metals, 55-51,5003-5007. 13. Shklyaev A.A., Miloshenko T.P., Lukovnikov A.F.(!988) ESR spectra and orientational order in semi-coke anisotropic structure. Khimiya Tverdogo Topliva, N1,70-75.
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New Prospects for Biocarbons Michael Jerry Antal, Jr., Xiangfeng Dai, Brent Shimizu, Man S. Tam, Hawaii Natural Energy Institute, University of Hawaii, 2540 Dole St., Holmes Hall 246, Honolulu, HI, 96822, USA; Morten Gramli, SINTEF Energy Research, 7024 Trondheim, NORWAY.
ABSTRACT: We present a concise review of recent advances in technology for producing biocarbons, and new results concerning the electrical conductivity of biocarbons. In light of the fact that biocarbons can be produced in yields which attain the theoretical limit, we recommend that their potential as fuels, reductants, adsorbents, and electrodes be more fully explored. In particular, we recommend research aimed at the development of a biocarbon fuel cell.
INTRODUCTION
In 1956 M. King Hubbert used a mathematical model to predict that oil production in the U.S.A. would peak around 1969, and that world oil production would peak early in the 2Istcentury (I). Oil production in the 48 states peaked according to the model, and there is now widespread agreement that world oil production will peak sometime during the next ten years (2). Most projections of future energy supplies indicate an important role for biomass and other renewables as a substitute for fossil fuels. To fulfill this role, biomass must be converted into electricity or a suitable transportation fuel. Conversion methods include gasification for the production of a low to medium Btu gas or hydrogen, and “liquefaction” methods to produce biocrude “oils” (by fast pyrolysis) or ethanol (by fermentation). Electric power can be produced by combustion of the low Btu gas or the biocrude in boilers, gas turbines, or diesel engines. Likewise, electric power can be produced from hydrogen by electrochemical combustion in a fuel cell. In the U.S.A. ethanol is currently the renewable fuel of choice for transportation, but hydrogen may become competitive if less costly methods can be discovered to produce and store the fuel. These technologies have been the focus of worldwide research and development efforts since the “energy crisis” in 1973. In addition to gaseous and liquid fuels, biomass can be converted to charcoal and carbonized charcoal (a relatively pure carbon). We call these products “biocarbons”. Unlike the gaseous and liquid fuels discussed above, biocarbons have attracted little interest. One reason for the disinterest is the assumption that the properties of biocarbons are similar to coal, and the world has plenty of coal. But in fact, the properties of biocarbons are quite different than those of coal (see below). Another reason for disinterest is the low efficiency of traditional methods for
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converting biomass to charcoal. Simple metal kilns used in the developing world often realize charcoal yields of 20 to 25 wt% on a dry basis (3). Such inefficient processes are the principal cause of the deforestation of many tropical countries (including Thailand, Haiti, and Madagascar). Moreover, because of pollution associated with charcoal production, the charcoal fuel cycle is among the most greenhouse-gas intensive energy sources employed by mankind (4). In the U.S.A. a typical yield of charcoal manufactured from hardwoods in a Missouri kiln operated on a 7 to 12 day cycle is about 28 wt% (5) on a dry basis. This yield is also low: simple stoichiometric arguments show that charcoal yields in excess of 40 wt?hshould be possible, and thermochemical equilibrium calculations indicate that a fixed carbon yield of about 30 wt% should be achieved when equilibrium is reached in a pyrolytic reactor operating at 300 "C.In 1986 our research led us to realize that high yields of charcoal can be obtained when pyrolysis is conducted at elevated pressure wherein the vapors are held in contact with solid products of pyrolysis. A grant from the State of Hawaii enabled us to build a pilot plant, which demonstrated charcoal yields of 42 to 62 wt% with fixed carbon contents of 70% or higher on a 1 hour operating cycle (6). In this paper we report measurements of the fixed carbon yield that result from pyrolysis of a variety of biomass feeds at elevated pressure. We compare these values to values predicted by thermochemical equilibrium calculations. For some of the feeds considered, the yield of fixed carbon delivered by our equipment attains the theoretical limit. Another assumption of the biomass community is that woody materials will be the principal biomass feedstock. Recent studies by Hiromi Yamamoto and his coworkers (7) call this assumption into question. Their work shows that by 2100 most arable land will be dedicated to crop production in order to feed ten billion hungry people. Consequently, there will not be much land available for energy plantations. Instead, the residues of cereal crops will represent the largest available biomass resource. In this work we show that grain residues (e.g. rice and oat hulls) can be converted to biocarbons with yields comparable to those of hardwoods. Moreover, nut shells, some food processing wastes (e.g. garlic wastes), and tropical species (e.g. bamboo and leucaena wood) offer higher yields of biocarbons than traditional hardwood feeds. In light of these findings, we suggest that there are many new and exciting prospects for biocarbons as fuels, reductants for metal ores, and adsorbents for water treatment. EXPERIMENTAL At SINTEF charcoals from birch and oak wood were produced at atmospheric pressure in a 3.7 L stainless steel retort placed in a Nabertherm muffle oven. The muffle oven is equipped with a programmable temperature controller, which provides control of the heating rate, temperature, and reaction time. The retort has a capacity of approximately 1 kg of dry wood per batch. The typical dimensions of the feedstock pieces were 3 cm x 3 cm x 5 cm. During heating and cooling, the retort was purged with N2 to prevent air leakage into the system. At HNEI high-yield biomass charcoals were produced in two reactors with similar configurations and operating procedures. The process development unit (PDU) has an internal volume of 80 L and can produce as much as 10 kg of charcoal per cycle (6). Because the pyrolytic reactions which transform biomass to charcoal are exothermic (8), the pilot plant requires little heat input. The laboratory reactor (9) is a pressure vessel with an internal volume of 7.2 L. At the bottom of this reactor there
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are two internal 4 kW cartridge heaters that are controlled by a temperature controller. Also, a rod heater is installed down the central axis of the reactor. The reactor's outer wall is heated by four band heaters totaling 4.4 kW. The pressure within the reactor is controlled by an on-off valve. Combustible gases released by the valve are burned in a flare. More details concerning the design and fabrication of these two reactors are available in the literature (6, 10).
RESULTS AND DISCUSSION Table 1 displays the results of rapid, and slow pyrolyses of birch and oak wood at 0.1 m a . The charcoal yield for rapid heating of birch is 29.5%, and the fixed carbon yield is 21.4%. An increase in the heating time from about 90 min to about 4 h followed by a 4 h soak at 450 OC increases the fixed carbon yield to 22.8%. The yields of charcoal (3 1.2%) and fixed carbon (24.0%) from oak wood are somewhat higher than those obtained by American industry with this feed in a Missouri kiln.
Table I Charcoal and fixed-carbon yields realized at atmospheric pressure (0. I MPa). Proximate Analyses' Feed
Average Charcoal Yield2 (wt"/.)
Average Fixed C Yield2
VM
Fix-C
ash
(we?)
(we/.)
(we?)
Birch Wood3
26.4
72.5
1.15
29.5
21.4
Birch Wood4
2 1.5
77.2
1.23
28.8
22.3
Birch Wood'
19.9
79.1
1.04
28.8
22.8
Oak Wood4
2 1.7
76.9
1.42
31.2
24.0
(wt%)
'%of dried charcoal. *% of dried feed material. Fixed-C yield = Charcoal yield * (100 - % Volatile Matter - %Car Ash) / (100 -%Feed Ash). 'rapid heating from 20450°C (within 90 rnin), followed by 10 rnin soak at 450°C. 4s10w heating from 20450°C (within 240 rnin), followed by 60 rnin soak at 450°C. 'slow heating from 20450°C (within 240 min). followed by 240 min soak at 450°C. Table 2 displays values of the charcoal and fixed-carbon yields obtained from a variety of biomass feedstocks by pyrolysis at elevated pressures. The f values associated with oak wood represent the sample standard deviation of four identical runs on the pilot plant. In Table 2, charcoal yields range from 34.6% (birch wood) to 58.6% (kukui nut shell). The average fixed-carbon yields for these two feedstocks are 27.2% and 41.4% (respectively). A comparison of the elevated pressure fixed-carbon yields reported in Table 2 with those obtained by SINTEF at atmospheric pressure and low heating rates (see Table 1) shows that elevated pressures significantly improve the fixed-carbon yields for each of the wood species studied by both groups. A more comprehensive study involving 5 different wood feedstocks, reported by the authors in reference (1 0), corroborates the beneficial effects of elevated pressure for charcoal production. Remarkably, carbonization at elevated pressure realizes high yields
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without a prolonged reaction time. When the pressure vessel is already hot (as would be the case in a continuously operating commercial reactor), the feed is completely carbonized after about 1’ h of reaction time. Table 2 Charcoal and fured-carbon yields realized at 1.O MPa Feed
Proximate Analvses FU-C (*A)
(we?)
Average Charcoal ash Yield’ (weh) (*A)
Average FixedC Yield’
Theor. FixedC Yield
(we?)
(we?)
Bamboo
27.0
67.1
5.9
45.9
32.1
33.1
Birch Wood
20.9
78.5
0.6
34.6
27.2
32.4
Garlic Waste
27.3
49.2
23.5
42.9
25.5
27.6
Kukui Nut Shell
26.9
69.6
3.5
58.6
41.4
42.0
Leucaena Wood
27.0
69.4
3.6
44.2
32.3
32.9
24.0 f 2.0
75.0
1.0
39.8 f 1.3
29.9 f 1.1
34.2
Oat Hull
24.9
62.3
12.8
40.2
26.3
32.1
Rice Hull
14.9
47.2
37.9
46.2
27.6
32.1
oak wood3
’% of dried charcoal. *% of dried feed material. Fixed-C Yield = Charcoal yield %Char Ash) / (100 - %Feed Ash).
* (100 - %Volatile Matter -
’Results from process development unit (PDU).
The thermochemical equilibrium value of the fixed-carbon yield constitutes a benchmark against which the experimental values can be compared. We refer to this value as the “theoretical” yield of carbon, which we hypothesize to be the upper limit attainable by thermal processes. We used the thermochemical equilibrium software STANJAN (1 1) to calculate the mass fraction of solid carbon present at equilibrium when the elements C, H, and 0 with molar ratjos dictated by the composition of each feedstock (see Table 3) are mixed at 1.O MPa and 300 OC. We neglect the presence of nitrogen and sulfur in the STANJAN calculations because these two elements compose only a small fraction of the mass of the biomass feed. The ash content of the feed is also neglected. In general, STANJAN predicts that solid C, and the gases C a , H20 and C h should be the only significant products present in equilibrium, and that the distribution of these products is not strongly dependent q,sn either the assumed pyrolysis temperature or the assumed pressure. The predicted thermochemical equilibrium values for the fixed-carbon yields are displayed together with the experimental values in Table 2. The measured values of the fixed-carbon yields for kukui nut shell, leucaena wood, and bamboo effectively equal the theoretical yields within the margin of error of the data.
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Table 3 Elemental analyses of feedstocks. C (wr??)
H (wr??)
(we!)
(we!)
N
S (wr??)
ash (wt%)
Bamboo
47.65
5.77
44.23
0.27
0.1 1
3.91
Birch Wood
48.07
6.00
45.56
0.17
<0.05
0.20
Garlic Waste
37.85
4.97
43.12
0.49
0.22
17.07
Kukui Nut Shell Leucaena Wood Oak Wood
55.76
5.60
37.99
0.34
0.03
1.45
45.95
6.06
41.23
2.42
0.27
5.1 1
50.13
5.98
44.76
0.08
0.03
0.19
Oat Hull
46.00
5.91
43.49
1.13
0.15
4.91
Rice Hull
38.86
4.86
37.15
0.42
0.06
20.97
0
Remarkably, agricultural wastes offer yields of charcoal and carbon that are comparable to hardwoods. Rice hulls are a particularly interesting example because of their widespread availability and difficulties associated with their disposal. Based on our measured fixed-carbon yields, we estimate that 370,000 mt of fixed carbon could be produced from the 1,700,000 mt of rice hulls that are generated annually in the USA. High-yield charcoal produced from these rice hulls could satisfy the Norwegian ferrosilicon industry's current and future demand for carbon. Because rice hull charcoal consists of small particles, industry might not consider it to be suitable for use as a reductant. But this charcoal can be formed into briquettes, which resist crushing and are durable. Inexpensive charcoal briquettes made from agricultural wastes could cause a paradigm shift in the way metal ore is refined around the world. When a proximate analysis of charcoal is conducted, the charcoal is heated in a covered crucible to 950 OC and held at this temperature for about 6 min. The carbonized charcoal product of this process possesses interesting properties. It is a relatively pure carbon. For example, the ultimate analysis of carbonized oak wood charcoal is 92.8% C, 1.09% H, 3.49% 0, 0.24% N, 0.04% S , and 1.46% ash. Table 4 presents measured values of the resistivity of ##40mesh oat hull charcoal carbonized at different temperatures. The resistivity decreases by six orders of magnitude and reaches a value of 3.7 R - cm for carbonization at 900 OC. For comparison sake, we used the same apparatus to measure the resistivity of a packed bed of commercial grade graphite powder. As expected, the value we obtkined, 0.16 R - cm, was somewhat smaller than that of the carbonized charcoal. Nevertheless, the low electrical resistivity of carbonized biomass charcoal opens some interesting doors. In their classic text on fuel cells, Bockris and Srinivasen (12) conclude that carbon fuel cells are impractical because (i) coal is not an electrical conductor, and (ii) graphite is too scarce and expensive to be used as a fuel. In light of the good electrical conductivity of carbonized biomass charcoal, we believe that the research community should give attention to the development of a biocarbon fuel cell.
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Table 4 Measured electrical resistivity of a packed bed of carbonized oat hull charcoal (#40 mesh) compressed at 846 H a .
Carbonization Temperature ("C) 600
Resistivity (a- cm) 7.4 x lo6
700
9.9 x lo4
800
35
900
3.7
Carbonized charcoal is also suitable for the production of very high quality activated carbons. The conventional method for physical activation of charcoal involves gasification in steam or carbon dioxide at high temperatures. The yield of activated carbon from dry biomass for this process is typically about 8 wt%. To improve this yield, and to develop a less costly method for activation, we initiated research on a novel activation process involving (i) carbonization of the charcoal at temperatures below 750 "C, (ii) air oxidation in boiling water (AOBW) of the carbonized charcoal at temperatures below 240 "C, and (iii) activation (a second carbonization) of the oxygenated carbon. High surface area (>900 m2/g) activated carbons have been produced in our laboratory from various agricultural wastes by the AOBW process (13). In step (ii) the boiling water controls the oxidation temperature and holds the reaction rate at a relatively low value; thereby precluding mass transfer limitations that quickly appear at higher temperatures. Under these conditions dissolved oxygen in the yvater is able to d i f i s e into the micropores and efficiently remove carbon. Note that the equilibrium concentration of dissolved oxygen in boiling water at 240 "C exceeds the concentration of oxygen in air at room conditions. Yields of activated carbons produced by the AOBW process are about 100% higher than those reported by industry for the synthesis of coconut shell activated carbon. CONCLUSIONS (1) In the case of oak wood used to manufacture charcoal in Missouri kilns, pyrolysis at 1.0 MPa increases the fixed-carbon yield from 124 wt% in an externally heated retort (or the Missouri kiln) to 29.9 wto/o in the pressure vessel. The time requirement is reduced from =10,000 min for the Missouri kiln to about 70 min in the pressure vessel. (2) High yields of fixed-carbon are realized at 1.0 MPa from a wide variety of agricultural wastes. For example, rice and oat hulls offer fixed-carbon yields of 27.6 and 26.3 wt% (respectively). These yields are higher than those obtained from traditional hardwood feeds used by industry to manufacture charcoal in Missouri kilns. (3) More attention should be given to biomass species that are prone to the formation of carbon by pyrolytic methods. Likewise, industry should give more attention to the praduction of carbon from agricultural wastes. Briquettes manufactured from rice hull charcoal could satis@ the Norwegian ferrosilicon industry's current and future demand for renewable carbons. The substitution of renewable carbon from biomass for coal in
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the metallurgical industry and elsewhere can reduce greenhouse gas emissions while creating new jobs in the developing world and enhancing the profitability of agribusiness. (4) Carbonized biomass charcoals are good conductors of electricity. Consequently, carbonized biomass charcoal could be used to fuel a biocarbon fuel cell. We urge the research community to give attention to the development of a biocarbon fuel cell. ( 5 ) Very high quality activated carbons can be produced from carbonized biomass
charcoals by conventional gasification processes, but the yield is low. We urge the research community to give attention to improving the yields of activated carbons from biomass.
REFERENCES 1. 2. 3. 4. 5. 6. 7. 8.
9. 10. 11. 12. 13.
M. K. Hubbert, Transactions of the American Society of Mechanical Engineers 101, 16-30 (1979). C. J. Campbell, J. H. Laherrere, Scientijc American, 78-83 (1998). N. Shah, P. Girard, C. Mezerette, A. M. Vergnet, FUEL 71,955-962 (1992). K. R. Smith, Energyfor Sustainable Development 1,23-29 (1995). “Charcoal: Production, Marketing, and Use” US.Dept. of Agriculture Forest Service Report #2213 (1961). M. J. Antal, Jr. et al., Energy Fuels 10,652-658 (1996). H. Yamamoto, K. Yamaji, J. Fujino, Applied Energy 63, 101-1 13 (1999). W. S. L. Mok, M. J. Antal, P. Szabo, G. Varhegyi, B. Zelei, Industrial and Engineering Chemistry Research 31, 1162-1 166 (1992); W. S. L. Mok, M. J. Antal, ThermochimicaActa 68, 165-1 86 (1983). X. Dai, Ph.D. Thesis, University of Hawaii at Manoa (1998). M. J. Antal et al., submittedfor publication in Ind Eng. Chem. Res.. W . C. Reynolds, STANJAN Thermochemical Equilibrium Sofrware (Stanford University, Stanford, CA, ed. 3.91, 1987). J. 0. M. Bockris, S . Srinivasan, Fuel Cells: Their Electrochemistry (McGraw-Hill, New York, 1969). M. S. Tam, M. J. Antal, E. Jakab, G. Varhegyi, Submitted to Ind Eng. Chem. Res.
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Issues in Value-Added Products from Biomass D. C. Elliott Pacific Northwest National Laboratory, Richland, Washington, USA
ABSTRACT: While methods for convertingbiomass to alternative energy products has been developing over the past 30 years, process developmentfor chemical products is an area of increasing interest and effort. Although the need for renewable resources in the fuels market is still likely to have the most economic impact worldwide, chemical products derived from biomass can also be more thanjust niche market applications. However, the specific chemical processing required for refined products requires improved chemical handling methods for separations and purifications, as well as improved catalyst systems. Development ofthese unit operations has lagged behind the process research focused on the finshed products. This paper describes the critical processing issues that must be addressed to more effectively use biomass feedstocks for producing value-added chemicals. SpeciSc mas include biomass component separation,catalyst developmentfor aqueous processing, and trace component effects in catalytic processing of biomass feedstocks. The paper also describes some of the process research which has been performed or is now underway in our laboratory and others.
THERMOCHEMICALCONVERSION TO VALUE-ADDED PRODUCTS Thennochemical processes for producing value-added chemicals from biomass have followed similar paths as processes for generating alternative fuel products, namely gasification and liquefaction methods. These methods are discussed below for both fuels and chemicals production to give an overview of the current technology.
PROGRESS IN FUELS PRODUCTION The impetus for using thermochemical conversion of biomass to produce value-added products began with the need for new fuels to displace limited petroleum resources. Since the oil embargo in.the mid-l97Os, work has been conducted worldwide to develop new processing methods for alternative fuel products. The bulk of the research and development has focused on generating liquid and gaseous fuel products, with only limited work on direct combustion of solid biomass. Liquefaction processes for biomass conversion can be categorized as slurryphase, high-pressure processing, usually with catalysts; or as dry pyrolysis at low or reduced pressure'. The research has concentrated on the fuel properties and has often overlooked the detailed chemical analysis of the biomass-derived products. Both process types produce a complex mix of oxygenated organic components. The
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mixqure has a lower heating value than petroleum liquids and is more chemically reactive. Gasification systems have typically been either pyrolysis systems, often with added steam, or partial oxidation systems, using either air or oxygen2. Direct addition of catalysts into these systems has usually not been successful. Analysis of only the low molecular weight components in the gas product has been the typical limitation of these process development efforts, with the residual high-molecular weight product classified as "tar". PROGRESS IN CHEMICALS PRODUCTION
The recent interest in chemical production is based on a higher return expected for chemical products versus fuels. For example, biomass gasification can be used to produce a synthesis gas of hydrogen and carbon monoxide. This gas can be used in catalytic synthesis of a range of chemicals, from methanol and formaldehyde to higher hydrocarbons, in the same way that synthesis gas derived from natural gas can be used3. However, by breaking down the biomass to the basic building blocks all product differentiation relative to fossil fuels is lost. Biomass liquefaction methods can also be used to produce chemical products. With the more mild treatment at liquefaction conditions, the intrinsic properties of biomass composition are retained to a greater degree. Both pyrolysis and highpressure process conditions can be used to generate value-added chemical products. Pyrolysis of biomass produces a complex product slate of oxygenated chemicals, only a few of wluch are present in more than 1% concentration. In earlier times, most of the methanol and acetic acid was produced from wood pyrolysis, but such processing is not economical in the current petroleum-based economy. Today, liquid smoke, a food flavoring containing a mixture of pyrolysis products, is a well-known chemical product from wood pyrolysis. New processing methods and new products are now being developed, such as resins4, levogluco~an~, and specialty chemicals6, that use flash pyrolysis-derived bio-oil. Pressurized catalytic processing may also provide future opportunities for chemicals production. Such systems have been developed for methane' and hydrogen* production. Commercialized recovery of liquid chemicals requires advancing the teclmology of separation and concentrationfor these processes, however. More recent chemical production methods using biomass seem to be utilizing pretreatment technologies from the past. Whereas pretreatment in thennochemical conversion is geared to material handling issues such as densification and feeding, pretreatment in bioconversion involves concentration and purification of the glucose or carbohydrate components to maximize the processing efficiency of the microbes". The preconcentration of the useful biochemical feedstock components and removal of deleterious components, as done in biochemical conversion, is also beneficial in thennochemical processing. Direct conversion of cellulose in an unpurified state has been shown in the case of levulinic acid production. Production of levulinic acid as a high-yield direct product of controlled acid hydrolysis of cellulose is an exception to the complex product slate typically produced from biomass feedstocks. While the levulinic acid is only derived from the cellulose portion of the biomass and other components end up as byproducts, technology has now been demonstrated for recovery of the levulinic acid product at yields and purity sufficient to generate market interest". Based on related research developments, levulinic acid may prove to be an important building block for
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chemical products derived from biomass'*. In order for new technologies to be developed for chemical production from biomass, issues related to the feedstock and aqueous processing environment need to be addressed. BIOMASS FEEDSTOCK SOURCES AND PRETREATMENT PROCESSING
Biomass feedstocks usually contain a wide variety of chemical functional types within the biopolymer structure. Since thermochemid conversion has typically focused on the use of the "whole" biomass, separating the chemical functional types has been of less interest. Using the "whole" has been viewed as the most cost-effective approach while the thermochemical processes were considered robust enough to handle the range of chemical functional types. As a result, the products were a complex mixture of chemical entities useful primarily for he1 applications. The costs of collecting individual chemical products could not be justified in most cases. A new processing paradigm is now being developed based on using cleaner, simpler biomass-derived feedstocks to address the issue of biomass complexity. The complexity issue has two facets, the mixture of organic functional types and the inclusion of inorganic components. Although wood is "clean," in terms of being low in sulfur and nitrogen (two major pollution sources found in coal and petroleum), it is still complex structurally, containing aromatic (lignin) components - like coal - but also both polymeric Ca and CS structure. It also has significant levels of mineral content much lower than coal but much higher than petroleum. Other biomasses can contain very high levels of mineral content (grass straws) or nitrogen (all types of green foliage or seeds). Complication of the organic structure with protein components is commonly found, and the presence of lipids is a complication in some other cases. The different specific components are useful as feedstock for Werent chemical products. In order to effectively use biomass for chemical production, the various components should be used separatelyfor feedstock.
-
PLANT-DERIVED SUCARBTARCH FEEDSTOCKS
There is no directly recoverable glucose feedstock from biomass. Even the sugar feedstock produced directly by plants in some cases, such as beet or cane sugar (sucrose), is actually a disaccharide. It must be processed through hydrolysis (in this case known as "inversion") to be useful as a glucosdfiuctose mixture for chemical production. However, these sugars are generally too expensive except for specialty chemical production, such as sorbitol or mannitol, or in cases of special economic circumstances, such as a captive market or high transportationcosts. Cornderived starch is a readily available feedstock in the U.S.It is produced efficiently in large quantities in corn wet mills and provides the basis for much of the biomassderived chemical production presently in the market. The starch is enzymatically hydrolyzed on a large scale to produce glucose, much of which is isomerized to fructose for use as food sweetener, and the rest mostly goes into ethanol production for automotive fuel. The several separation processing steps in the corn wet mill are summarized in Figure 1. Equivalent processing could be applied to wheat and potato or other starch when the feedstock is available at a suitably low price. Some of the cornderived glucose goes into various fermentation processes that lead to other chemical products. For example, large-scale production was recently announced for lactic acid and polylactide plastic derived from it". Citric acid and lysine are two other important fermentation products.
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Corn,
to oil
conversion
Figure I Glucose Recovery from Corn HYDROL Yz4TE FEEDSTOCKS AND DERIVED FEEDSTOCKS
Hydrolysis of biomassderived carbohydrate could produce inexpensive feedstock for processing to chemicals. [The technical aspects of hydrolysis are discussed in detail in the next section.] The major stnrctural components of biomass, cellulose and hemicellulose, can be broken down to the individual sugar monomers which can be used as feedstock for chemical production. The monomers can be processed directly by thennochemical means, such as catalybc hydrogenation, or can be biopmessed into to various fermentationderived feedstocks. Processes are being developed for fermentationproducts, including lactic acid (mentioned above) and succinic acid. The succinic acid has potential application for polyester monomer (1,4-butanediol)and solvent (tetrahydrofuran) pr~duction'~. PULPING ANALOGS DERIVED FEEDSTOCKS
Various pulping technologies for wood and other biomass have been used for centuries to produce fiber, primarily for paper production. These processes effectively separate
and clean the cellulose for use as fiber. However, the hemicellulose and lignin byproducts are not typically used for chemical products. The depolymerized hemicellulose in pulping liquor is a large source of biomass carbohydrate that could potentially be used for chemical production.
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Another related processing byproduct, available in large quantities, is the unrecovered fiber disposed as paper mill sludge. This cellulosic feedstock is usually contaminated with various paper-making chemicals, such as clay and lime which would likely require separation before the cellulose could be used. However, processes are being developed that directly use the cahohydrate, including the levulinic acid process mentioned above. New pulping technology is in development that could lead to more effective recovery of the biomass components in a higher purity form. Such "Clean Fractionation" technology is expected to produce high purity streams of cellulose, lignin (in solution), and hemicellulose (as dissolved sugars) by a specially developed solvent pulping process". The cellulose and lignin would be particularly clean products, while the hemicellulose stream would contain residual dissolved mineral matter. ANIMAL WASTE FEEDSTOCKS
The use of animal livestock waste for chemical production would provide a solution for this otherwise problematic material. Because of increased large-scale livestock operations, the waste has become locally concentrated to levels that often exceed environmentally desirable disposal methods. The composition of animal waste is not adequately documented to determine which chemical manufacturing processes would benefit from its use as a feedstock. The limited information on the composition of carbohydrates, protein structure, and other components needs to be expanded. The secondary issue is the need to develop appropriate separation and cleanup processes to produce useful and consistent feedstock for chemical production. It is clear that there are economically interesting amounts of waste available for processing if sufficient amounts of feedstock can be economically derived from it. BIOMASS UTILIZATION ISSUES
Processes for depolymerization of the biopolymer, separation of the various chemical types, and aqueous phase processing all must be established for biomass conversion to chemical products to become widespread. DEPOLYMERIZ4 TION
Depolymerization of the biopolymer structure is an issue in almost all utilization concepts. The established chemical technology using cellulose as a base material is a signficant exception. Recovery of proteins is proving to be a second major exception. For the most, part the carbohydrate structure needs to be broken down into useful monomers like glucose or xylose. Over the years, the assessment of hydrolysis with catalysts, both mineral and biological, has been evaluated with a range of biomass feedstocks. Acict/basehydrolysis
Autohydrolysis is the hydrolysis without added catalyst. In fact, it requires the presence of water and heat that likely generate the initial acid sites (acetate in the case of biomass), which then further catalyze continuing hydrolysis. Hydrolysis is generally practiced in a staged method, using Merent concentrations of acid and
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processing temperatures to maximize the yield of desired sugar products and minimize the formation of degradation products. Acid hydrolysis of wood is a well-known procedure, but it has never been considered an economical process for sugar production. Most of the technology developed for wood has broader application to most biomass. In Europe initial efforts focused on high concentration acid (40% H2S04 or HCl) at high temperatures (150250°C). Later efforts focused on a low-concentration acid first-step for hemicellulose removal, followed by a higher concentration acid step to break down the more refractory cellulose. Refinements of this type of processing have proceeded unabated. A recent search of the U.S. patent database shows over 200 patents describing acid hydrolysis of biomass just since 1976. Dilute acid hydrolysis has also been developed as a pretreatment for enzymatic hydrolysis’6 Base hydrolysis has also been practiced for depolymerization of biomass. Alkali pulping of biomass is an industrial example of a well-known way to break down the hemicellulose without destroying the cellulose. In other cases, base hydrolysis has been used as a pretreatment to facilitate microorganism access to the carbohydrates for bioprocessing of biomass. In some cases, the base is added in the form of ammonia which has the advantage of providing nitrogen nutrient value for the microorganisms as well. More intensive base treatment has been used to attack cellulose for the production of organic acid products. Krochta’s group has evaluated a range of base treatments that result in the production of acids like formic or acetic, as well as several hydroacids and diacids”.
Envmatic hydrolysis Enzymatic hydrolysis has received great attention in recent years. In combination with fermentation, it has been found partmdarly attractive. The value of the enzymatic process is its specificity for the reaction and resulting high product selectivity. This also means that the enzymes have limited application, and often a suite of enzymes must be developed to achieve an effective depolymerization of the complex structure in a typical biomass material. Wood hydrolysis has been limited to cellulose degradation by cellulase enzymes. These enzymes are typically low activity and highly inhibited by the glucose product. As mentioned above, mild acid pretreatment has been found to be an important first step in the biomass utilization process. The pretreatment is used both to break down the hemicellulose to sugars and to disrupt the lignocellulosic structure and the crystallinity of the cellulose. Depolymerization of hemicellulosic structure is still being developed. Because of its more complicated structure involving several kinds of sugar units in variously linked conformations, more than one specific enzyme is needed. More simple plant starches can be effectively broken down to sugars with combinations of enzymes, including amylases and glucoamylases. COMPLEX MIXTURE SEPAM TION
Following the depolymerization of the carbohydrate structure of the biomass, numerous separations will likely be required. These complex mixture separations are at the forefront of the process development effort for biomass utilization for chemicals production. Several technologies are available for other applications and are being developed for the specific needs of biomass processing.
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An interesting model can be found in the existing and developing technology for cheese whey utilization. Cheese whey, the residual aqueous product from cheese production, is a dilute solution (6% dry solids), with a major component of carbohydrate (75% lactose, dry solids) and lesser amounts of protein, and minerals as well as trace amounts of lipids and vitamins. Separations have been based on precipitation of lactose and minerals following concentration to produce food-grade products. A more recent development is the recovery of a protein concentrate by ultrafiltration. The resulting permeate still contains a significant amount of lactose with minerals, vitamins, and non-protein nitrogenous compounds. The mineral components may be removed by ion exchange or precipitation. Lactose recovery may either be by utilization of the solution or by concentration, precipitation, and recrystallization to a purified form. Recovery of the sugars from solution may or may not be an important processing step. Separation of individual sugars may allow production of higher valued specific products. Such separation is possible using simulated moving bed continuous chromatography processing. This type of processing can also be used for final purification by removing residual mineral and protein from the carbohydrate stream. However, the several sugars derived from biomass may be processed to final products before separation in some cases. This option is particularly applicable to thennochemical processing, in which the catalyst and heat are less specific in the reaction mechanism than enzymatic processing, in which the biocatalyst differentiates between various isomers. The use of the sugars in solution or concentration and recovery for final use is also an important question. The aqueous processing environment raises important processing issues as will be described below. Use of the dilute solution requires considering the tradeoff of additional reactor volume versus effects of concentration on thermodynamic equilibriums of the reactions involved. The solution can be concentrated through straightforward evaporation, but due to the high heat of vaporization of water, there is a significant energy cost. While water is an excellent solvent for the sugars, it is low cost as well. Because of the good solvent properties of water for the sugars, concentration by solvent extraction is not an attractive or effective alternative, even disregarding the fact that such an option almost undoubtedly introduces a less environmentally friendly and more expensive solvent to the process. The most cost-effective approach appears to be to remver the sugars from the water but to proceed with the processing to the point that a product can be more efficiently recovered. AQUEOUS PHASE CATALYTIC PROCESSING
Use of catalysts in an aqueous processing environment has received much less attention than more traditional catalytic processing. Most heterogeneous catalyst development work is focused on gas phase processing or on hydrocarbon (petroleum) processing. In either case, there is only a limited partial pressure of water in the reactor and no liquid phase water. Use of the aqueous phase to process biomassderiv,ed sugars requires understanding the high-pressure liquid water phase as the reaction environment. In this environment, conventional catalyst formulations have limited lifetimes because of catalyst metal oxidation or sintering andor catalyst support hydrolysis or dissolution. The use of biomassderived feedstock can also lead to the contamination of the catalyst with dissolved mineral species and other inorganic components such as sulfur.
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Catalyst material stability in the aqueousphase
Chemical process development underway in our laboratory is using high-pressure liquid water as the processing medium. Catalysts would be incorporated in many of the processes being studied. To date, commercial catalyst formulations tested have been unacceptable for chemical processing in the high-pressure liquid water environment. We have reported earlier that conventional support materials, including y-alumina (and other high-surface area aluminas), cements like calcium aluminate, silica, silica-duminas, hydrous titanates, and natural minerals like kieselguhr, have not been usable at 350OC and 200 a d 8 . Pelletized forms of these materials soften, and actual dissolution has occurred in some cases. Reaction of the aluminas to a hydroxide form (bohmite) has also been confirmed. Catalytic metal species are often unstable in high-pressure liquid water environments. Of the first row Group VIII metals, only nickel remains reduced at 350°C and 200 atm; both iron and cobalt are readily oxidized. Of the other transition metals tested to date, copper and Group VIII noble metals remain d u d , but zinc, chromium, tungsten, and molybdenum are oxidized. Even though nickel remains reduced at operating conditions, it is not stable with respect to crystallite size. Based on x-ray diffraction approximations, average crystallite size increase can account for dramatic reduction of nickel surface area. In addition, nickel metal is subject to dissolution into the aqueous processing medium leading to loss of catalyst material and contaminationof the products. We have found, however, that a ruthenium metal catalyst on titania support has high activity in these systems. Because the titania is stable in the aqueous processing environment, the catalyst remains physically intact. The ruthenium is also stable in tlus environment and does not oxidize or sintedagglomerateor dissolve into the water. In addition, because of its high specific activity, the ruthenium can be used at low metal loadings of only 3 wt% compared to nickel metal catalysts of 50 wt%, or nearly 100% in the case of Raney nickels. Also because of its high activity, this catalyst can be used at significantly lower processing temperatures and still attain high processing effectiveness. The form of the titania is also an important issue. We have found that the rutiie form is the preferred component for use in the aqueous processing environment. The anatase form is readily converted to the rutile in the aqueous environment over a period of hours to days depending on the operating temperature. As a result, ruthenium catalysts formulated on anatase support lose activity over time, while those formulated on mile maintain their activity at operating conditions. Cutulystprocessing results with inrprovedformulatwns
We have demonstrated that the ruthenium on titania catalyst is useful for the specific applications of lactose, glucose, and xylose hydrogenation to lactitol, sorbitol, and xylitol, respectively. In these applications, the use of the new catalyst allows the process to be accomplished at lower processing temperature while maintaining high processing rates and high conversion and product selectivity. As an example, some typical processing results are presented in Table 1 for the new ruthenium on titania catalyst. This high activity is maintained from startup, while the nickel catalysts lose activity over a period of hours to arrive at a typical operating activity.
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Table 1 Hydrogenations of lactose, glucose, and xylose.
Feedstock (20wt%)
Lactose Lactose Lactose Glucose Glucose Xylose
Temp. (OC) 100 100 100
100 100 100
Pressure (pig) 1900 1900 1900 1900 1900 1900
LHSV L/Lh 2
3 4 4 6 4
Feedstock conversion (%) 100 99.7 99.2 100 99.95 100
Selectivity to sugar alcohol 95 96 98 97 99.8 99
Comparativetests have been performed in the semi-batch reactor system to evaluate the Ru/Tiq catalyst versus a more conventionalnickel-based catalyst. These tests show that ruthenium at only 3% metal loading has about the same activity as nickel at 50% metal loading. This comparison is only for short-term activity of the catalyst. As demonstrated in the continuousflow tests, the nickel catalyst loses activity readily in tlie first hours on stream, while the ruthenium maintains its activity. Table 2 Semi-batch hydrogenations of lactose, glucose, and xylose.
Feedstock
Temp.
(12 wt%)
("c)
lactose/Ru lactosemi lactose/Ru 1actoseNi
100 100 140 140
Pressure Time at (psig) Temp. (hr) 1500 3 1500 3 1500 1500
1 1
Feedstock conversion (%) 98.8 98.8 95.3 98.6
Selectivity to sugar alcohol 97 88 85 64
Poisoning of cutubsts.in aqueousphase processing of biomass-derivedfeedstocks
The processing of mineral-loaded, biomass-derived sugar feedstocks can be problematic. We have reported in tests with cheese whey and brewer's spent grain the calcium and magnesium precipitate on the catalyst as various phosphate minerals, e.g., whitlockite or d~rapatite'~. Removing these elements from the feedstock may be required even at the lower temperatures used for chemical production, e.g., catalytic hydrogenation at 15O-25O0C, as opposed to the higher temperature tested in the reference. However, avoiding the near-critical temperature processing condition may facilitate the use of the contaminated feedstocksfrom biomass. The sulfur content of biomass can also be important when metal catalysts are involved in the processing. Sulfiu poisoning of nickel metal catalysts is well known, particularly as a result of petroleum processing. The difficulty led to the development of nickel modified molybdenum catalysts, which are used in the sulfided form at relatively high temperatures for some hydrogenation mctions in treatment of petroleum distillates. In the case of biomass, sulfur content is mostly contained in protein components or as sulfates. Removing the protein can likely reduce the sulfUr content to an insignificantly low level. Sulfate removal may require anion exchange.
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CONCLUSIONS Effective utilization of biomass for value-added chemical product synthesis will require development of new applications of important unit operations. Carbohydrate recovery from the biomass is the key near-term application for production of commodity chemicals. Protein recovery will continue to have an important niche market in the purified form as food and a larger low-value market in the crude form as animal feed. Important processing information for carbohydrate depolymerization can be found in the literature from biochemical conversion of biomass. New process applications of separation technologies are just now being developed and refined for use with biomass-derived carbohydrate and protein streams. The use of an aqueous processing environment for carbohydrates will require careful consideration of the differences that type of environment entails, such as the effect on catalyst formulations.
REFERENCES 1 . Elliott, D. C., Beckman, D., Bridgwater, A. V., Diebold, J. P., Gevert, S. B., and Solantausta, Y. (1991) Developments in Direct Thermochemical Liquefaction of Biomass: 1983-1990. Enetgy & Fuels, 5(3), 399-410. 2. Beenackers, A.A.C.M. and Maniatis, K. (1996) GasificationTechnologies for Heat and Power from Biomass. In: Biomassfor Energy and the Environment, (Ed. P. Chartier, G.L. Ferrero, U.M. Henius, S. Hultberg, J. Sachau, and M. Wiinblad) pp. 228-259, PergamonElsevier, Oxford, United Kingdom. 3. Anikeev, V. (1999) Modem Technologies of the Biomass Conversion for Chemicals, Carbon Sorbents, Energy, Heat and Hydrocarbon Fuels Production. In: Biomass: A G m h Opportunity in Green Energy and klue-Added h d u m . (Ed. R. €? Overend and E. Chomet), Wume 1, p p 563-570. Elsevier Science, Ltd., Kidlington, United Kingdom. 4. Roy, C., Calve, L., Lu, X., Pakdel, H., and Amen-Chen, C. (1999) Wood Composite Adhesives from Softwood Bark-Derived Vacuum Pyrolysis 0ils.h: Biomass: A Growth Opportunity in Green Energy and Wue-Added h d u m . (Ed. R. I? Overend and E. Chomet), Ulume 1 , p p 521-526. Elsevier Science, Ltd., Kidlington, United Kingdom. 5. Pernikis, R., Zandersons, J., and Lazdina, B. (1997) Obtaining of Levoglucosan by Fast Pyrolysis of Lignocellulose. Pathways of Levoglucosan Use. In: Developments in ThermochemicalBiomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), pp. 536-548. Blackie Academic & Professional, London. 6. Scott, D.S., Legge, R.L., Piskotz, J., Majerski, P., and Radlein, D. (1997) Fast Pyrolysis of Biomass for Recovery of Specialty Chemicals. In: Developments in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), pp. 523-535. Blackie Academic & Professional, London. Butner, R.S., and Sealock, L.J., Jr. (1993) Bench-Scale 7. Elliott, D.C., Baker, E.G., Reactor Tests of Low Temperature, Catalytic Gasification of Wet Industrial Wastes. Journal of Solar Energy Engineering, 115,52-56. 8. Minowa, T., Ogi, T., Yokoyama, S . (1997) Hydrogen Production from Lignocellulose Materials by Steam Gasification Using a Reduced Nickel Catalyst.
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In: Developments in ThermochemicalBiomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), pp. 932-944. Blackie Academic & Professional, London. 9. Miles, T.R. (1984) Biomass Preparation for Thermochemical Conversion. In: Thermochemical Processing of Biomass. (Ed. A.V. Bridgwater), pp.69-90, Butterworths, London. 10. Robinson, J.S. (1980) Primary Biochemical Conversions, In: Fuelsfrom Biomass: Technology and Feasibility. pp. 167-183, Noyes Data Corporation, Park Ridge, New Jersey. 11. Denn, J. (1997) Test Plant Aims to Make Something of Nothing. Albany Times Union, p. E l & E4, November 29, 1997, Albany, New York. 12. Elliott, I3 C,Fitzpatrick, S W,Bozell, 1 1,Jarnekld, 1 L., Bilski, R. 1, Moens, L., Frye, 1 G., JL, Wing, Y and Neuenschwndel; G. G. (1999) Production of Levulinic k i d and Use as a Platform Chemical for Derived Products. In: Biomass: A Growth Opponuniry in Gmen Energy and klue-Added h d u c t s . (Ed. R. I! Overend and E. Chomet), Wume 1, pp 595-600. Elsevier Science, Ltd., Kidlington, United Kingdom. 13. Brown, R. (2000) CDP Will Build Huge PLA Unit at Cargill's Site. Chemical Marketing Reporter, 257(3), p. 1. 14. Werpy, T.A., Frye, J.G., Jr. et al. (2000) Production of Succinic Acid Based Renewable Chemicals. In: ACS Symposium Series, to be published. 15. Black, S.K., Hames, B.R., and Myers, M.D. (1998) Method of Separating Lignocellulosic Material into Lignin, Cellulose and Dissolved Sugars. U.S. Patent No. 5,730,837. 16. Torget, R., Hiinmel, M., Wright, J.D., and Grohmann, K. (1988) Initial Design of a Dilute Sulfuric Acid Pretreatment Process for Aspen Wood Chips. Applied Biochemistry and Biotechnology, 17,89-104. 17. Krochta, J.M., Tillin, S.J. and Hudson, J.S. (1988) Thermochemical Conversion of Polysaccharides in Concentrated Alkali to Glycolic Acid. Applied Biochemistry and Biotechnology, 17,23-32. 18. Elliott, D.C., Sealock, L.J., Jr., and Baker, E.G. (1993) Chemical Processing in High-pressure Aqueous Environments. 2. Development of Catalysts for Gasification.Industrial and Engineering Chemistry Research, 32, 1542-1548. 19. Elliott, D.C., Phelps, M.R., Sealock, L.J., Jr., and Baker, E.G. (1993) Chemical Processing in High-pressure Aqueous Environments. 4. Continuous-Flow Reactor Process Development Experiments for Organics Destruction. Industrial and Engineering Chemistry Research, 33, 566-574.
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Combined Chemicals and Energy Pmduction From Biomass Pymlysis D.A. Himmelblau Biocarbons Colporation, 71 Cum m ings Park, Wobum, M A 01801, USA and R.W . Beck, P.O. Box 9344, Framingham, M A 01 701, USA
ABSTRACT: Whde the emphasis on biomass pyrolysis liquids has been fuel production, higher or added value can be obtained by chemical production. Pyrolysis oil yield and composition are functions of reactor design and operating conditions, especially temperature, and the raw pyrolysis oil composition can be modified to improve yield and selectivity for higher value chemical products in order to produce a better return than from fuel use alone. The pyrolysis products that are currently commercial chemicals will have to compete with conventional petrochemical or agrichemical production, especially in terms of product purity and scale of production. Unfortunately, ready markets do not exist for many of the potential chemical products from biomass pyrolysis such as levoglucosan, because they have not been previously available in large quantities at reasonable cost. The markets cannot be developed without additional application research and distribution channels. Chemical company partners will be needed. This paper uses pyrolysis products from an air-blown, fluidized-bed reactor designed to produce liquids for use in adhesives as an illustration of what might be produced, separated and sold commercially for higher value. Guidelines for chemicals production and product recovery are suggested for pyrolysis processes in general. Recommended research and development topics to aid commercialization and increase chemical product recovery and project cash flow are presented.
INTRODUCI'ION The purpose of this paper is to guide or at least stimulate thinking about obtaining as much value as possible from the chemicals produced from biomass pyrolysis. At optimum, all solid and liquid products could be used for chemicals, rather than fuel. At minimum, additional cash flow could be produced by recovery or production of any high value chemicals in the liquid product, leaving the balance of the products to be used as fuel. The data in this paper is specific to the air-blown, fluidized-bed reactor system used
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by Biocarbons Corporation (5,6) that produces a water-insoluble liquid yield of up to 25 weight percent (of dry feed), a char yield of up to 15 weight percent and the balance as gas, mostly CO,, I&, CO, 50 and CH,. The reactor system was specifically chosen to produce a liquid stream that could be used (as it turns out in its entirety) as a lower cost substitute for phenol in the phenol-formaldehyde adhesive resins used for making water-resistant wood panels such as plywood and oriented strandboard. Recovery and/or production of higher value materials from the original products increases the net revenue for this process. The first part of this paper will present the chemical recovery options for the product spectrum produced by Biocarbons Corporation's reactor system. This is followed by a general discussion and guidelines for chemical recovery from pyrolysis. This paper will also present suggestions for additional research and development to create new high value chemical products and recovery methods. BIOCARBONS CORPORATION'S CHEMICAL PRODUCTS
By operating an air-blown, fluidized-bed reactor at about 590 to 600"C, Biocarbons Corporation has been able to produce a water-insoluble liquid product that is phenolic in nature. While product yields are lower than for fast pyrolysis processes designed to maximize liquid fuel production from biomass, the selectivity is 100%; no further product separation is required before the oil can be used to make adhesive. Like other pyrolysis oil processes,Biocarbons Corporation's reactor produces a large number of oil compounds. For mixed hardwood (maple, birch and beech) pyrolyzed at typical operating conditions, 69 peaks were found by GCMS analysis. Of these, the 14 peaks present at above 2 mole percent, represented 45 mole percent of the product that came through the GC. The 27 peaks between 1 and 2 mole percent, represented an additional 37 mole percent of the product. These compounds that were identified are listed in Table 1, in order of appearance (time). Several of the #4-position groups could also be occurring at the #3 position. All are reactable to make a Dhenolf o r m a l d e h v d e a adhesive. Pyrolysis oil from pine that was made at the same operating condition (but has not yet been tested for adhesive use) had essentially the same compounds present at >1 mole percent, but at different relative concentrations. Some lower concentration compounds such as fatty acids are only produced from pine, but these compounds are specific to softwoods and the composition of softwoods. A comparison between the mixed hardwoods and pine products is shown in Table 2. Although some have three or more, most of the compounds in Table 1 have two available positions for methylene linkages, compared to three for phenol. Some have only one site, and are polymer chain breakers. Most of the aldehydes are capable of linking at positions in addition to the HC=O group. The overall "linkability" of the compounds present - an average of 2.0 positions per molecule - helps explain why complete substitution of pyrolysis oil for phenol does not produce a suitable thermoset, and why 50 percent or less phenol is still needed to provide an adequate. network of methylene linkages. Adhesive production and testing results are reported elsewhere (3,4). Operating the reactor at about 500°C increases the liquid product yield by about 10 weight percent, but changes the product composition, making it less suitable for use in an adhesive (see Table 3). Most of the compounds produced are 2,4 substituted phenols that only have one position available for polymerization.
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Table 1 Compounds Identified in Air-Blown, Bubbling Bed Pyrolysis Oil At > 1 Mole Percent In Order of Appearance From the Column (Time).
Phenol 1,2-Cyclopentanedione, 3-methyl Phenol, 2-methyl Phenol, 4-methyl Phenol, 2-methoxy Phenol, 2,4-dimethyl Phenol, 4-ethyl Phenol, 2-ethyl Phenol, 2-methoxy-4-methyl 1,ZBenzenediol Phenol, 2 4 1-methylethyl) or Phenol, 2-ethyl-4-methyl 1,2-Benzenediol, 3-methyl 1,4-Benzenedicarboxaldehyde-2-methyl
1H-Inden- 1-one, 2,3-dihydro 1,ZBenzenediol, 4-methyl 2-Methoxy-4-vinylphenol Phenol, 2-methoxy-3-(2-propenyl) 4-Ethylcatechol Vanillin Phenol, 2-methyl-6-(2-propenyl) Phenol, 4-ethyl-2-methoxy Phenol, 2-methoxy-4-(2-propenyl) 1,3-Benzenediol, 4-propyl Ethanone, 1-(4-hydroxy-3-methoxyphenol) Benzaldehyde, 4-hydroxy-3,5-dimethyl 4-Hydroxy-2-methoxycinnamaldehyde
Table 2 Comparison Between Mixed Hardwood and Pine Pyrolysis Oils Made At 590°C Using Selected Compounds.
Compound
Hardwood GC/MS Area Pine GC/MS Area
Phenol 2.18% Acetophenone 2.3 1 Phenol, 2-methyl 3.86 1,ZBenzenediol 1,2-Benzendio1-3-methyl 2.93 5.26 4-Ethylcatechol 2-Methoxy-4-vinylphenol 1.04 1.54 Vanillin n-Hexadecanoic acid Furan, 3-phenyl 0.74 4-Hydroxy-2-methoxycinnamaldehyde
1.84% 0.60 1.66 2.87 2.71 4.22 1.67 1.46 0.61
----
____
Ratio 1.18
____
1.39 1.34 1.08 1.25 0.62 1.05
____
Compounds in the water-insoluble liquid produced at around 600°C with current high value (greater than S2.00kg) and commercial uses include vanillin and catechol(l,2benzenediol, pyrocatechol). However these compounds are difficult to remove economically as discussed below and do have adhesive value. It is more feasible to remove catechol fiom the scrubber water as discussed below. To recover the liquid product from the reactor product gas stream, Biocarbons Corporation uses a venturi scrubber with a closed water recirculating loop that (for commercial operation) is kept above the water dew point of the gas to avoid net water production (An electrostatic precipitator would be used downstream of the scrubber
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Table 3 Major (>2% of Area) GCMS Compound Peaks Identified In 11/99 Mixed Hardwood Oils.
Compounds Identified At 500°C
Compounds Identified At 590°C
Phenol, 2-methoxy Phenol Phenol, 2-methyl Phenol, 2-methoxy-4-methyl Phenol, 4-methyl 2-methoxy-4-vinylphenol Phenol, 2-methoxy Phenol, 2,6-dimethoxy Phenol, 2,4-dimethyl Phenol, 2-methoxy-4-( 1-propenyl) Phenol, 2-ethyl-6-methyl Eugenol 1,2-benzenediol, 4-methyl Phenol, 2,6-dimethoxy-4-(2-propenyl) 2-Methoxy-4-vinylphenol Benzaldehyde, 4-hydroxy-3,5-dimethyl 3,5-Dimethoxy-4-hydroxycinnamaldehyde Phenol, 2,6-dimethoxy Eugenol 2-Methoxy-6-methylphenol Vanillin Phenol, 2-methoxy-4-( 1-propenyl) 5 additional peaks 3,5-Dimethoxy-4-hydroxycinnamaldehyde
to collect the remaining oil fume.). As the various compounds reach saturation in the water, they deposit into the water-insoluble oil phase that is heavier than water and separates from the water phase by gravity. To prevent water soluble acids from reaching saturation and entering the water insoluble phase (if a third phase is not formed), the scrubber water would be treated with metal hydroxide to form precipitating salts after the treatment shown in Figure 1. Compounds typically found in the 600°C reaction scrubber water at high concentrations and their relative concentrations are shown in Table 4. Of note are catechol (and levoglucosan (1,6anhydro-P-d-glucopyranose).Catecholis present in the water at a higher concentration than in the oil phase, because of a high solubility in water. The catechol can be recovered by proprietary chemical reactions used to convert commercially produced catechol into even higher value products. An alternative method is shown in Figure 1. A side stream of the scrubber water at about 40 weight percent "solids" is removed, heated and flashed to separate the higher boiling point catechol and sugars from the water and organic acids, aldehydes and ketones (an additional drying step may be needed). If not separable by melting point difference, the catechol can be selectively washed out from the sugars with methanol. The water, hydroxyacetaldehyde and formic acid are distilled overhead from the higher boiling compounds -acetic acid, hydoxyacetone, etc. The formic and other organic acids can be removed from water by adding a metallic hydroxide (sodium, calcium, copper, iron, etc.) to produce the corresponding lower solubility formate, etc. Catechol yield from dry wood seems to be in range of 0.65 to 1.35 weight percent. Levoglucosan could also be recovered from the scrubber water, and is of special note since it is a major high temperature decomposition product of cellulose for all reactor systems. Levoglucosan has the potential to be used as a precursor for a wide
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Table 4 Selected Relative Concentrations In Scrubber Water of Compounds Made At 590°C From Mixed Hardwoods.
1-Hydroxy, 2-propanone (acetol) Acetaldehyde Phenol 1,2-Benzenediol (catechol) 1,ZBenzendiol, 3-methy1 1,2-BenzendioI, 3-methoxy 1,2-Benzenediol, 4-methyl Levoglucosan
1.35 1.95
1.00 4.42 1.12 1.20 2.07 12.04
Acids Ketones, etc
T A
Water Acid
A c:gq ~
Flash Vessel Methanol Steam From
Acetic Acid HydroxyAcetone, etc
r
Figure 1 Product Separation And Catechol Recovery.
range of industrial chemicals including levoglucosenone, linear dextrans, vitamin H, polyurethanes, epoxies, surfactants and chiral ligands (8). Finally, while the char produced from the pyrolysis process can be sold for making charcoal, depending on the cost of transporting the char to a charcoal-producing facility, the char can also be used to make activated carbon. Because the char is what remains after compounds have rapidly volatilized out of the original structure, the removal of the remaining hydrogen and oxygen, and the opening of the pore structure
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is rapid and can be done below typical steam activation temperatures and residence times. The yields are approximately 50% by weight with surface areas of at least 500 m2/g and Iodine numbers over 500. The optimization of the activated carbon grades and yields is ongoing.
REVENUE BENEmTS Recovering (a 0.6 weight percent yield of) catechol for sale at about 75% ($1.85/kg) of the large volume/contract sales price, and selling charcoal at $77/tonne, F.O.B. the plant would increase the gross cash flow from a 180 dry tonne of wood per day plant making 14,750,000 kg per year of pyrolysis oil for adhesive use (at $0.077/kg) by about $1,140,000 per year or about 20%. However, the net cash flow increase, exclusive of any additional cash flow for other by-products is about $700,000 per year or about 30%. A catechol yield of 1.2 weight percent would increase the net cash flow by an additional $600,000 per year. Converting the char to activated carbon with a yield of 50% and a sales price of $O.SO/kg, F.O.B. the plant, would increase the net cash flow by about another $1,000,000.
COMMERCIAL CONSTRAINTS ON CHEMICAL RECOVERY ECONOMICS Three strong warnings need to be attached to t h s type of economic analysis. First, there is not an infiite market for any chemical product. For many of the chemicals that potentially can be produced from biomass, the markets are small relative to commodity petrochemicals. Most of the chemicals that can be produced from biomass should be considered specialty chemicals. Unless new large uses can be found for them, e.g. using levoglucosan as an intermediate building block, large scale production of any of these currently high-value chemicals will drop the price. Second, attention must be paid to the current purity and grades of product being sold, and the costs required to achieve these specifications. Even if absolute purity specifications are achieved, the types of impurities remaining will be different than those in the conventional chemical. The effects of these impurities on current uses will have to be evaluated. If the amount used in any most applications is small, this could become very expensive. It is also unlikely that small-volume users of any chemicals recovered from pyrolysis oil will be willing to change their manufacturing processes to accommodate new impurities. Finally, because of the small yields and scale relative to even many specialty petrochemicals, one should not automatically assume that the recovery "production" costs of chemicals from biomass pyrolysis are necessarily competitive with the actual production costs from p.etrochemicals; there is usually a large mark-up over costs, because the markets are small, the number of producers is small and there are technical barriers to entry, sales and application development. It would be wise to codevelop recovery with a current manufacturer, just to assure distribution channels. If recovery costs are lower than conventional production, and the market is growing, there will likely be interest, because expansion capital costs can be avoided, and incremental additional production becomes feasible.
1202
GUIDELINES FOR CHEMICALS PRODUCTION AND RECOVERY
USE PROCESS AND REEDSTOCK PRODUCTION
SELECTWITY
PROLYSIS
OIL
While all pyrolysis oil production reactor systems produce similar materials, each reactor produces a unique compound slate. The first decision, especially for a potential chemical or fuel producer, rather than a reactor developer, is to determine what products to make and which reactor system to use. The operating parameters of any reactor system designed to produce pyrolysis oil, especially temperature, can be altered to change the pyrolysis oil product composition and yield. Different feedstocks will produce different pyrolysis oil compositions and by-products, e.g. amorphous silica from rice hulls or rice straw, fatty acids from pine. Finally, feedstock pretreatment and/or catalysis, or reactor-bed catalysis can be used to improve specific product yields (7). Reactor system developers need to examine what they can produce and make this d o r m a t i o n available to chemical manufacturers and suppliers/owners of biomass feedstocks. This assumes that analysis of the entim liquid pmduct fmm thermal conversion can be made, including quantitative analysis for any compounds that am being consideted for mcovery. Physical characterization - pH, viscosity, solids content, e t c h also needed. However, what can be produced is of no value, if it cannot be recovered or used economically. This involves examining the trade-offs between yield and current commercial value, recovery costs, and potential commercial value.
SEPARATION AND RECOVERY HURDLES Recovery of specific compounds from pyrolysis oil is made difficult by the thermal sensitivity of most pyrolysis oils. The aldehydes produced can react with the phenols, furans, ketones, etc. present, forming viscous polymerized compounds that have no value as distinct chemicals (but still may be usable as adhesive feedstock). These reactions are accelerated by temperature. Many of the compounds produced are photosensitive or are oxygen sensitive, again polymerizing when exposed. Addition of bases to adjust pH by organic acid neutralization will catalyze most of these polymerizations. Divalent cations can polymerize some of the vinyl groups present from the lignin similar to styrene polymerization. Some of these reactions are useful for making adhesives, but make thermal separation for chemical product recovery difficult without using low temperature vacuum distillation. Initial separation of water-soluble material from water-insoluble material is usually easy, because most pyrolysis oils are really oil-in-water emulsions, with the organic acids and/or compounds with hydroxyl groups usually serving as the emulsification agent. The water-insoluble compounds - phenolics, etc., can be separated from the water-soluble products (organic acids, etc.), by adding water, changing the pH, heating (the riskiest) or some combination of all three, similar to water separation from crude petroleum. Solvent extraction can be used to remove groups of compounds (1) from water soluble or water-insoluble fractions, but solvent extraction is not particularly selective, and recovery of the solvent can be expensive.
1203
WH4 T CAN BE PRODUCED ECONOMICALLY While multiple products can be produced from biomass pyrolysis, pyrolysis and product recovery is not really a refinery, because thermal processing, the backbone of petroleum processing is not usually a viable option for product separation. There are also large differences in the economies of scale between a petroleum refinery and a biomass conversion plant, because of the differences in cost for transporting low density biomass by truck or rail versus transporting petroleum by ship or pipeline. The economic radius for a petroleum refinery feedstock can be worldwide, versus about 100 kilometers for a biomass plant, unless the biomass is the gathered residue from an agricultural processing plant. Storage requirements and feedstock availability are also major differences. Consequently, biomass processing plants, even paper mills utilize feedstock on the order of 1000 tomes per day or less, rather than 40,000 tomes per day or more for a economic world-scale refinery. The by-products of petroleum r e f i g that are the building blocks of the petrochemical industry are produced in larger quantities than a typical biomass-to-energy conversion facility. For example, a single fluidized catalytic cracker for upgrading heavy oil fractions can produce over 1000 tomes per day of propylene for making polypropylene. Finally, biomass and biomass compounds contain much more oxygen than petroleum products. Duplicating petrochemical building blocks from biomass can require a large amount of expensive hydrogen and processing to remove this oxygen. Consequently, making most conventional petrochemicals from thermal conversion of biomass is not going to be economic until petroleum prices actually reach the dizzy heights predicted in the early 1970's, i.e., not in your lifetime. Other than fuels, the products that make economic sense to try and produce and then use as either end products or synthesis intermediates are the chemicals that can actually be produced and recovered from biomass pyrolysis without a high recovery costs materials that can be recovered in an acceptable form for $0.20 to 0.50 per kg or less. This means using the solid by-products, organic acids, aldehydes, ketones, phenols, hrans, sugars, pentanones, etc. that are "currently" produced, trying to improve specific product yields and selectivity as much as possible. This rule of thumb might be ignorable for low concentration pharmaceutical precursors such as steroids or fragrance precursors that have been found in pyrolysis oil, depending on the value and recovery and purification costs. The pyrolysis products that can be produced fall into two categories- compounds that are or have already been used, and compounds that are newly available in bulk and at lower cost such as levoglucosan. For the former there is a wealth of literature that exists prior to the widespread of use of petrochemicals (for example reference 2) that is usually ignored by lazy investigators that end up reinventing or ignoring prior research. Much of the use of the chemicals that used to be made from biomass and coal was cast aside when cheaper petrochemical synthesis pathways became available. However, many of these forgotten applications and syntheses should be newly competitive with higher volume production and therefore lower cost biomass chemicals. This also may apply to biomass chemicals newly available in "bulk" as specialty chemicals for which bench studies were done, but whose commercial prospects were dim, because of the actual or perceived cost of making the chemicals. The new "bulk" intermediaries should also be evaluated for use via all the chemistry, especially for making polymers, that has been developed since the advent of the
-
1204
petrochemical age in the 1930's.
PRODUCT RECOVERY TECHNIQUES Product recovery is some combination of rote methods of using processing tools and inspired application of new and conventional separation tools. For example, if aldehydes make distillation infeasible because of polymerization products andlor vacuum distillation costs, then one can consider utilizing the Cannizzaro reaction to make organic acid salts and ketones from the aldehydes before proceeding. Many bench-scale processing techniques such as thin-film evaporation have been scaled up to handle larger quantities of temperature sensitive materials. Assistance in utilization of these techniques can usually be obtained from the equipment vendors. Other suggested product recovery tools that may not be readily apparent are: (1) Organic acid removal by resin separation in a chromatographic simulated moving
bed (2) Water removal by membrane separation to promote selective precipitation of lower solubility compounds, followed by solvent separation of the precipitates (3) Melting point separation of thermally stable solids (4) Adding or seeding compounds to cause precipitation of specific chemicals ( 5 ) In-situ chemical reactions such as oxygen initiated reactions, e.g., aldehydes to acids
SUGGFSTED W E A R C H A N D DEVELOPMENT In order to effectively produce chemicals from biomass pyrolysis, developers need to think beyond the mindset of maximizing production of liquids and stuffing these liquids into a combustion device. The proper thought process should be that of a chemical engineer - what can I make from this and what do I then do with it (to become rich and retire early). For a few specific chemicals thaf are identified through analysis and are easy to recover, no additional research and development are needed other than scale-up; the commercialization problems are market related. For most potential products, some additional help is needed. The following are suggested research and development topics: (1) Polymerization of the sugars such as levoglucosan into thermoplastics
(2) Utilization (beyond polymerization) of any oxirane sugars via ring opening (3) Fluidized bed catalysis for better selectivity and removal of compound oxygen (4) Conversion of "new" bulk five and six carbon sugars to furans and pyrans, respectively ( 5 ) Organic acid and other oxygenated compound separation by resins using a simulated moving bed (6) Identification of specific solvents for compound classes, i.e., aldehydes, ketoneqetc. (7) Fragrance and flavors from the aldehydes produced (8) Increased steroid precursor yields and better selectivity (9) Pyrolysis oil mixture properties and quantitative analysis techniques (10 Recovery and usage of pentanones and quinones
1205
(1 1) Increased yields of polyhydric phenols (12) For any reactor system, testing of feedstocks known to be rich in valuable,
extractable components (13) Production of silicons and silicates from amorphous silica from rice hulls and
straw (14) Increased selectivity and/or recovery of vinyl and propenyl phenols for
polymerization ,etc.
ACKNOWLEDGEMENTS The author would like to thank Carrie Davis and Lynn Gennaro of Northeastern University and Dietrich Meier of the Institute for Wood Chemistry, Hamburg, for the analytical assistance provided.
1. Chum H. et al. (1990) Process for Fractionating Fast-Pyrolysis Oils and Products Derived Therefrom. U.S.Patent 4,942.269. 2. Ellis C. (1935) The Chemistry of Synthetic Resins, Reinhold Publishing, New York. 3. Himmelblau D.A., Grozdits G. A. and Gibson M. A. (2000) Performance of Wood Composite Adhesives Made With Biomass Pyrolysis Oil, To be published as part of Wood A dhesives 2000, Forest Products Society, Madison, WI. 4. Himmelblau D.A. and Grozdits G.A. (1999) Production and Performance of Wood Composite Adhesives with Air-Blown, Fluidized-Bed Pyrolysis Oil. In Proceedings, Fourth Biomass Conference of the Americas, Elsevier, New York. 5. Himmelblau D. A. and Grozdits G.A. (1999) Production of Wood Composite Adhesives with Air-Blown, Fluidized-Bed Pyrolysis Oil. In International Contributions to Wood Adhesion Research, Forest Products Society, Madison, WI. 6. Himmelblau D.A. (1996) Phenol-Formaldehyde Resin Substitutes from Biomass Tars. In Wood Adhesives 1995, Forest Products Society, Madison, WI. 7. Piskorz J, Radlein R, et.al. (1995) Fast Pyrolysis of Pretreated Wood. In Proceedings of the Second Biomars Conference of the Americas, National
Renewable Energy Lab, Golden, CO.
-
8. Resource Transforms. Levoglucosan A Chiral Raw Material, Waterloo, Ontario.
1206
Sibunit Supported Catalysts for Hydrogenolysis of a C - 0 Bond in 'Bio-Crude-Oil' Components D.G. Aksenovl, A.N. Startsev' and B.N. Kuznetsov2 1 Boreskov Institute of Catalysis, 630090 Novosibirsk, Russia Institute of Chemistry and Chemical Technology, 660049 Krasnoyarsk, Russia
ABSTRACT: New catalysts for hydrogenolysis of C-0 bond are proposed. The catalysts were prepared by anchoring of Mo, (Ni,Mo) and (Co,Mo) complexes to the surface of a new carbon support Sibunit. Two types of the active component were prepared - oxide and sulfide forms. The catalysts were tested in a model reaction of tetrahydrofuran hydrogenolysis. As shown, the catalysts are active in the purposehl reaction of C-0- bond hydrogenolysis and do not catalyze the side reactions polymerization and dehydration of tetrahydrofuran.
INTRODUCTION Utilization of the liquid synthetic raw materials produced on thermal conversion of biomass has received much attention in recent years [ 11. "Bio-crude-oil" possesses high viscosity, contains solid resin precipitates, and a large number of oxygen containing contamination's. For this reason "bio-crude-oil" cannot be used directly as a fuel. Hydrorefining of "bio-crude-oil" allows one to decrease the concentration of oxygen to the required level, to decrease viscosity, and to remove solid resin materials. Hydrogenolysis of alcohol and ether groups proceeds easily. By contrast, hydrogenolysis of tetrahydrofuran (THF) and its homologues runs with great difficulty, though their concentration in "bio-oil" is rather high [2]. The traditional catalysts have usually a short period of operation because of high rates of deactivation processes, such as polymerization, carbonization, and decomposition of the active component. The catalysts are routinely prepared by precipitation, which results in formation of a wide range of surface species, providing a number of side reactions. The side reactions may also occur on alumina, the main support for the hydrogenolysis catalysts. Alumina is an active catalyst for dehydration processes as well; In our work, Sibunit, an artificial carbon material, was used as a support. The methods of Sibunit synthesis permit one to vary surface and size of the support pores within a wide range. Because of high thermal stability, mechanical strength, and
1207
inertness to many chemical reactions, Sibunit is very promising as a support for the hydrofining catalysts operating under severe conditions. EXPERIMENTAL The catalysts were prepared by supporting oxalate Mo and Mo-Ni (Co) complexes [3]. After drylng the catalysts were calcined at 500°C in air (CatMo, CatNiMo, and CatCoMo). Some catalysts were sulfirized with hydrogen sulfide (CatMoS, CatNiMoS, and CatCoMoS) at 400°C. The concentration of Mo in all samples was about 10%. For bimetallic catalysts, the ratio Mo :Ni(Co) is close to 2. The catalysts were tested on hydrogenolysis of tetrahydrofuran (THF). The reaction was performed in a flow setup at the hydrogen pressure of 20 atm and 250400°C. After the reactor we installed a vessel, filled with glass fiber, to trap the nonvolatile reaction products. RESULTS AND DISCUSSION Hydrogenolysis of THF may follow different schemes to produce either butane (scheme l), or butenes (scheme 2) or butadiene (scheme 4). In addition, polymerization, yielding crown ethers (scheme 3), may also occur.
C4Hg + H20 3
In contrast to alumina, Sibunit is not active on hydrogenolysis of THF (Table 1). The products of polymerization were not observed even at high temperature (Table 2). The monometallic molybdenum catalysts exhibit low activity. The oxide catalysts are more active than sulfides by a factor of 3. At high temperature, their activity is comparable (Table 1). 1208
Table 1. THF hydrogenolysis over the oxide and sulfide catalysts. Catalysts
Sibunit K M O
KMOS KNi KNiS
KCoS
Reaction temperature, "C
Activity, mol(THF)/(molCMet*h)
-
300 350 400 300 400 300 350 400 250 300 300 350 400 300 350 400
596
31,O 138 11,9 34,8 324 31,4 895 13,2 16.6 2,7
13,3 26.3
Both catalyst types preserve activity for several hours (Fig. 1). The activity and stability of the bimetallic catalysts strongly depend on their type. The oxide catalysts exhibit higher activity and stability than the sulfide catalysts. In three hours, the activity of sulfide catalysts decreases by -20%, and that of the sulfide catalysts, by a factor of 3. For bimetallic catalysts, there is the synergetic effect. At the initial reaction stage, the activity of oxide Mo-Ni catalysts is hgher than that of Mo catalysts by a factor of 5. In three hours, this value decreases to 4 (Fig. la). For the sulfide samples, thls effect is less pronounced (Fig. lb). In the next run,the catalyst activity was determined at 400"C, then the temperature was decreased to 300°C. This procedure was repeated for several times. For both oxide and sulfide molybdenum catalysts, activity and the product composition are reproduced with high accuracy in a number of experiments (Fig. 2). At high temperature, the bimetallic catalysts are quickly deactivated during the reaction (and the experimental cycles are not reproducible). Butadiene was not observed in the reaction products on all studied catalysts, including the pure support. The composition of the hydrogenolysis reaction products depends primarily on the reaction temperature (Table 2). For both oxide and sulfide molybdenum catalysts, the ratio between butane and butenes is much higher than 1. As temperature rises, thls ratio decreases for the oxide forms and increases for the sulfide forms. Polymerization begins at T > 300°C on both catalyst forms. For all bimetallic catalysts, excepting Mo-Ni oxide catalysts, the ratio butane : butenes is higher than 1 at 300°C. The amount of resin materials holds with time only for Mo catalysts. For bimetallic catalysts, the conversion of THF into resin materials increases with time by -10-15%. For all catalysts, excepting Mo-Ni sulfide catalysts, the yield of resin complexes increases with increasing temperature.
1209
molTHF/moIC Met’h 35
A
301
15
2o
-m,
Ltn\C
I
\&4----*
MoIC m
0
20
40
60
Mo-NilC
80 100 120 140 160 180 200
t, min
molTHF/molZMet*h
B
’1
Mo-Ni/C
5 1
4
3
J
i
-7
0
20
40
60
\
80 100 120 140 160 180 200
t, rnin
Fig.1 THF hydrogenolysis over the oxide (A) and sulfide (B) catalysts. Reaction temperature is 300OC. At high temperature and in the presence of Mo complexes, THF is polymerized to yield crown ethers (scheme 3) [4]. Probably, this reaction is responsible for deactivation of the bimetallic catalysts supported on Sibunit. The monometallic catalysts are not deactivated even at high yields of resin complexes and high temperature. It is likely that the polymers do not deposit on the active component and the support surface.
1210
mol(THF)/mol(Mo)*h
04OOOC 3OOOC
0
1
2
3
N2 of experiment
Fig.2 Changes in activities over monometallic catalysts at different temperatures in the consecutive cycles of the reaction temperature Table 2 Product composition of THF hydrogenolysis after 30 min of beginning of reaction. Catalysts Sibunit
Reaction temperature, "C. 300
Butanehutenes ratio
a*
-
-
4nn
For bimetallic catalysts, strong bonds between oxygen of the formed polymers and nickel are formed. According to the structure of the active component of the sulfide catalysts, electron density on the nickel atoms is reduced, which provides
121 1
formation of strong bonds between nickel and the molecules possessing lone electron pairs [ 5 ] . Because of this phenomenon, the number of active sites in the sulfide bimetallic catalysts quickly decreases, which results in a sharp decrease in the catalyst activity. In oxide catalysts these bonds are less strong. As a result, a major part of the formed polymers is removed from the reactor, and the catalyst deactivation proceeds slower, though the number of resin complexes is large (Table 2). The fact that butane is not found in the reaction products indicates that THF is not dehydrated (scheme 3) on the study catalysts and oxygen is removed from THF by hydrogenolysis (schemes 1 and 2). On the oxide and sulfide molybdenum catalysts, butane is the main reaction product, that is why hydrogenolysis predominantly follows scheme 1 . T h s is also true for bimetallic sulfide catalysts, if the reaction occurs at high temperature. For the rest catalysts, a part of butenes in the reaction products is rather high. This suggests that hydrogenolysis of the most part of THF follows scheme 2. CONCLUSIONS: 0
0
0
0
Sibunit is the most suitable support for the catalysts of hydrogenolysis of the C-0 bond in THF as a model of the “bio-crude-oilcomponent; dehydration does not occur on the study catalysts; hydrogenolysis of the C-0 bond follows scheme 1 on the sulfide and oxide molybdenum catalysts and bimetallic sulfide catalysts, and schemes 1 and 2 on the other catalysts; polymerization as a side reaction does not proceed on the monometallic catalysts at low temperature; bimetallic catalysts are quickly deactivated because of strong bonds formed between the polymers and the active component.
REFERENCES
1. Bridgwater A.V. (1994) Catalysis in thermal biomass conversion. Appl. Catal. A: General, 116,5-47. 2. Veldsink J.W., Bouma M.J., Schoon N.S.-H., Beenackers A.C.M. (1997) Heterogeneous Hydrogenation of Vegetable Oils: A Literature Review. Catal. Rev.- Sci.Eng., 39,253-318. 3. Aksenov D.G., Klimov O.V., Startsev A.N. (1997) Alumina supported sulfide catalysts. VI. Hydrogenolysis of tetrahydrothiophen.Kinet. Katal. 38, 903-7. 4. Yukelson I.I., (1969) Methods of the Basic Organic Synthesis, Nauka, Moscow. 5. Startsev A.N. (1995) The Mechanism of HDS Catalysis. Catal. Rev.-Sci. Eng., 37, 353-423.
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Multi-parameter Assessment of Sunflower Husksawdust Layer Hydraulic Resistance Gubynskyy M., Shishko Y., Usenko A., Vvedenskaya T. National Metallurgical Academy of Ukraine 4 Gagarin Avenue, Dniepropetrovsk 49005, Ukraine
ABSTRACT Modelling and computing of thermal biomass treatment processes involving thick layers requires that the gas permeability and hydraulic resistance of such layers be known. The paper presents the results of experimental research into the hydraulic resistance of the thick layer comprised of two biomass varieties: sunflower husks and sawdust. The research was conducted within the range of Re alteration from 5-70, and layer porosity change 0.2-0.7. INTRODUCTION The utilization of agricultural and wood biomass waste for power generation requires knowledge of the physical properties of the wastes. These include low filling density and significant compressibility of waste during transportation in the channel. The last factor is efficiently used in the technology of sawdust briquetting by biomass friction against the channel walls. Biomass compression diminishes unit dimensions and intensifies the process of thermal treatment, which is especially important in conditions of excess pressure. Gas permeability of the layer and its hydraulic resistance are generally considered to be the major parameters determining the thermal treatment processes in the layer. The regime of gas motion is one of the essential factors influencing the process. As the gas is moving along the channel the particles of the layer are streamlined, the laminar regime giving way to the turbulent one at R-4. In this phase the actual value of gas filtration velocity may significantly exceed the conventional gas velocity related to the channel cross-section The movement of the gas in the layer depends on dimensions, amount, and shape of the channels formed by the particles. These parameters stand in finctional relation to the granulometric composition and porosity of the layer, as well as particle dimensions and shape. Porosity is greatly determined by the particle shape. Ball shape allows for the minimum void volume in an irregular filling. However, the layer porosity characteristics obtained for the simplest case of ball-shaped evenly-sized particles cannot be applied for determining the type of the filling, or the shape of the channels formed by the particles and, consequently, for estimating the layer resistance.
1213
Literature on the issue provides a sufficient number of equations for computing the hydraulic resistance of the layer [1-6]. In processing raw data the authors departed from two theoretical approaches modelling the pattern of particle motion in a gas: - A capillary model presenting a system of channels with developed roughness; - A model of the particle ensemble, with gas streamlining particular elements (particles) of the layer. Hydraulic channel diameter d, was assumed to be a defining parameter for the first model, while equivalent particle diameter dk was chosen for the second model. The conventional equation for computing hydraulic resistance is known to be:
Where
6 I = f (Re, E) - coefficient of hydraulic resistance; Vnjh E = __ - layer porosity;
v,
wi - gas velocity related to the layer cross-section area. Experimental relations resulted in different expressions for computing Re and
4 1. The value of EJ is as a rule inversely proportional to Re first power and E first or second power, i.e.
-
1
1
The above-mentioned relations are valid only for certain shapes and dimensions of particles as well as for a limited range of porosity values. At E=0.4 the range of variation does not exceed M.08.These relations are rather generalized which tells on the calculation error and calls for fi~rtherexperimental proof of their validity. A series of experiments has been undertaken with the view to determine hydraulic resistance of sunflower husks and sawdust of coniferous wood.
RESEARCH METHOD The scheme of experimental unit is presented in Figure 1. Biomass layer under analysis is placed in the cylindrical container 200mm0. The layer is enclosed from the top and the bottom by the specially designed grates whose hydraulic resistance is extremely low in comparison with the layer hydraulic resistance, the fact that was taken into account in the processing of experimental data. The grates prevent the material from being carried away and ensure that the biomass layer be compressed and held in the experimental chamber. The layer height (250mm) remains constant in the process of the experiment as extra portions of the material are added into the cylinder. Then the layer is reduced by the hydraulic press to the initial volume. The amount of additionally-charged material was measured by weighing the material after the experiment. The absolute error in weighing did not exceed 0.5g. Porosity and density of the layer varied from 0.73 to 0.21, and from 1OOkglm3 to 400ker/m3respectively.
1214
2
Fip4r-e.I Experimental unit. 1-container with material under investigation
Z-bIower;3-U-shaped manometer;4-calibrated n o d e ; 5-chamber for the flow velocity levelling ;6-hydraulic press TB-1-25 blower blew the air through the material. The regulating valve in the blower’s inlet maintained the air consumption. The amount of the consumed air was measured by a calibrated round-edged nozzle (length-diameter ratio: Vd = 4.3) which refers to the consumption coefficient p=0.97.The nozzle was placed in the chamber for gas velocity levelling. The air pressure near the nozzle was measured by a liquid Ushaped manometer. The amount of air passing through the layer was calculated by: = 2 Q=v-d 4 Where H - air pressure near the nozzle, mm of water column.
@
The difference of pressure in the layer was measured by the manometer: the absolute error being 0.5mm of water column. ANALYSIS OF EXPERIMENTAL RESULTS
The results of exqxnmental research into husks and sawdust hydraulic resistance as related to air velocity and layer density are presented in Figure 2 and Figure 3.
1215
AP. Pa 10000
8000 6000
4000 2000
0
Figure 2. Experimental data on the sunflower husks layer hydraulic ra&ttnncc!
Figure 3 Experimental data on the sawdust layer hydraulic resistance
1216
Conventional empirical equations for computing hydraulic resistance in the thick layer [1,2] were used for evaluating pressure loss AP. Computationswere based on the actual geometric characteristics of the particles and the data received fiom studying the layer granulometric composition as obtained by direct measurements [7] alongside with the sieve analysis of sawdust. The comparison of the computed and experimental values (Figure 4 ) testify that the obtained relations ensure the required accuracy of calculations only within the range of high biomass porosity (without external layer compression). The value of error grows significantly with the increase of the air velocity and particles packing density. That is why it seems incorrect to use conventional computational instruments in the given situation.
0,4
0,3
0,s
0,7
0,s
E
1-experimental values; 2,3 -computed values [1,2]. F'pre 1.Pressure losses in sunflower husk layer at air velocity w=0,2d s Depending on the type of relations between resistance coefficient and Re the value of 5 can be computed by the power equation: <=a*Reb, where coefficients a and b are functions of the layer porosity. Computational equations for coefficients a and b are presented in Table 1. Table I.Computed equations for calculating empirical coefficients a and b
Biomass type Sunflower husks Sawdust
I Coefficient a
I
Coefficient b
I a =105(-9,69E3+19.34E2 - - I b = 1,48E2- 1,ME + 0,80 12,97E + 2,94) a = 4-lO6EXP(-12,7E)
1217
b = -7,56E3 + I 0,69E2-
- 3.88E - 0.44
CONCLUSIONS
The empirical equations obtained allow to evaluate biomass layer hydraulic resistance with an error up to f 10%. The resulting computational instruments will be hrther applied for fixed bed pyrolysis units designing as well as for constructing chargers providing continuous feeding of material and pyrolizer’s hermetic sealing by creating the necessary hydraulic resistance of the biomass layer.
REFERENCES 1.
2.
3.
4.
5.
6. 7.
Aetov M.E.( 1998) Hydrindic nnd thermal j d a m e ~ r t a l .qf~ the apparatus operution for Bxed ardfluidized berls./Russianl, Moscow, Chymyia Publishing House. Kutateladze S.S. (1995) Reference manuai on heat exchcvge. /Russian/, Moscow, Energia. Sethonenko B.D.. Orlik V.N., Alekseenko V.V.(1998) Matematicheskov mc&lirovmie teploperedachi i gidruvlicheskogo soprotivlenicr v ~rgencrali~~non~ teplwbmennike s sharovoi nusadkoi Russian/, Ekolotechnologei i rcsursosberegenie ,#5. Telegin A.S. ( 1 993) Tepfotechnicbeskrerascheti irtetdlttrgicheskichpechei /Russian/.. Moscow, Metallurgia Agafonova M. E., Toritsin L. N., Lekomtseva E. D. (1990) Gidravlichiskor soprotivlenie sloewi t w d k i ~jpduchona~e~atelia i MetaSlurgicheskaia ntpljtechnika./ /Russian/,. Moscow ,Metallurgia Kiet F., BIek Y.(1993)Osnovi teploperedachi. &ussian/,. Moscow, Mir. Gubinsky M.(1999) Biomass pyrolysis of organic waste in a rarefied layer. Proceedings of the YhBiomass Cotlference of Americas, Oakland, CA.
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Organic composition of liquidized model kitchen garbage S. Inoue, T. Minowa, S. Sawayama and T. Ogi National Institiite.for Resources and Environment, I6-3 Onopwa, Tsrikirba, Ibaraki 305-8549, Japan
ABSTRACT: Model kitchen garbage (87.1% moisture content, 11.3% organic content) was liquidized at 175 "C for 111. The liquidized garbage was filtered into a liquid phase and a solid phase and then the organic compositions of the liquidized garbage and its centrifuged phases were analyzed and compared to mechanically disrupted garbage. The total carbon amounts in 100 g of the mechanically disrupted garbage and liquidized garbage in the liquid phase were 2 168.1 and 2297.0 mg, respectively. The total carbon amount was almost similar value. The solubilization of solid andor insoluble material was not caused through the liquidization process. The liquidization phenomenon is mainly the thermal rupture of the cells. During liquidization, low molecular organic acids were produced and their amounts increased compared with that in the mechanically disrupted garbage. The increment of the low molecular organic acids might be due to the degradation of the proteins and sugars. The total sugar content in the liquidized garbage. including simple sugars, oligosaccharides. polysaccharides. and their derivatives. was lower than that in the mechanically disrupted garbage. Considering that low molecular organic acids can be easily decomposed by anaerobic digestion. these results show that the anaerobic digestion of the liquid phase could contribute to a higher digestion rate compared with that of the mechanically disrupted garbage.
Introduction
The kitchen garbage can be converted from a solid to a liquid slurry by thennochemical liquidization at 150-250 "C.[ 11 The liquidization of kitchen garbage has wide range of applications. such as pretreatment for transportation using pipelines and anaerobic digestion. By applying the liquidization process. we examined the anaerobic digestion of model kitchen garbage and the filtrate of the liquidized model kitchen garbage. The anaerobic treatment of the filtrate indicated fast digestion compared with the mechanically broken up original model kitchen garbage.[2] Upflow anaerobic sludge blanket (UASB) systems have been widely used for the high-
1219
rate treatment of various organic wastewaters since UASB reactors have exhbited superior performance compared to other types of anaerobic reactors when utilized with high loading rates. By using this system. the filtrate of the liquidized model kitchen garbage was continuously digested for 82 days.[3] The anaerobic digestion of a liquid phase of the liquidized model kitchen garbage exhibits fast digestion rates. An analysis of the organic constitute of the liquidized food waste during the liquidization process is important for providing preliminary information and optimum conditions about the potential for the pretreatment of the anaerobic digestion.
Experimental Model kitchen gurhuge
Garbage discharged from homes or food industries has various properties. depending on the season or region. In this work. model garbage[4-61 was used as a starting material. The components of the model garbage were sliced cabbage (92.4%). boiled rice (5.3%), dried fish (0.6%). shell (1.1?40) and butter (0.6%). Table 1 shows the properties of model kitchen garbage. The moisture content was determined by weighing the garbage after drying at 105 "C for 24 h. The ash content was determined by weighing the residue after heating at 600 "C for 1 11. and the organic content was obtained by difference. Table I Properties of model kitchen garbage
Model garbage (%) Moisture content
87.1
Organic
11.3
Ash
1.6
Mechunicul[v disrupted model kitchen gurbuge
Mechanically disrupted model kitchen garbage was obtained by the addition of the Table 2 Properties of mechanically disrupted garbage and its filtered phases
Mechanically disrupted garbage (%)
Liquid phase (%)
Solid phase (%)
Moisture content
93.8
96.5
83.6
Organic
5.9
3.4
15.7
Ash
0.3
0.1
0.7
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same weight of distilled water to the model food waste followed by disruption using a mixer for 1 min.[2] The product was filtered into a liquid phase (81.0 wt%) and a solid phase (19.0 wt%) using a 20 mm pore size nylon mesh. Table 2 shows the properties of mechanically disrupted garbage and its filtered phases. Liquidizution of model kitchen gurbugr
The model kitchen garbage was placed into a I L autoclave made of stainless steel. After purging with nitrogen gas. it was charged to 2 MPa. The reaction was started by heating the autoclave with an electric furnace. After heating the autoclave to 175 'C. the temperature was maintained for 1 11 and then the autoclave was cooled to room temperature. Separation of the liquidized model kitchen garbage into solid and liquid phases was conducted by filtration using a 20 ~i m pore size nylon mesh. Table 3 shows the properties of liquidized model kitchen garbage and its filtered phases. Table 3 Properties of liquidized model kitchen garbage and its filtered phases Liquidized garbage (%)
Liquid phase (%)
Solid phase (%)
Moisture content
89.7
93.2
74.0
Organic
9.7
6.0
25.1
Ash
0.6
0.8
0.9
Analysis ofwuter solubles in the liquidphuses
The carbon amount in the liquid phase was measured using a total organic carbon (TOC) meter (Shimazu. TOC-5000A). Organic acids were analyzed with a Dionex DX-500 ion chromatograph. The total sugar content was measured by the method of Dubois et a/..[7] Sugars were analyzed by using a high performance liquid chromatography (HPLC) (Shimazu. LC- 10AT).
Results and Discussion Liquidizution
The TOC in the liquid phase of mechanically disrupted garbage and liquidized garbage in the liquid phase were 12803 and 31479 mg/l. respectively. The total organic carbon (TOC) for the liquid garbage was 2.5 times as that for the mechanically disrupted garbage. However. the addition of water of same weight as the model garbage was needed to obtained the mechanically disrupted garbage. The total carbon amount in the liquid phase of mechanically disrupted model garbage and liquidized garbage on the basis of 100 g-model garbage were 2168.1 and 2297.0 mg. respectively (shown in Table 4). The water soluble organic carbon amount was similar value in spite of the liquidization; the solubilization of solid andor insoluble material did not occur by the liquidization process. Minowa et at. reported that 1221
liquidization phenomenon of garbage is due to the thermal rupture of cells: thus releasing intra-cell water from the cells.[8] Table 4 TOC in the liquid phases (in 100 g-model garbage)
Oreanics
Model garbage (liquid phase)
Liquidized garbage (liquid phase)
TOC
2168.1 me-C
2297.0 mg -C
Orgunic ucids in the liquidphuses
Figure 1 shows the concentration of organic acids in the liquid phases as measured by ion chromatography. The low molecular organic acids concentration in the liquidized model kitchen garbage is higher than that in the mechanically disrupted garbage. It has been reported that protein denatures at 150 OC[9] and proteins in sewage sludge hydrolyze during the liquidization process at 175 OC.[lO] Krochta et af. reported the thermochemical degradation of cellulose into organic acids.[ 111 Therefore, organic acids could be obtained by the hydrolysis of protein and the thermochemical degradation of sugar. The major phenomenon of liquidization is rupture of cells. The decomposition of soluble material might have occurred by minor reaction, such as hydrolysis.
I
0
Mechanically disrupted m d e l kitchen garbage (liquid phase)
Liquidized model kitchen g a h a g e (liquid phase)
I
20
15
10
5
0 I
Formic acid
Acetic acid
Propionic acid
Lactic acid
Malic acid
Figure I Organic acids concentration in the liquid phases (in 100 g-model garbage)
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Sugar in the liquidphases Table 5 shows the total sugar content in the liquid phases. The total sugar amount in the liquid phase for liquidized model kitchen garbage, includmg simple sugars, oligosaccharides, polysaccharides, and their derivatives, is lower than that for the mechanically disrupted garbage. The results of the HPLC showed that the fructose and glucose contents decreased during the process. (shown to Figure 2) They were decomposed to other compounds, such as sugar derivatives and acids. The decrease of sugar amount consists with the increase of organic acids amount as mentioned above. Table 5 Total sugar content in the liquid phases (in 100 g-model garbage)
Organics
Model garbage (liquid phase)
Liquidized garbage (liquid phase)
Sugar amount
89.9 mg
59.9 mg
I
Mechanically disrupted model kitchen garbage (liquid phase)
Liquidized model kitchen garbage (liquid phase)
15
10
5
Figure 2 Sugar concentration in the liquid phases (in 100 g-model garbage)
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Applicution of liquidized gurbuge
As mentioned above. additional water is needed for mechanical disrupting. One of the advantages of the liquidization is that the model garbage is changed to liquid slurry without an increase of volume. The liquid phase of liquidized model kitchen garbage contained mainly organic acids and sugars and their derivatives. In our previous study. anaerobic digestion of liquid phases obtained from liquidization of the model kitchen garbage was reported. The biogas yield from the liquid phase obtained from the liquidized model kitchen garbage was higher than that froin the mechanically disrupted garbage.[2] Furthermore. the liquid phase was continuously anaerobically digested using a UASB method.[3] The decomposition and solubilizatioii of solid organic materials is one of the key steps for anaerobic digestion. These results suggest that the anaerobic digestion of liquid phase could contribute to a higher digestion rate compared with that of inechanically disrupted garbage.
Conclusion Model kitchen garbage was liquidized by a thermocheinical process. The organic composition of liquidized model kitchen garbage and its separated phases were analyzed. The following conclusions were obtained: (1) The liquidization phenomenon is due to the thermal rupture of cells: the intra-cell water could then be released froin the cell. (2) The low inolecular organic acids concentration in the liquidized model kitchen garbage is higher than that in the mechanically disrupted garbage. The increment of the low molecular organic acids might be due to the degradation of the proteins and sugars. (3) The sugar concentration in the liquidized model kitchen garbage is lower than that in the inechanically disrupted garbage.
Acknowledgements We thank Ms. Yukiko Fukuda, an assistant staff member. for her help during the experiments.
References 1. Minowa T., Murakami M., Dote Y.. Ogi T. and Yokoyaina S. (1995) Oil production from garbage by therinochemical liquefaction. Bioniass Bioenergy. 8
(2) 117-120 2. Sawayama S.. Inoue S.. Minowa T., Tsukahara K. and Ogi T. (1997) Thermochemical liquidization and anaerobic treatment of kitchen garbage. J. Ferment. Bioeng.. 83 ( 5 ) 451-455
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3. Tsukaliara K.. Yagisliita T.. Ogi T. and Sawayama S. (1999) Treatment of liquid fraction separated from liquidized food waste in an upflow anaerobic sludge blanket reactor. J. Biosci. Bioeng.. 87. (4) 554-556 4. Koyaina K. Tanaka N. and Iiioue Y. (1983) Study on anaerobic decomposition of municipal solid waste. Proc. Environ. Snni. Engng Rex 19. 136-145 5. Taiiaka E.. and Morisliita K.( 1985) Citizens and waste disposal. Eriviron. Rex Quart. 56. 79-88 6. Tanigawa N.. Wakabayashi M. and Noguchi T. (1989) 1988 FY research of waste composition. Report of TokvoMetropolitan Cleansing Laboratory. pp. 1-42 7. Dubois M.. Gilles K. A,. Hamilton J. K.. Rebers P. A. and Smith F. (1956) Colormetric method for determination of sugars and related substrances. .4i?al. Cliem.. 28. (3). 350-356 8. Minowa T.. Dote Y.. Sawayaina S.. Yokoyaina S. and Murakami M. (1995) Pliase cliaiiging of garbage from solid to liquid slurq by thermal liquidization. J. “hem. Etig. of Japan. 28 (6) 727-73 1 9. Dote Y.. Inoue S.. Ogi T. and Yokoyaina S. (1996) Studies on tlie direct liquefaction of protein-contained biomass: The distribution of nitrogen in the products. Bioniass Bioenergv. 11 (6) 49 1-498 10. Iiioue S.. Sawayaina S.. Ogi T. and Yokoyaina S. (1996) Organic coinpositioii of liquidized sewage sludge. Bio/wa,s.sBioenergv. 10 (1) 37-40 11. Kroclita J. M.. Hudson J. S. and Drake C. W. (1984) Alkaline therinocheinical degradation of cellulose to organic acids. Biotechnol. Bioeng. S)mp., 14, 3 7-54
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The volatility of tars from pyrolysis of biomass materials V. Oja’ and M. R. Hajaligol Research Center, Philip Morris USA,Richmond, Virginia, USA. I Current Address: Chemical Engineering Department, Tallinn Technical University, Tallinn, Estonia.
ABSTRACT: The volatility of tars from pyrolysis of biomass and biomass-derived materials is of special interest. Thls interest is related to the question whether or not biomass tar evaporation has a significant effect on the total rate escape of tar from pyrolyzing substance and therefore on pyrolysis lunetics. In fact, in many practical applications (especially in fossil fuel thermal conversion processes), tar vaporization is known to be an important step during pyrolysis that influences both yield and composition of pyrolysis products. As part of the answer lies in ability to describe the volatility of biomass tars, the work was primarly undertaken to get some insights into this area. Due to the very complex nature of pyrolysis tars, there are not suitable estimation methods available experimental data are needed for further development of these techques. The experimental data are needed to increase our understanding of this phenomenon. In this paper, we present results from tobacco, hemicellulose and lignin tars. Our preliminary results indicate that the heats of vaporization of the “low volatility” fractions studied here fall somewhere between 110 and 140 kJ/mol. Therefore, incorporation of the heats of vaporization into the global activation energy should be relevant for an accurate assessment of pyrolysis kinetics.
INTRODUCTION Understanding the volatility of tars from pyrolysis of biomass materials should be an important aspect in the design of bimass thermal conversion processes, including gasification and combustion. So far, the effect of external transport limitation due to tar vaporization has been quantitatively described for coal pyrolysis and this transport limitation has been shown to cause changes in pyrolysis product yields, compositions and even in pyrolysis lunetics. Coal (as well as biomass) is a macromolecular network in which thermal degradation in a non-oxidizing environment includes, among other processes, formation of volatile (semi-volatile) fragments or species. Smaller and higher volatile fragments vaporize, bigger and less volatile tend to form char as a result of longer residence times. It could be said that there is a competition between vaporization and char forming reactions. In some cases (especially in softening coals)
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the impact on product yield and kinetics due to vaporization is more significant than in others (non-softening coals), but qualitatively the external mass transport limitation is correct for all fossil fuels including lignites [Suuberg et al., 19781. Therefore, advanced pyrolysis modelds originally developed for coal pyrolysis [Serio et al., 1987; Niksa and Kerstein, 1987; Grant et al., 19891 include tar vaporization as one relevant step in the modeling schemes and use vapor pressure correlations to predict volatilities of pyrolysis tars. Recently, some of these advanced models have been applied to pyrolysis of biomass materials [Serio et al., 1994; Chen et al., 19981. One modeling code to be unproved for the advanced pyrolysis models is the ability to predict tar volatilities. The complex nature of pyrolysis tars makes estimation of the effects of tar volatility unpredictable without experimental evidence for each individual biomass tar. There is evidence that external transport limitations due tar vaporization contribute to the overall kinetics of bimass pyrolysis. Suuberg et al. [1996, 19971 first addressed this problem by measuring cellulose tar volatility. They hypothesized that tar escape via vaporization was an important parameter in cellulose pyrolysis and therefore in determining the global mass loss kinetics at high heating rates. However, in t h s respect, it should be remembered that opposite to the other type of biomasses, cellulose tars, depending on the pyrolysis conditions, can contain a significant amount of low volatility levoglucosan and its condensation products (up to 60 %). In present context it may be stated that the volatile species from biomass pyrolysis have much smaller molecular weights and therefore are much more volatile than coal tars. On the other hand, biomass tars generally have much higher heteroatom content, which is known to lower the volatility significantly, especially when hydrogen bonds forming functional groups are present. The subject of tar vaporization in biomass pyrolysis should be given more attention. EXPERIMENTAL SAMPLES
The tars studied were generated from lignin, xylan, and tobacco starting materials. These afforded representative samples of a wide variety of tar chemical and physical properties including a cellulose-like material, hemicellulose, a sample with a high phenolic content (lignin) and a material with a large nitrogen content - tobacco. Tobacco used in this study was a low alkaloid tobacco (to minimize the contribution of the nicotine to tar volatility). Birch wood xylan is representative of hemicellulose, and was purchased from Fluka. Alkali lignin was obtained from Aldrich. TAR PREPARATION
This study was primarily focused on the volatility of tars that were prepared under conditions that minimize the extent of secondary reactions. Tars were produced using a fast pyrolysis techmque (heating rate greater than 5 "Chec). About 2 or 3 grams of sample on a wire mesh holder was pushed into the pre-heated tubular pyrolysis reactor. The sample temperature was measured by a thermocouple and the holder was pulled out from the heated region after the sample reached desired temperature. The final temperature was selected based on thermogravimetric experiments at heating rate of 20 "C /min for each material studied and are shown in Table 1. The inert gas (N2) flow was 300 rdmin at room temperature. That corresponds to the gas phase tar maximum
1227
residence time of 3 seconds in the hot zone of the reactor. This approach was selected in order to produce enough material to analyze, but at the same time to minimize the extent of secondary reactions. Volatile room temperature condensibles pyrolysis products were collected at cold parts within the reactor and in a collecting system made of a liquid trap followed by a Cambridge filter. The liquid trap contained inhibitor free tetrahydrohan (THF) at O’C. The filter and the cold parts of reactor were washed with inhibitor free THF. Dry tars were prepared by evaporating the solvent in a vacuum oven at about 45°C. DEFINITION OF TAR
The definition of tar is arbitrary and operationally defined. The tar preparation technique used resulted in loss of light materials (low molecular weight) or the “high volatility” fraction of room temperature condensable organic pyrolysis products -- the tokl tar. The material remaining, or “low volatility fraction”, was defined here as the tar to be analyzed. The relative amount of “low volatility’’ and “high volatility” fractions in the total tar depends mostly on the starting materials and pyrolysis conditions. For example, in coal pyrolysis, the “high volatility” fraction is no more than 5-10 weight percent of the total tar, but in the case of biomass, it can contribute up to 80% by weight. Table 1 shows the yields of pyrolysis products based on dry and ash free starting material. Table 1. Yields of biomass pyrolysis products (dry and ash free bases).
To get a better estimate of the amount of “low volatility fraction” in total room temperature condensable products, crude semi-quantification of evolved gases (only CO, C02, H20) was camed out in a thermogravimetric analyzer (TGA) coupled with a mass spectrometry. This estimation was based on mass spectrometric data of pyrolysis volatile products at a heating rate of 20 “C/min. We are aware that the gas yield during pyrolysis varies somewhat with heating rates, but here our aim is to obtain an approximation. Based on our estimations the amount of gases (CO, C02 and HzO)were 40% for tobacco, 25% for lignin and 10% for xylan of the total weight. This suggests that the amount of “low volatility” tar in the total tar would be about 30% for tobacco, 55% for lignin and 20% for xylan by weight. It should be mentioned that a sharp borderline between “low volatility’’ and “high volatility” fractions does not exist. There is a continuous change that is sensitive to temperature and vacuum fluctuations/uncertainties in a vacuum oven. Therefore, based on our observations, the amount of “low volatile” fractions might be up to 15 to 20% greater than the values shown if conditions were such that the maximum amount of “low volatility” tars was produced.
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RESULTS AND DISCUSSION Biomass tars are known to be very complex mixtures which contain significant amounts of polar compounds that can potentially form hydroxyl bonds. This means that current vapor pressure estimation techques are not only difficult to apply, but are inaccurate for mixtures with significant oxygen fimctionalities. Therefore we have chosen to approach this problem from an experimental standpoint. BIOMASS TAR VOLATILITY
It is known that volatility (vapor pressure) of tars depends to a great extent on molecular weight (on sum of van der Waals forces). Therefore, one of the objectives of our volatility concept was the volatility relation to the molecular weight distribution of the volatilizing material. Field Ionazation Mass Spectrometric technique was adopted for thls purpose. This technique is especially usehl because in addition to volatility data, it provides molecular weight distribution information. As our main interest was vaporization data, the prepared tars were reheated in the Field Ionization Mass Spectrometry (FIMS) lnlet probe at a heating rate of 3"C/min. T h s differs somewhat from traditional molecular weight distribution determination experiments by FIMS, where samples usually are pyrolysed within the inlet of the mass spectrometer. Figure 1 shows vaporization behavior for tars at pressure lo4 torr at a heating rate of 3 "Chin. These are relative vaporization data that relate volatilities of tars to the number average molecular weights under these specific experimental conditions.
400 8
6300
I 0
! m E g 100 E
$
0
8
J81. a
*HPtobac~ot~
b
8
c,
tobacco tar
0
'
'xylantar x
A'
hgnlntar
0 -100
Fig 1. Temperatures of volatility as a function of molecular weight.
Some general outcomes can be drawn from the data:
(1) First, volatility depends to a significant extent upon the molecuiar weight of the tar species. As a frst approximation, it can be assumed that all compounds with the same molecular weight generate similar vapor pressures. From this it fallows that for any tar, it is reasonable to say that an average molecule with higher molecular
1229
weight has lower volatility, that is vapor pressure decreases as molecular weight increases. (2) Second, the overall tar vaporization processes (positions of the vaporization curves in Fig. I) is determined by the chemical composition of the tar. The comparison between tobacco tar (solid squares) and its hexane phase partitioned fraction (referred as “HP tobacco tar”; solid circles) indicates clearly this important effect of chemical composition and functional groups. Note that the hexane phase partitioned fraction represents the less polar fraction of tobacco tar obtained by partitioning tar species from methanol solution into a hexane phase. These data indicate that one cannot describe volatility without obtaining information on molecular weight distribution and chemical composition of the vaporizing species. In order to more conveniently interpret the data in the Fig. 1, following are some characteristics of tars studied. Molecular weight distributions fitted form FIMS data on the “low volatility” fraction of tars are shown in Fig. 2.
I
h
800 M l ~ ~ ~ [ d a I t o ~
0
200
600
400
Fig 2. Molecular weight distributions fitted fiom FIMS data. Thn figure shows that all three tars contain a quite broad range of pyrolysis products with very different molecular weights and the distribution curve shape and broadness varies from tar to tar. At this point it is worth noting that presenting the FIMS data as a molecular weight distribution is only a rough approximation. Although practically useful, the assumption of continuous spectra of compounds may not be adequate for hemicellulose and lignin tars - these tars contain strong signals of individual compounds. Table 2 shows elemental composition for each tar studied here.
Table 2. Elemental analysis results for tars studied (drybases). tar tobacco lignin xvlan
M w c 219 246
71.71 75.93
207
69.99
~
H
O
N-
7.03 6.34 6.56
11.3 15.36
9.95 0.78
23.45
0
1230
~~~~
S
WC
O/C
N/C
1.59
1.18 1.00
0.12 0.15 0.2s
0.119 0.009
1.12
The oxygen content indicates that these tars are quite polar in nature. For example, an average molecule of tobacco tar would have 3, lignin tar 2.5 and xylan tar 3 polar atoms, which are likely to form hydrogen bonds. WC ratios between 1 and 1.2 indicate cyclic and saturated or aromatic structures. This might suggest that some secondary reactions had occurred despite our efforts to avoid these. Therefore, the work currently in progress is expected to overcome this problem by using a wire mesh reactor (also referred as a heated grid reactor) instead of the previously used tubular reactor. HEATS OF VAPORIZATIONAND BOILING POINT ANALYSIS The non-isothermal Knudsen effusion technique was used to study vapor pressures of tars. The experimental details have been described previously [Oja and Suuberg 1997, 19981. About 10 mg of dry tar was placed into a hermetic e f h i o n cell with a small orifice, from which the saturated vapor effises out into the vacuum outside the cell. To overcome effects caused by changes in tar composition during effiion, the experiment involved first a continuous cool-down followed by a continuous heat-up of the sample. From the mass loss data (by taking into account both cool-down and heat-up as a whole cycle) the vapor pressure was calculated using the Knudsen equation. The goal was to determine vapor pressures of tars cycle-by-cycle, where between each cycle, 'certain amounts of higher volatility species were evaporated. As the effision technique built in tlus laboratory was best suited for measuring vapor pressures fiom 10" to 10;' torr, the temperature had to be continuously increased as more and more volatile species were lost in the process. The process is fixther complicated by the fact that pyrolysis tars generally have a tendency to "age" with time, especially, at higher temperatures. It is known that condensatiodpolymerization type reactions become considerably facile at temperatures above 100 "C. Therefore, the main problem related to the experimental determination of vapor pressure was the thermal instability of tars. Significant signs of thermal instabilities for all three tars were observed around 140 -150 "C. Because of the tar thermal instabilities, we present here only results from vapor pressure measurements performed below 110°C and between 110 and 140 OC. Table 3 shows heats of vaporization and boiling points determined from vapor pressure measurements below 110 "C. It is reasonable to assume that tars under these conditions can be considered "fresh" i.e. no significant condensation type reactions had taken place. Table 3. Results fiom vapor pressure measurements below 110 "C. sample tobacco lignin xylan
Mass loss [wt.%]
Heat of vaporization [kJ/mol]
6% 6% 4%
132 119 121
boiling point at 760 torr ["C] 234 292 307
boiling point at 10 torr C"C] 170
-
210 220
boiling poipt was estimated assuming that the heat of vaporization is constant. Table 4 shows heats of vaporization and estimated boiling point values from vapor pressure measurements between 110 and 140 "C.
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Table 4. Results from vapor pressure measurements between 110 and 140°C.
sample
Mass loss [wt%]
tobacco lignin xvlan
3AOL 24% 19% 1J-70 15%
~~~
I
Heat of vaporization [kJ/mol] 1 ?a 129 138 1 11 111
I
boiling point at 760 torr [“C] 1AC 345 335 420 4LU
Boiling point at 10 torr [“C] 3-m 250 250 I
290 LYU
I
As evident from above data, the boiling points increase as molecular weight increases as a result of losing higher volatility spieces. In order to more conveniently interpret the data shown above, Table 5 shows pyrolysis decomposition temperatures and differential mass loss peak temperatures for the materials studied. The temperatures were obtained by thermo-gravimetricanalysis at low heating rates.
sample tobacco lignin xylan
Start of weight loss [“C] 150 200 190
Major peaks [“C] 275 and 315 335 240 and 290
SUMMARY The situation regarding the biomass tar vaporization is more complex, and the work above was an attempt to bring some insights into biomass tar volatility. We believe that the question of biomass tar volatility should be given more attention in the discussion of yields of biomass pyrolysis products and even pyrolysis kinetics. More specific analysis of tars and correlations of vapor pressure data as a h c t i o n of molecular weight or/and some chemical characteristics were beyond the scope of this presentation, although these are important considerations. Detailed pyrolysis models [Chen et al., 1998; Serio et al., 19941 include a tar vaporization step and use
1232
correlations for predictions of vapor pressures of primary pyrolysis tars. Unfortunately, very few experimental data to develop such correlations are available. REFERENCES
Chen, Y.,Charpenay, S., Jensen, A., Wotjowicz, M.A., and Serio, M.A. (1998) 2Th Symposium (Int.) on Combustion, p. 1327, The Combustion Institute. Grant D.M., Pugmire, R.J., Fletcher, T.H., and Kerstein, A.R. (1989) Energy Fuels 3, 175. Milosavljevic, I. and Suuberg, E.M. (1995) Ind. Eng. Chem. Res. 34, 1081. Niksa, S. and Kerstein, A.R. (1987) Fuel 66, 1389. Oja, V. and Suuberg, E.M.(1997) Anal. Chem. 69,4619. Oja, V. and Suuberg, E.M.(1998) Energy and Fuels 12, 1313. Serio, M.A., Hamblen, D.G., Markham, J.R., and Solomon, P.R. (1987) Energy Fuels 1, 138. Serio, M.A., Charpenay, S., Bassilakis, R., and Solomon, P.R. (1994) J. Biomass Bioenergy 7, 107. Suubereg, E.M., Milosavljevic, I., and Oja, V. (1996) 26& Symposium (Int.) on Combustion, p. 1515, The Combustion Institute. Suuberg, E.M., Peters, W.A., and Howard, J.B. (1978) Znd. Eng, Chem. Process Des. Dev. 17,37.
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Release of Chlorine From Biomass and Model Compounds at Pyrolysis and Gasification Conditions B. Stromberg and F.Zint1, TPS Termisku Processer AB, Studsvik 461 1 82 Nykoping; Sweden
ABSTRACT Experiments with KC1-pretreatedbiomass and different organic model compounds have shown that they release chlorine when they are heated to moderate temperature levels (300-500°C). The experiments included both pyrolysis and gasification conditions. Both low-molecular model compounds and macromolecular compounds were included, of both biological and technical origin. The results allowed a ranking of the functional groups concerning their contribution to the release of C1 from the materials. The following functional groups contribute significantly to the release of chlorine when heating the samples to 300 "C: At pyrolysis conditions: Free carboxylic acid 7 ester in lactone ring z hemiacetal > acetal z open ester a hydroxymethyl on ring or chain >> terminal hydroxymethyl At gasification conditions: Hemiacetal > free carboxylic acid 3 ester in lactone ring z acetal > open ester. It must be noted that all common biomasses contain biopolymers (carbohydrates or others) which are rich in these hctional groups. Some of them as for example crops which have been fertilised with KCl, can contain elevated amounts of chlorine. The release of significant parts of this chlorine from chlorine containing fuels can probably not be suppressed by pre-treatment under pyrolysis or gasification conditions at low temperature.
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INTRODUCTION Emission of hydrogen chloride is the third most important contribution to the global acidification from human activities. The two first are SO2 and NO,. The HCI is a local pollutant, contrary to the other two. It is soluble in water and easily dissolved in rain droplets and, therefore, usually falls down near the emission source. The hydrogen chlorine emissions from combustion and gasification processes has been calculated to 3.5 Mt./year. The major part of the estimated global contribution of HCI to the atmosphere is evaporation from the oceans. Even with a redeponation of 90-% HCI to the oceans, the estimated emission will reach approximately 120 Mt./year. The majority of the emitted chlorine from a combustion process will leave as HCI in the gas phase which may cause problems like corrosion and formation of dioxins. Both the amount and the origin of the chlorine varies between different types of fuels. In coal, the concentration varies normally between 50-2000 m a g . The origin is mainly groundwater which has been incorporated into the coal after its formation. In biomass the chlorine content can vary from less than 100 mg/kg up to 7000 mg/kg [ 13. Biomass fuels show comparably low ash content. The uncombustible fuel components which form the ash are dominated by compounds of alkaline and alkaline earth metals (K,Na,Ca,Mg) and non-metallic anion-forming elements (CI,S,P,Si). Woody biomass is low in ash content, sulphur, chlorine and silica, and the ash itself is dominated by compounds of Ca and K. Biomass fuels f?om crops are higher in ash content, sulphur, chlorine and silica. The ash is dominated by oxides of Ca, K and Si. The fuels based on the growing parts of trees (bark+debries) show ash compositions between the two extremes, wood and acre crops. Therefore the largest chlorine emissions are expected from energy crops (grasses) and from agricultural by products (straw etc). Little is known about the bonding and chemical forms of the ash-constituting elements in the plant materials. The alkaline metal ions can easily be leached from the fuels and are probably present as inorganic or organic salts, or associated to parts of the organic macromolecular structures. A first hypothesis in this project has been that the main part of the chlorine in the biomass fuels is present as simple chloride ions, forming some type of salts with inorganic or organic cations. The high degree of water-leachability of CI from the acre crops suggests this type of weak bonding in water-soluble salts. Since K is the ash constituent with the highest affinity to chlorine, easily forming KCI, it was supposed that CI was present in the fuels mainly as KCI, or that this salt was formed by reaction after degradation of other forms of potassium and chlorine when the fuel was burnt or pyrolysed. It would then be expected that the release of chlorine or chlorine compounds (even HCI) to the gaseous phase during thermal treatment of biomass would be limited by the equilibrium vapour pressure of the potassium chloride. Very little volatilisation of CI was therefore expected for temperature lower than the melting point of KCI (770°C), since the vapour pressure of solid KCI is negligible. A number of chlorine-rich biomass fuels were studied in small-scale laboratory pyrolysis experiments. Contrary to the above mentioned hypothesis it was shown that during pyrolysis of biomass 20 to 50% of the total chlorine evaporated already at 400°C.[2]. See Figure 1.
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80-
m40-
P0-
n Fig. 1 Percentage of released chlorine as a function of the pyrolytic temperature for different biomasses.
To explain the release of chlorine at low temperature, a mechanism has been suggested in which the chlorine in the fuel during pyrolysis can be transformed to the gaseous and volatile hydrogen chloride.[2] This mechanism suggests that: (1)
(2) (3)
The organic matter in the fuel contains or develops under pyrolysis a number of carboxylic acid (-COOH) groups. These carboxylic groups react with KC1 in an ion exchange reaction which sets free HCl: R-COOH(s) + KCl(S) + R-COOK(s) + HCl(g) HCl is then evaporated to the gas phase. The potassium carboxylate is more stable but can be deteriorated in a later step e.g. R-COOK + R + CO&) + K(g)
The hypothesis is that the pyrolytic thermal degradation of the biofbels leads to the formation of chain-bound carboxylic groups which then react with KCl or NaCl under formation of alkaline carboxylate and gaseous HCI. This mechanism would then explain: (1) (2)
(3)
The early release of a significant part of the fuel chlorine during biomass pyrolysis. The observation that the early release of C1 not requires any hetero atoms (e.g. N,S,P) in the structure. Why the early release of C1 needs a threshold temperature when the CI and the alkaline metal ions are present in an organic material without initially accessible
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carboxylic acid groups. These groups are then formed on heating and thermal degradation. The suggested mechanism was confirmed in series of experiments with three model compounds cellulose, oxicellulose, and xylan [3]. In the present work a great number of new model compounds were chosen to study the activity of different functional groups more in detail. The aim with the work was to verify the suggested mechanism and to investigate the influence from other functional groups on the early release of chlorine from organic materials. EXPERIMENTAL
METHOD The model compounds were all pretreated with a saturated and filtrated solution of potassium chloride KCI (2%) in methanol. A solution in methanol was chosen to establish a good contact of the materials with the KCl. Water solution was avoided, since the methanol can more easily be removed from the samples without heating. Approximately 2 g. of the desired model compounds were heated in a stream of dry nitrogen (pyrolysis) or of a mixture of 92 % nitrogen and 8% oxygen (gasification). The sample was placed in a combustion boat of porcelain. The boat was heated in a quartz tube which was placed in an electrically heated furnace. The furnace was heated with 50°C/min to the desired temperature which was kept for 30 minutes, and the sample was cooled under nitrogen and weighed. The fraction of solid residue left in the vessel was calculated and analysed on its chlorine content. Whenever possible, the three temperature levels 300, 400, and 500 "C were studied for all materials and both gas atmospheres. In some cases the experiments failed, or were impossible. This was the case for e.g. adipic acid at temperatures above 300 "C, since the vapor pressure of adipic acid is too high
MODEL COMPOUNDS The model compounds were chosen so that they contained structural elements and functional groups which are usual in the natural materials that constitute the biomasses. Both macromolecular (polymeric) and low-molecular compounds were chosen. Some of the model compounds were natural materials or extracted from them, some were synthetic, some were natural materials after chemical treatment. Table 1 gives a list over all model compounds, and which functional groups they contain.
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Table I ;
Model compounds and their functional groups.
Model Compound Cellulose Oxicellulose Xylan Cellulose triacetate [CTA] Pectin Poly(acry1ic acid) [PAA] Poly(methy1methacrylate) [PMMA] Copolymer of ethylene and acrylic acid [PE-PAA] Polyesterpolyol [Poly-poly] ester groups building the chain Polyvinyl butyral [PVB] Glucose Sucrose hemiacetal Maltose hemiacetal Lactose hemiacetal Cellobiose hemiacetal Sorbitol Gluconolacton group attached to lacton ring Mucic acid Adipic acid BM: BM-Cl?
SM:
I1
Type Functional groups BM exocyclic hydroxymethyl groups, acetal BM-CT free carboxylic acid groups, acetal BM acetal BM-CT acetal, open ester free carboxylic acid, open ester, acetal BM free carboxylic acid on polymer chain SM open ester, attached to polymer chain SM ~~
SM SM
free carboxylic acid attached to polymer chain hydroxymethyl groups attached to polymer chain
SM BM BM
acetal groups (rings) attached to polymer chain exocyclic hydroxymethyl group, hemiacetal exocyclic hydroxymethyl groups, acetal,
BM
exocyclic hydroxymethyl groups, acetal,
BM
exocyclic hydroxymethyl groups, acetal,
BM
exocyclic hydroxymethyl groups, acetal,
BM-CT terminal hydroxymethyl groups BM-CT ring-constituting ester (lacton), hydroxymethyl BM-CT SM
free carboxylic acid free carboxylic acid
biological material or extractedfromsuch material biological material, modied by chemical treatment synthetic material abbreviation used in figures etc.
RESULTS AND DISCUSSION In figure 2 results from the model compounds cellulose, oxicellulose, and xylan, are shown.
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100
90
80 70
60
50 40
30 20 10
0
Model compounds
Fig. 2 Release of chlorine (%) from KC1-pretreated model compounds cellulose, oxicellulose, and xylan, during pyrolysis (Nz) and gasification (92Y0N2+ 8%02) experiments. Figure 2 shows the previously presented results from pyrolysis [ l ] together with the new results from gasification of cellulose, oxicellulose and xylan. Chlorine is released to a significant degree from both cellulose and oxicellulose at 300°C, but not from xylan, during pyrolysing conditions. No results from gasification of cellulose could be obtained since there was to small amounts left after the experiments to analyse. For xylan a small release of
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chlorine can be observed at 5OOOC but more significant release of chlorine do not occur until 600°C. This result confirms the suggested reaction mechanism. The details have been discussed previously [3]. Figures 3 and 4 show all the results from the chlorine release obtained from experiments with model compounds. 100
90
80
10
1
00
J
e
50
h
L,
40
30
20
10
0
Fig. 3 Release of chlorine (?h)from KCl-pretreated macro-molecular model compounds during pyrolysis (N2) and gasification (92%N2+ 8%02) experiments. Abbreviations explained in table 1. From the results the strong effect of the free carboxylic groups on the chlorine release is verified. The chlorine release from the compounds with free acid groups showed no or little dependence from temperature. The Polyvinylbutyral PVJ3 with its high concentration of acetal groups shows a significant chlorine release during pyrolysis, and a very high release for gasification. The latter shows that the acetal during gasification probably is rapidly oxidized under formation of free acid groups. Even the polyesterpolyol shows a significant activity, but only for gasification, not for pyrolysis. The reason is unclear. Also the model compounds with esterified carboxylic acid groups showed significant chlorine release. Their chlorine release was strongly temperature dependent, since higher temperature supports the cleavage of the ester bond under formation of free acid.
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10
60
50
I
:40
1
-5 -
30
5
10
10
0
Fig. 4 Release of chlorine (%) from KC1-pretreated molecular model compounds during pyrolysis (Nz)and gasification (92%N2 + 8%02). Figure 4 shows the results for sugars and sugar derivatives. The sorbitol, which lacks the hemiacetal and acetal groups of the sugars, shows weak chlorine release only during pyrolysis at 500 "C. In some model compounds, especially in the sugars in figure 4 , hydroxymethyl groups and groups like acetal and hemiacetal have been the dominating functional groups. For hydroxymethyl it was necessary to make a distinction between the hydroxymethyl in terminal positions (at the end of open chains) and those which are attached to rings, chains or rings which constitute chains. The gasification conditions could not be studied for some of these substances at higher temperature levels since they gave too little solid residue. It can clearly be seen that at 300°C more chlorine is released during gasification than during pyrolysis for most of the model substances. The results from all model compounds show that, with a few exceptions, a significant release of chlorine at pyrolysis conditions occurs already in the temperature range 300-500 "C, which is too early for a gasification of undissociated KCI. The exceptions have been: xylan, sorbitol, and polyesterpolyol. This strongly confirms the ion exchange mechanism which initially had been suggested. The model compounds with free carboxylic groups show the strongest effect, which is consistent with the theory. The other compounds show weaker, but temperaturedependent effect. This means that even other functional groups, such as chemically bound carboxylic acid (open or ring-shaped esters) or chemical precursors for such carboxylic acid
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(hydroxyrnethyl groups, acetals, hemiacetals) can be activated at elevated temperatures, probably by chemical rearrangements, in the case of gasification additionally supported by partial oxidation. From the results it can be concluded that the following types of functional groups contribute more or less strongly to the early release of chlorine: Free carboxylic acid Carboxylic acid in ester Acetal Hemiacetal Hydroxymethyl Hydroxymethyl
(strongest effect, temperature-independent) (open ester or lacton ring)
(when attached onto chains or rings) (when at the end of a chain, “terminal”)
RANKING Based on this conclusion, a method has been developed to determine the relative activity of the functional groups. This was done based on the assumption that the observed effect (chlorine release in %) is caused by the simple sum of the contributions from the different functional groups, which are present in the model compounds. For all model compounds where reliable values for the chlorine release had been obtained , the concentrations C~unctional of all functional groups from the materials “as used” were determined. A value S was defined for each model compound, which was calculated from these concentrations:
These S-values for the model compounds were plotted against the results for chlorine release for the temperatures 300,400, and 500 “C (pyrolysis) and for 300 “C (gasification). For gasification at higher temperatures there were too few reliable results. Using linear regression the relative influence from each functional group was calculated. The factors (a,b,c,d,e,f,g) were chosen in that way that their total amount equals 100 (normalized results). These factors give a crude ranking of the activity of the functional groups. The results are presented in Figure 5.
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El300"CW .WCW
BWCN2
300°C W+o2
Figure 5. The relative effect of the presence of certain functional groups in the model compounds on the chlorine release during pyrolysis or gasification in the temperature range 300-500 "C.
SUMMARY
The earlier results from pyrolysis of cellulose, oxicellulose and xylan together with the new results from gasification of these substances and all results from experiments with organic model compounds confirm the previously suggested mechanism (ion exchange with carboxylic acid). (1)
The release of chlorine at pyrolysis or gasification of chlorine-containing biomasses or similar model compounds is significant already at temperature levels which are several hundreds of degrees lower than what would be expected for the volatilisation of KCI. In some cases, less than 300 "C can be sufficient.
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Many model compounds which only consist of chemically bounded C, H, and 0,and pretreated with KCI, can show the early chlorine release when pyrolysed or gasified at 400 "C. No heteroatoms such as S, N, or others, are required for the effect. In most cases, the release at gasification conditions is higher than at pyrolysis conditions. As expected, the model compounds with significant amounts of free carboxylic acid groups showed the strongest chlorine release. Other model or natural compounds which contained no free carboxylic acid groups, showed significant and temperature-dependent chlorine release at pyrolysis or gasification between 300 "C and 500 "C when at least one of the following functional groups was frequent in the substance: Ester in open or ring-bound (lactonic) form Acetal Hemiace tal Hydroxymethyl (-C&-OH), when attached to a ring or onto the side of a chain. Hydroxymethyl was found to be inactive at 300 "C, when it formed the end of a straight unbranched molecule (or polymer chain), but its activity increases strongly with increasing temperature. A simple ranking of the relative activity of these functional groups showed that the relative activity of free carboxylic groups is strongest at 300 "C at both pyrolysis and gasification conditions. This effect is more or less ''wiped out" at higher temperatures at pyrolysis. The effect of the other (non-carboxylic acid) functional groups can be explained by their ability to rearrange under formation of active forms like carboxylic acid at elevated temperatures. CONCLUSIONS
In all biomass fuels, several of the functional groups active in the low temperature release of chlorine are present in significant amounts in one or several of the biopolymers, which constitute the biomass. The consequence for practical applications is that the effect of chlorine release at the pyrolysis or gasification of chlorine-containing fuels will be significant as soon as a temperature level of about 300 "C is exceeded. The ability to release a part of their chlorine concentration already at comparably "moderate" temperatures seems to be an inherent property of these fuels. This seems to be caused by the chemical properties of the natural substances that constitute the main part of these fuels.
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ACKNOWLEDGEMENT This work has been performed within the program “FB firbr%nning/fdrgasning” under contract no P1286. The support f”rom STEM, Swedish National Energy Administration, is gratefully acknowledged.
REFERENCES 1. Mojtahedi,W. & Backman,R. (1989) The fate of sodium and potassium in the pressurised-bed combustion and gasification of peat. J.Inst.Energy, 1989, 189-196. 2. Bjorkman,E.& Strbmberg,B. (1997) Release of chlorine from biomass at pyrolysis and gasification conditions. Energy&Fuels, 11, 1026-1032. 3. F. Zintl, B. Stromberg, and E. Bjorkman (1998) Release of chlorinefiom biomass at gasification conditions. In: “Biomass for Energy and Industry”, 10’ European Conference and Technology Exhibition, Wiirzburg, Germany 8-1 1 June 1998, pp. 16081611.
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The Char Residues from Pyrolysis of Biomass some Physical Properties of Importance E.M. Suuberg, I. Aarna, I. Milosavljevic Brown University,Division of Engineering, Providence, R.I. 02912 USA
Abstruct:The pyrolysis of pure cellulose and actual biomass materials has been studied with an eye towards characterization of the char products formed during pyrolysis. Very few data exist regarding the thermal properties of biomass or cellulose chars, particularly at temperatures above ambient. Those engaged in modeling the pyrolysis behavior of these materials have had to guess at many of the key parameters. Several of these parameters have been here experimentally determined as part of an examination of heat transfer controlled pyrolysis of pure cellulosic material. It is likewise the case that the development of porosity in biomass materials during pyrolysis is not yet well understood, despite the significant commercial interest in producing activated carbons from these materials. This study also provides experimental data on this topic.
INTRODUCTION THE THERMAL PROPERTIES OF A MODEL CELLULOSIC MATERIAL The pyrolysis and combustion of cellulosic solids have been extensively studied phenomena. There have been many reviews of different aspects of the problem (e.g., biomass pyrolysis[ 1-31, modeling of combustion [4] and smoldering [ 5 ] ) . There remains some controversy regarding even the most basic aspects of the pyrolysis process, including even the apparent kinetics of pyrolysis of pure cellulose [6-lo]. It is clear that transport limitations of various kinds are responsible for some of the disagreement in the literature, though the relative importance of heat and mass transport limitations have yet to be fully sorted out. The process of evaporation of pyrolysis tars (or partially decomposed sugars) has been seen to be an important facet of both the heat and mass transport processes. In a study on the thermal effects of cellulose pyrolysis [ l l ] , it was concluded that the endothermicity of pyrolysis arises mainly from the latent heat of tar evaporation, and is counterbalancedby exothedc char formation due to tar cracking. The magnitude of the pyrolysis endotherm is comparable to the sensible enthalpy requirements, under certain conditions. Thus the thermal requirements of the pyrolysis process are themselves quite complicated. In addition, while the processes governing the transport of heat to the reaction zone may be qualitatively understood, there remain
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significant uncertainties regarding the rates of heat transport. In this paper, some of the key thermal and transport properties governing the rate of the pyrolysis process in cellulosic solids are examined. The original motivation for studying the thermal properties of cellulosic chars came from our study on bulk cellulose pyrolysis under conditions simulating those existing in a fire. In such a situation, the flame over the surface of the solid supplies heat to the pyrolyzing solid. In our work, the radiative and conductive feedback of heat from the flame to the surface was simulated using radiant heaters. The experiments were carried out in an inert gas environment, to maintain as well-defined a heat transfer environment as possible, free from complications due to actual combustion heat sources. A convective flow of the inert gas was used to sweep away volatiles from the vicinity of the surface, and the heat transfer effects of the sweep gas were also taken into account. The material selected for study was pure cellulose. This material was selected in an effort to produce “realistic”, but highly reproducible samples. Samples were pressed to densities typical of those encountered with wood. Such samples had the advantage that a large amount of kinetic work on cellulose could be brought to bear in understanding the results of the experiments.The work was concerned with the behavior of bulk material, in which sample dimensions are at least of order centimeters in all directions. There have been many other studies of cellulosic pyrolysis on this scale [e.g., 12-21]. Some have noted a well-defined pyrolysis front [12-161 and a period of relatively constant mass loss rate [15,16,19,21],particularly with high imposed radiative fluxes. Others have seen a sharp rise followed by a slow decrease in mass loss rate [17]. Generally, the data revealed a significant sensitivity to char thermal conductivity [20]. The effects of oxygen [I81 and retardants [21] can be significant in affecting observed behavior. In summary, experiments with bulk cellulosic samples have provided evidence of a conduction limited process, often characterizedby a very sharp boundary between char and unreacted cellulose, particularly at high incident radiative fluxes. The pyrolysis wave is often seen to move at near-constant velocity through a significant portion of the process, as our own results have also shown [22]. In attempting to model the heat transfer-controlled behavior, we were surprised to find, in the literature, a significant variability in the key thermal properties that such a model required. Quite different predictions could result from using different assumptions regarding the thermal properties. To address this obviously significant impediment to reliable model development, an experimental program focusing on thermal property measurements was initiated. The results are summarized below.
THE POROUS PROPERTIES OF BIOMASS CHARS The interest in the properties of the chars derived from cellulosic or biomass solids extends beyond those associated with thermal transport in the char. Insofar as the char residue from a pyrolysis process must typically be burned, gasified, or put to use as an activated carbon product, there is also a need to examine the porous nature of the char. In activated carbons, the pore structure is key to adsorption performance. In combustion or gasification, the porosity can play a role in determining conversion kinetics in the intrinsic rate controlled or pore diffusion controlled regimes. There exists significant empirical understanding of how to prepare useful activated carbon products from biomass materials. Still, there is relatively little information available on this topic to guide those whose main interest is in the pyrolysis of biomass to produce oils or combustion energy. In these types of processes, the conditions are often
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far removed from those used to produce activated carbons. This paper therefore also outlines some of the main characteristics of the porous structure developed in biomass materials during pyrolysis and during subsequent oxidative gasification.
EXPERIMENTAL MATERIALS The cellulose studied in this work was Whatman CF-11 powder, with an ash content of 0.009% and a moisture content of 7.68% (as-received). Powder samples were formed into pellets using an ordinary laboratory press, without any additives. Three different density samples were prepared by pressing for varying lengths of time at various pressures. The final pellet densities were: 0.965k0.041 glcc, 0.69 h0.035 g/cc and 0.458k0.026 g k c . These densities are comparable to those of woods ranging from softwoods to hardwoods. All samples were 38 mm in diameter and had a thickness of about 10 mm. The pellets were generally quite hard; only the lowest density material showed some propensity to lose powder from near its edges. These pellets were used in the simulated fire apparatus mentioned above, in order to examine the kinetics of volatile release in situations in which the screening of oxygen from the solid surface (by the volatiles efflux) is simulated by performing the experiments in inert gas [22]. These experiments will not be described further here, except as needed to explain the thermal property results. It was observed that in pressing the samples, the cellulose fibers tended to mat down in the plane perpendicular to the pressing direction. This gave samples a somewhat banded appearance. Experiments were conducted with the banded structure (grain structure) both perpendicular to and parallel to the incident radiative flux. This seemed to make little difference in the results that were obtained. Generally, the simulated fire experiments were conducted with the samples having the bands perpendicular to the incident heat flux, and this was the orientation for the thermal property measurements as well. Again, for these pure cellulosic materials, the orientation seemed to be of little significance. In the case of the porosity measurements, in addition to the same cellulose samples, samples of oak and white pine were also examined. These samples of wood were ordinary commercial lumber samples (the same pine having been previously used in fire experiments at the Center for Fire Research of the National Institute for Standards and Technology). The pine had an initial dry density of 0.377 glcc, and the oak 0.734 glcc. Both of these wood samples were pyrolyzed for two hours at 1273 K, in helium, in order to prepare the chars for testing. The char yields were in the neighborhood of 28%by mass of the starting oak and 26%by mass of the starting pine.
RADIATION ABSORPTIVITY MEASUREMENTS Virgin cellulose pellets and cellulose chars produced in the simulated fire apparatus were both examined. Two different measurements were made. One involved measuring the reflected radiation in the mid-infrared from 2.5 to 25 pm (4000 to 400 cm-1). These measurements were performed in a diffuse reflectance cell within an FTIR spectrometer. These experiments revealed some wavelength dependence of reflectivity. Reflectance was also measured in-situ in the simulated fire apparatus, by arranging the samples, a fluxmeter, and the heating lamps such that surface reflection of the incident radiation
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could be determined over a fairly wide range of angles around the specular angle. This arrangement did not assure the collection of radiation over all angles, as did the FTIR measurement. There will be a dominance of a specular component in the reported values below. The results from this technique were, however, compared with independently calculated values from a surface energy balance on a sample instrumented with nearsurface thermocouples. The agreement appeared to be reasonable, implying that the diffuse reflectance component did not contribute much to the energy balance. Measurements of the reflectivity were performed with the cellulose or char surface still cold, such that reradiation from the pellet surface was not significant.
THERMAL DIFFUSIVITY MEASUREMENTS Thermal diffusivity measurements were performed on pellets by the use of a transient response technique. Pellets of cellulose or char were instrumented with thermocouples at known distances from the surface. Samples were then brought into contact with a hotplate whose surface temperature varied in an approximately sinusoidal manner. Samples of virgin cellulose could only be tested up to surface temperatures of about 500 K, to avoid pyrolysis. Char samples could be tested to temperatures up to those seen during pyrolysis (in excess of 750 K). Samples were covered by a bell jar, purged with nitrogen, in order to avoid reactions with oxygen. The pseudo-one dimensional conduction problem defined by the experimental system has a well-known solution [23]. While it proved difficult in practice to satisfy perfectly the boundary conditions for the ideal solution, a good enough approximation was possible such that the phase lag could be used to calculate the diffusivity directly [24]. Together with information on the sample pellet density (from measurements of its mass and volume) and heat capacity (see below), the thermal conductivity of the sample could be calculated.
HEAT CAPACITY Heat capacities were determined by using a differential scanning calorimeter (DSC). A sapphire standard was used for DSC cell calibration. Samples were cut directly from the samples used in the thermal diffusivity measurements. Heat capacity was determined, as usual, from the heat requirement of the sample in response to a particular change in temperature. Multiple scans were performed to verify that sample degradation did not influence the results. Char samples could be heated up to 850 K in these experiments. The DSC cell was continuously purged with nitrogen, in order to avoid oxidation of the samples.
POROSITY MEASUREMENTS Adsorption isotherms were determined in an automated volumetric gas adsorption apparatus (Autosorb 1, Quantachrome Co.). Adsorption of nitrogen was performed at 77 K . Before measurements, samples were outgassed at 672 K for at least 8 hr in vacuum. Where some burnoff of the char was desired, the reactions were performed in an Online Instruments TG-plus thermogravimetric analyzer. The reactions were performed in a mixture of helium and oxygen, flowing at a rate of about 220 cc/min. Samples of 30-50 mg were dispersed on a circular platinum pan with a large flat surface and raised sides, resulting in a particle beds of about I mm thickness. Temperatures between 573-748 K
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were used for gasification. The partial pressures of oxygen was 2.02 kPa. Char samples were outgassed at 1173 K for 30 minutes prior to reactivity measurements. Samples were then reacted to the appropriate level of burnoff and then quenched. Burnoff is expressed on a dry,ash-free basis.
RESULTS AND DISCUSSION CHAR YIELD AND DENSITY The char yield is defined here as the fractional mass left after the mass loss ceases under given conditions. The char yield was not observed to be a very strong function of initial cellulose sample pellet density, over the range examined here. The char yield did appear to be a function of the temperature that any particular point in the sample achieved. The asymptotic char yield for these bulk samples was in the neighborhood of 10%. Unfortunately, the role of mass transfer cannot be disregarded in predicting what the char yield should be under any particular conditions. It is possible that deep within a sample, the char yield is higher than at its periphery due to a greater opportunity for secondary cracking of tar volatiles to take place due to their hindered escape from the hot solid [7]. Our earlier work clearly established the importance of the interplay of thermal history and mass transport limitations in determining char yields. Slower heating and more resistance to escape of volatiles clearly promotes higher char yields [7]. The above naturally raises a question as to what exactly is meant by the term “char” in the context of bulk pyrolysis experiments. The definition can clearly only be operational, and is simply the mass that is left when an apparently steady state condition is achieved, regardless of how incomplete the pyrolysis is or however much carbon has resulted from processes involving secondary carbonization reactions of tars and their precursors. Thus all reports of “properties” of cellulosic chars must be viewed in this light. Figure 1 shows the variation in actual char density, as a function of char yield. These data were obtained from various experiments in which bulk samples were pyrolyzed in the simulated fire apparatus to a surface temperature of around 800 K. Different ultimate char yields were observed as a function of position in the samples. Different ultimate char yields were observed as a function of position in the samples. It should be carefully noted that the variation in final char density with char yield is not inconsistent with the above observations concerning the lack of effect of initial cellulose density on the ultimate fractional char yield. The important conclusion is that char density is a nearly linear function of char yield, as might be expected. Such an assumption has been commonly employed in modeling of the pyrolysis behavior. The slope is, however, not what would be expected from a simple extrapolation to zero char yield. The true density would be underpredicted by such a common extrapolation. The reason is that shrinkage during pyrolysis is significant. Such shrinkage would tend to increase the density of the char in comparison with a simple extrapolation to zero char yield. The error associated with neglect of shrinkage can be seen by a comparison of the dashed line and the solid line in Figure 1. The shrinkage is observed during all experiments, and makes the calculation of actual temperature gradients in a char problematic.
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1
-
a
2
0.8
0.6 0.4 0.2 0
0
20
40
60
80
100
120
Char Yleld [%I
Figure I. The ratio of final to initial cellulose sample densities for various char yields.
HEAT CAPACITIES OF CELLULOSE AND CHAR Values of the heat capacity of the cellulose and chars are shown in Figure 2. The different curves for “chars” reflect the values obtained for cellulose pyrolyzed between 300 and 600°C.
I--6m
3
0” ,, /
0.5
300
500 700 TEMPERATURE [K]
900
Figure 2. The heat capacities of virgin cellulose (solid line) and cellulose chars (heavy broken lines). Graphite is also shown for reference (thin dashed line).
There was, within the uncertainty of the heat capacity data, no variation with preparation conditions (neither initial pellet density nor pyrolysis temperature made much difference). The values in Figure 2 are at considerable variance with some values used in modeling of wood and biomass pyrolysis, depending upon temperature. For example, values for wood and biomass chars have been taken to be 1.38 [25], 2.5 [26], 1.0 [27-29],0.71 [30],0.67
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[31]. 2.0 [32], and 2.3 1331, all in J/g-K. The measured values of Figure 2 show that the true heat capacity values, for the cellulose studied here, are in the range from 1.5 to 2.5 J/g-"C, irrespective of initial density. The results show that the variation in heat capacity with temperature should not be neglected, as it often is in modeling studies. Uncertainties in the magnitude of the char heat capacity have, however, been judged to be of only modest importance in determining pyrolysis rates [34,35], so it is not necessarily surprising that this parameter has not been earlier identified as a potential source of significant error. Stil1,where there is concern with the accuracy of net sensible enthalpy effects, it appears that there could be difficulties encountered in using many commonly assumed values. This problem is exacerbated by large uncertainties in enthalpies of pyrolysis [I 11, which may be significantly endo- or exo-thermic, depending upon conditions and the position of the pyrolysis front. THERMAL CONDUCTIVITIES OF CELLULOSE AND CHARS Values of the thermal conductivity of different chars are shown in Fig. 3, together with the conductivities of the cellulose pellets, and for reference, gaseous nitrogen. The chars tested for thermal diffusivity (conductivity) had fairly uniform properties because they were prepared in a pyrolysis furnace, and not in the simulated fire apparatus. This tended to minimize temperature gradients, but there was no assurance of absolutely uniform density, for the reasons noted above. Preparation of the chars followed a temperature history designed to simulate that in the simulated fwe apparatus. /.--
7
,_.-I
____.-
0.14 -
/.AA
0.06 -
F'
...................
.............
...---
0.02 I ' I
400
450 500 550 Temperature [K]
600
Figure 3. Thermal conductivity of virgin cellulose (0.458 g/cc - heavy solid line; 0.678 g/cc - thin dashed line; 0.928 g/cc - thin solid line), cellulose chars (heavy dotted lines)
and nitrogen gas (thin dotted line). The thermal diffusivities of different density cellulose samples were all comparable, meaning that their thermal conductivities scaled roughly with density. The conductivity values obtained for chars from different density samples were all rather close, because it was observed that the final char densities of samples prepared for these tests did not vary significantly, or in a consistent manner, with initial pellet density. The final char densities (0.33 to 0.52 glcc) often bracketed that of the low density unpyrolyzed cellulose sample, thus it is little surprise that these samples exhibited conductivities comparable to that of the low density unpyrolyzed cellulose pellets. This implies that it is the porosity
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of the sample that is the primary factor in determining the thermal conductivity, since the true solid densities of the char and cellulose are somewhat comparable. The char thermal conductivity values obtained here are again in good agreement with some and at variance with others that have been utilized or reported in the literature, e.g., 0.19 [271,0.05 [301,0.04 [251,0.03 [321,0.11 [33], of order 0.1 [36], around 0.15 [29] and 0.08*10-’k (“C) [37], all in W/m-K.
SURFACE ABSORPTIVITY Surface absorptivity (emissivity) was initially determined using an FTIR spectrometer operating in the mid-infrared. This wavelength was not entirely commensurate with the wavelength of the radiative heaters, but little attention was paid to this point initially, because the measurements were performed to “confirm” the black nature of the char surface. Results are shown in Figure 4, as integral reflectance through the mid-IR, as a function of pyrolysis time. The cellulose itself is seen to have a reflectivity of only about 10% to begin with, and this rapidly drops as pyrolysis begins. This is what would normally be expected as the surface blackens, and becomes more absorbing (and less reflective) during pyrolysis. Interestingly, the reflectance value began to climb again as pyrolysis continued, implying a lowering of absorptivity with carbonization. 1
o
kl8
8-
Pyrolysis Begins
; I
I b I I
64-
.
b \
21 0
500
. *. e-
,
r
a
*-
1
1000 1500 2000 2500 Time [s]
Figure 4. Surface reflectance of high density cellulose pellets as measured by FTIR in the mid-infrared. The times refer to the time of heating under an incident flux of 40 kW/rn*. Difficulties in closing a calculational energy balance in the simulated fire experiments prompted review of the FTIR absorptivity measurements (absorptivity being taken as one minus reflectivity, since the sample transmits no radiation through its interior). The alternate procedure described above was therefore employed for obtaining absorptivity in the wavelength range of the actual heat lamps. Again, this involved the use of the simulated fire apparatus and its own heaters as a radiative source, and using a radiation fluxmeter looking towards the surface for reflected radiation. The result of these experiments was that the cellulose gave an average absorptivity over the wavelength range of heater operation (about 0.5 to 8 pm) of 0.53 (k0.05). Various chars were tested at various flux levels and gave an average absorptivity of 0.77 (H.09).These latter two
1253
values may be viewed as maximum values for the respective materials since, again, they came from measurements near the specular angle. These values gave energy balances that were in considerably better agreement than those using the higher absorptivity values from the mid-IR measurements. The difference is presumably associated with the fact that the bulk of the radiation from the heat lamps is of considerably shorter wavelength than mid-infrared. There must be a rather significant wavelength dependence of absorptivity in this range. There is obviously an important consequence of using char absorptivitites of 0.950.98 as implied in Figure 4, as opposed to using the value of 0.77 which better describes the real situation in a fire. The former will not result in as rapid heating of the sample. Rarely has much concern been shown for such important departures from near-black char surface behavior in the modeling of pyrolysis in cellulosic systems. Another area of concern related to such radiative heating calculations concerns the “screening” of the surface (during active pyrolysis) by what might be described as “smoke” or an aerosol of tar droplets. In our work with the simulated fire apparatus, it appeared that the degree of screening depended upon the inert purge gas flowrate over the sample surface. Higher flowrates. not unexpectedly, helped clear away the smoke, and an optical pyrometer trained on the surface revealed an enhanced reflected light signal with increase in purge flowrate. It was difficult to obtain quantitative results from the particular experiments that were performed as part of this study. There was strong experimental evidence, though, that the extent of surface screening was at least lo%,in experiments involving heat fluxes of around 40kW/m2. This evidence was gathered in experiments in which a fluxmeter was placed into the center of a pyrolyzing sample. Taken together with the lower than expected surface absorptivity,the calculation of heat absorption rate is seen to be potentially subject to some rather severe overestimates.
POROSITY OF CELLULOSE AND BIOMASS MATERIALS The porosity of the cellulose chars was studied using nitrogen adsorption at 77K.The results for the fresh cellulose char and the char burned off to differing extents in oxygen are shown in Figure 5.
0
0.2
0.4
0.6 PIPo
0.8
1
Figure 5. Nitrogen isotherms at 77 K on cellulose chars. The values on the graph refer to the degree of burnoff in 2% oxygen at 723 K.
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The results in Figure 5 display a very typical pattern of behavior for microporous carbons. The definitions of pore sizes recommended by IUPAC are micropores (<20A), mesopores (20A-500A) and macropores (>500A), and the micropores will adsorb nitrogen at the lowest values of relative pressure, P/Po [38]. The apparent non-zero intercept on the ordinate implies an important role of the smallest micropores. The dramatic increase in the value of the intercept implies that a large amount of new microporous volume is being opened up by burnoff. It is of course precisely this sort of porosity development that is desired during the “physical activation” processes for producing activated carbons.The so-called BET surface area [38] of the cellulose char starts at a value of around 265 m2/g and grows to 1374 m2/g at 49.8% burnoff. The latter value is typical of good commercial activated carbons. These carbons have mesopore volumes of between 0.1 and 0.3 cc/g. Figure 6 shows a comparison of the nitrogen isotherms obtained on the pine and oak samples, in comparison to the isotherm on cellulose. All are for the raw chars, with no burnoff. This figure illustrates that the initial microporosity in many biomass materials is much lower than it is from pure cellulose. It is believed that this is associated with the chemical differences in the starting materials; the hemicellulose and lignin components tend to promote filling of the porosity during pyrolysis. Because the fresh wood chars have much smaller amounts of accessible microporosity, they also exhibit very low BET surface areas (37 m2/g for pine and 27m2/g for oak char). Figure 7 shows that the development of microporosity is very significant in the wood chars, just as it was in the cellulose char. In fact, a very similar pattern of surface area development is seen in the wood chars as in the cellulose chars. The implication of the above resuks is that cellulosic and biomass materials tend to form microporous chars. This is already well known in the activated carbon community. The tendency to form microporous materials has to do with the forming of a rigid, highly crosslinked structure during pyrolysis. Materials that tend to melt during pyrolysis generally give rise to ordered, mesophasic carbon materials with low microporosity. These results also tend to emphasize the fact that it is inappropriate to characterize biomass char porosity solely on the basis of fresh char porosity. While the fresh chars have very low microporosity, they are quickly activated, or “opened up” upon exposure to oxidizing gases at high temperatures (oxygen in this case, but carbon dioxide, nitric oxide and steam also are effective). The rapid opening of porosity with bumoff at low burnoffs carries with it another important consideration. Because the reactivity of the char to oxidizing gases depends upon the ability of the gases to reach reactive surface, the development of porosity at low burnoffs carries with it an important increase in apparent reactivity of the char with burnoff. Thus it is important in characterizations of reactivity of pyrolysis chars to recognize this possibility, and to conduct reactivity experiments at a range of burnoffs, under which conditions this trend will become more apparent. This topic will be explored in a forthcoming paper by our group.
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0
0.2 0.4
0.6 PIP0
0.8
1
Figure 6. The nitrogen 77 K isotherms on raw cellulose char (filled circles), pine char (open circles) and oak char (open squares).
A /
ti m
0 0 10 20 30 40 50 60 70 Burnoff [%I
Figure 7. The nitrogen BET surface areas for raw cellulose char (filled circles), pine char (open circles) and oak char (open squares).
CONCLUSIONS The literature shows a wide variation in values of key thermal properties of cellulosics and their chars. Values of thermal conductivity, heat capacity, enthalpy of pyrolysis and surface emissivity apparently cannot yet be safely taken from one study and applied in another. The values of these properties obtained in this study suggest that the specific heat capacity of cellulose char varies significantly with temperature between about 1.3 J/g-K at ambient temperature to about 2.5 Jlg-K at 800 K. The heat capacity of the raw cellulose is comparable to that of the char at low temperatures, but quite a bit higher at temperatures just below the onset of pyrolysis (around 550 K). The thermal conductivity of cellulose chars is a strong function of final sample density. Thus the sample porosity (on a macroscopic scale) plays a key role in
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determining the conductivity, which was typically found to be in the range from 0.04 to 0.08 W/m-K at temperatures from ambient to 575 K. The cellulose samples shrink during pyrolysis, and achieve a density in the range from 0.33 to 0.53 g k c , regardless of initial density. The surface absorptivity or emissivity of cellulose chars cannot be safely assumed to be near unity. Near unity values are found for wavelengths in the mid-infrared, but at the shorter wavelengths characteristic of thermal radiation in combustion environments, the emissivity may be closer to 0.8. Significant energy balance errors may be made in assuming higher values. Whole biomass samples are characterized by rather low microporosity and surface areas when freshly prepared. This is in contrast to pure cellulose, which exhibits significant microporosity initially. The biomass samples do, however, develop a significant amount of microporosity (activate) with modest amounts of burnoff in oxygen. In this regard, cellulose is a good model for whole biomass behavior.
Acknowledgment The financial support of the U.S Department of Energy (under grant DE-FG2699FI'40582) and the National Science Foundation (under grant BES-9523794) is gratefully acknowledged.
REFERENCES 1. Antal, M.J. Jr. (1982) in Advances in Solar Energy, (K. Boer and J. Duffie, Eds.), American Solar Energy Society, Vol.1, p. 61. 2. Antal, M.J. Jr. (1985) in Advances in Solar Energy, (K. Boer and J. Duffie, Eds.), American Solar Energy Society, V01.2, p. 175. 3. Antal, M.J., Jr. and Varhegyi, G. (1995), I&EC Res., 34, 703-717. 4. Di Blasi, C. (1993) Prog. Energy Comb. Sci, 19,71-104. 5. Ohlemiller, T. (1985) Prog. Energy Comb. Sci, 11,277-310. 6. Milosavljevic, I. and Suuberg, E.M. (1995) Z&EC Res., 34, 1081-1091. 7. Suuberg, E.M., Milosavljevic, I. and Oja, V. (1996) 26th Symp. (Int.) on Combustion,, The Combustion Institute, Pittsburgh, pp 1515-1521. 8. Antal M. J., Jr., Varhegyi G. and Jakab E. (1998) I&EC Res., 37, 1267-1275. 9. Reynolds, J.G. and Burnham, A.K. (1997)Energy and Fuels, 11,88-97. 10 Antal, M.J.Jr. and Varhegyi, G. (1997), Energy and Fuels, 11, 1309-1310. 11. Milosavljevic, I. and Suuberg, E.M. (1996) I&EC Res., 35:653-662. 12. Kanury, M. A. and Blackshear, P.L.(1966) Pyrodynamics 4,285-298. 13. Blackshear, P.L., Jr., and Kanury, M.A. (1965) 10th Symp. (Int.) on Combustion, The Combustion Institute, Pittsburgh, pp. 91 1-923. 14. Kanury, M.A. and Blackshear, P.L., Jr. (1970) Comb. Sci. Tech., 1,339-355. 15. Lee, C.K., Chaiken, R.F. and Singer, J.M. (1976) 16th Symp. (Inr.) on Combustion, The Combustion Institute, Pittsburgh, pp. 1459-1470. 16. Lee. C.K. and Diehl, J.R. (1981) Comb. Flame,42, 123-138. 17. Ohlemiller, T.J., Kashiwagi, T. and Werner, K. (1987) Comb. Flame,69, 155-170. 18. Kashiwagi, T., Ohlemiller, T.J. and Werner, K. (1987) Comb. Flame,69,331-345.
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19. Becker, H.A. and Phillips, A.M., Comb. Flame (1984) 58,255-271. 20. Chan, W.-C., Kelbon, M., and Krieger, B.B. (1985)Fuel,64, 1505-1513. 21. Chen, Y., Frendi, A., Tewari, S.S., and Sibulkin, M. (1991) Comb. Flame,84, 121- 140. 22. Milosavljevic, I. and Suuberg, E.M. (1992) ACS Div. Fuel Chem. Prepr., 37, 1567-1574. 23. Carslaw, H.S. and Jaeger, J.C. (1959) Conduction of Heat in Solids, Clarendon Press, Oxford. 24. Milosavljevic, I. (1994) Ph.D. Thesis, Division of Engineering, Brown University. 25. Atreya, A. and Wichman, I.S. (1989) J. Heat Transfer, 111,719-725. 26. Kung, H-C. (1972) Comb. Flame, 18, 185-195. 27.Kung, H.-C. (1975) 15th Symp. (Int.) on Combustion, The Combustion Institute, Pittsburgh, pp. 243-253. 28. Holve, D.J. and Kanury, A.M. (1982) J . Heat Transfer, 104,344-350. 29. Evans, D.D. and Emmons, H. (1977) Fire Res., 1,57-66. 30. Kansa, E.J., Perlee, H.E., and Chaiken,R.F. (1977) Comb. Flame,29,311-324. 31. Kung, H.-C. and Kalekar, A.S. (1973) Comb. Flame,20,91- 103. 32. Kanury, A.M. and Holve, D.J. (1982) J . Heat Transfer, 104,338-343. 33. Weatherford, W.D. and Sheppard, D.(1965) 10th Symp. (Znt.) on Combustion, The Combustion Institute, Pittsburgh, pp. 897-910. 34. Desrosiers, R.E. and Lin, R.J. (1984) Solar Energy,33, 187-196. 35. Chen, Y., Delichatsios, M., and Motevalli, V. (1993) Comb. Sci. Tech., 88, 309328. 36. Parker, W. J. (1988) D S c . Thesis, Dept. of Mechanical Engineering, George Washington University. 37. Koufopanos, C.A., Papayannakos, N., Maschio, G., Lucchesi, A. (1991) Can. J . Chem. Eng., 69,907-915. 38. Gregg, S.J., Sing, K.S.W. (1982) Adsorption, Surface Area and Porosity, Academic Press, New York.
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Biomass Fast Pyrolysis in an Air-blown Circulating Fluidized Bed Reactor I. Boukis, M.E. Gyftopoulou and I. Papamichael Centrefor Renewable Energy Sources (CRES) 1qhkm Marathonos Ave., 190 09 Pikermi, GREECE
ABSTRACT: A novel circulating fluidised bed reactor for biomass fast pyrolysis integrally utilizing the energy content of the by-product char in the pyrolysis reactor has been operated and M h e r amended. This paper is aiming to examine the amendments adopted for the reliable operation of the downstream processing as well as other system components, to analyse the results and to explore scale-up possibilities. In addition, potential end-use applications are discussed and first operational results of a bio-oil fuelled Stirling engine are presented.
INTRODUCTION Biomass valorisation to high value added energy products is of growing interest, since bioenergy has a series of socioeconomicadvantages, such as: (1) Direct reduction of fossil C02emissions. (2) Energy independence on a regional scale via the rational exploitation of forestry and agricultural residues. (3) Reduction of agricultural surpluses to comply with the Common Agricultural Policy. (4) Income increase in rural areas followed by a reduction in urbanisation rates. Several potential routes have been developed to evolve the vast bioenergy potential. As such, the production of pyrolysis liquids (also known as biosil or biocrude oil) via biomass fast pyrolysis is approaching market introduction and precommercialization in Europe, Bio-oil, as a liquid fuel with the strategic advantage that it can be stored until required or readily transported to where it can be most effectively utilised for power and/or heat applications. Besides the remarkable progress accomplished regarding biomass fast pyrolysis, further challenges to be faced exist in improving the technology and adapting applications to cope with the unusual behaviour and characteristics of the liquid product (1). On this background, a Circulating Fluidised Bed (CFB) reactor for biomass fast pyrolysis was extensively tested by CRES in the framework of previous JOULE I and I1 Programmes. The CFB reactor configuration conceived is differentiated from
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any other reactor system considered for biomass fast pyrolysis, since it incorporates the integral utilisation of the energy content of the by-product char in the same reactor vessel. Char derived by biomass fast pyrolysis is captured in the solids recovery system and recycled in the lower part of the CFB reactor, where it is combusted with air, thus providing the necessary energy requirements for the subsequent biomass pyrolysis carried out in the riser section (2).
IMPROVEMENTS IN DOWNSTREAM PROCESSING Design aspects and operational problems faced during the development of the CFB reactor have been discussed and reviewed in the past (3). During the first period of the CFB reactor operation (2) a number of experimental runs were successfully performed. At that time, the condensible vapors recovery system consisted of an indirectly cooled shell-and-tube heat exchanger (STHE)and a cotton wool filter (in the following, Downstream Configuration I, DC-I). The relatively lower liquids yelds obtained, in comparison to those obtained by other researchers (4), were mainly attributed to this, DC-I, configuration and particularly to the indirect pyrolysis vapors recovery system practiced, since: (1) Deposits of heavy components formed in the tubes, due to indirect cooling of gadvapors in the STHE, act as the foreftont in polymerisation reactions, converting primary o r p c vapors to secondary and tertiary tars. (2) The layer, formed by the deposits created in the heat exchanger tubes, reduces the heat transfer coefficient leading to ineffective cooling of the gadvapor products, prolonging their residence time at higher temperatures and consequently reducing the product yield. (3) The deposits created, reduced the free area and initiated tube plugging and excess pressure build-up, whle resulted to limited system availability. It has been reported (1) that the recovery of pyrolysis liquids is not a simple condensation process, because their rapid indirect cooling is leading to the formation of stable aerosols and micron-size droplets, which are usually entrained in the gaseous stream, thus avoiding capture. Therefore, impingement and coalescence of the pyrolysis vapors is considered an essential feature in any liquids recovery process. On the aforementioned grounds, a scrubbing system for the recovery of pyrolysis vapors was designed, constructed, installed and tested. This system, fiuther referred to as Downstream Configuration 11, DC-I1 (3,is actually based on the existing STHE, and is consisted of (1) A scrubbing section, virtually a plate-and-ring disk system arrangement for the proper distribution of the scrubbing medium and the maximisation of heat transfer area. (2) The previously used STHE for the indirect cooling of the scrubbing medium. (3) A properly designed bottom vessel for disengagement of liquids and pyrolysis gas, through which the pyrolysis gas coalesces hence enhancing mass transfer and improving vapors capture. (4) A recirculating Moyno-type (screw) pump with independent flow control.
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Pyrolysis experiments with wheat straw were conducted to commission DC-11. Although the experiments exhibited substantially improved system availability, they suffered lower total liquids yields (loss of char and heavy liquid fractions, see further on the experimental results section). An improved configuration of the DC-I1 gadvapors recovery system integrated an electrostatic precipitator (EP) replacing the cotton wool filter (denoted as DC-111). Gas enters the bottom of the precipitator and leaves the top, while a liquid outlet situated at the lower part of the precipitator serves for the collection of the captured pyrolysis liquids. More experiments conducted in order to validate the performance of the new downstream components, revealed improved availability and higher mass balance closures (above 95%). The final configuration of the biomass fast pyrolysis experimental set up is depicted in Figure 1 (5).
Flla
0-
nnc 41
i
I
Fig. 1 CFB reactor with DC-111.
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EXPERIlVZENTAL RESULTS
Five experimental runs (R7-RI 1) were performed with DC-I and pine of 1-2 mm as a feedstock with acceptable mass balance closures (91.99 - 97.52 %) and liquid yields up to 6 1.5 % wt on moisture and ash free (maf) feed basis. Two experiments (RSI,RS2) were performed with DC-I1 and Swedish wheat straw with particle size in the range of 1 to 2 mm. During these experiments the solids recovery system (cyclone and impinger, Figure 1) did not perform satisfactorily with the very fragile straw char. The c-har, exposed to the very harsh environment (fast moving sand particles) of the CFB reactor and atritted to very fine particles, almost submicron powder, could not be collected by conventional solids recovery systems. Consequently, the char was either adhered with the liquids heavy fraction to the STHE inner tubes or readilly accumulated in the liquid recovery system. In addition, sintering of the sand occured, due to the straw ash melting at sigmficantly lower temperatures than the respectives for wood ash. It was observed in form of white spots at the char combustor walls. After the installation of the DC-111, i.e. scrubber, STHE and EP, two more runs (R12, R13) were conducted with a mixture of softwood with a particle size in the range of 1.5 to 2 mm. Dunng run R12, increase of the EP’s supply voltage resulted in improving the collection efficiency, as indicated by the change of colour in the gaseous stream leaving the EP from brownish to transparent. Short-circuiting, attributed to dropletsassisted bridgmg between the two electrodes, occurred during the EP operation and was followed by severe instability problems to all the nearby electronic devices, as well as significant reduction of the EPs voltage. After the completion of R12 the scrubber was dismantled and cleaned. It was noticed that the solids loading was much less than the respective of the straw experiments. In run R13, the EP performed satisfactorilyand the pyrolysis liquids recovered from this point had very different physical appearance (higher density and viscosity and lower moisture content) compared to the liquids recovered from the scrubber. However, short-circuiting occurred again in the EP approximately after 30 minutes of operation, lading to the thought of installing a pump for the continuous removal of the collected liquids. In total, five pyrolysis experiments were conducted with the latest configuration (DCIII, Figure 1) and a mixture of hardwood in order to achieve longer and more reliable operation of the system. The particle size of the feedstock was within the range if 1.5 to 2 mm. Operating conditions for all runs are presented in Table 1. Operation of the biomass fast pyrolysis pilot plant with DC-III and woody feedstocks, specifically mixed softwoods (mainly pine) and hardwoods, revealed the following: (1) The reactor operability is easily sustained once steady-state operation is reached. (2) The woody feedstocks produced much less fine char than straw, and the gas
entrained solids were more efficiently recovered by the solids recovery equipment. (3) The modified vapor recovery system (DC 111) resulted in the higher pyrolysis liquids yields obtained (63.4% on a maf feedstock basis) for the feedstock size
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(1.5%<2.0 mm) tested. Nevertheless, these yields are still lower than those achieved by others researchers due to reasons referred to above. The quality of the bio-oil was significantly improved with respect to the liquids water content. During the earlier DC-I and I1 experiments, the loss of the heavy fraction of the pyrolysis liquids (retained in the cotton-wool filter) resulted in a wattery liquid (> 35% wt water), while during DC-111 experiments the water content was reduced to approx. 20% wt. A sigruficant problem remains the relatively high solids (char) content in the pyrolysis liquids, due to the fragmentation of char and wood fines in the harsh environment which are subsequently entrained in the gaseous stream, an inherent characteristic encountered in all CFB reactor systems, (ineffective solids recovery system).
Table 1 Operating conditions and total liquid yields for the experiments conducted.
RUn
number R7 @GI) R8 @GI) R9 @GI) R10 @GI) R11@C-I) Rs-1 @C-11) Rs-2 @C-11) R12 @C-UI) R13@C-III) R14 @C-111) R15 @C-111) R16 @C-111)
T TCom- Biomass Fluidizing Vapor Totalliquids Mass Riser bustor flow-rate air residence yield balance ('c) ('c) (kg flow-rate time (ms) (%wt on maf closures -) feedstock) (%)
&a)
581 605 579 496 550 630 670 470 520 450 500 535
678 700 645 748 740 700 720 610 700 520 580 605
8.44 10.45 10.20 8.96 10.22 13.20 9.90 8.50 8.50 10.30 9.64 10.17
16.00 14.70 13.90 12.10 13.00 17.00 13.80 18.40 17.50 14.70 13.88 14.78
374 388 396 458 373 360 477 557 557 459 491 465
40.12 54.46 57.43 61.50 49.91 43.00 38.00 61.00 56.00 63.40 62.80 61.90
94.00 95.36 97.52 94.18 91.99 89.78 91.42 97.29 97.61 95.14 94.91 95.96
FURTHER IMPROVEMENTS IN THE PILOT PLANT
After the second period of the CFB reactor operation (DC-II, DC-111) further amendments in the plant were decided to solve respectwe problems. Firstly, the existing feeding system incorporated in the pyrolysis pilot plant has been replaced, due to the following reasons: (1) Excessive wear of motors, axes and gears. (2) Not very reliable feeding (variations of up to 10-15% from average have been
observed during calibration tests). (3) Insensitivity of feeder to deliver reliably over a wide span of flowrates.
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A new feeding system has been designed, constructed and installed in the plant. It comprised of
(1) A feedstock hopper (capacity 40-45 kg of biomass) with a stimng device. (2) A system of three screws, a horizontal one, a vertical one and another horizontal
one and their relevant motors. The horizontal screw in the hopper is the main screw that regulates the biomass flow rate. The other two screws are just transporting the biomass into the reactor. The second horizontal screw is connected directly to the reactor. The feeding system has been calibrated with feedstock in the range of 1 to 12 kg/h on a wet feedstock basis. It must be noted that during the calibration of the system the repeatability and the reproducibility of the results were proved to be very accurate, especially when they were compared with the respective ones of the previous feeding system (2). The biomass flow remains relatively steady, even while the hopper is discharging to the dosing screw. A sigdicant deviation in the results appears only when less than 1 kg of biomass is left inside the hopper. This feeding system is expected to result in better control of the process and more reliable and accurate mass balances.
Secondly, aiming to avoid the short-circuiting occurring in the EP, due to droplets-assistedbridgmg between the two electrodes, a peristaltic pump was installed in the liquid outlet situated at the lower part of the precipitator. During commissioning of the EP connected with the peristaltic pump apparent improvement was noticed. Although the nearby situated electronic devices failed again at approximately 30 minutes of steady state operation, the voltage provided by the power supply unit to the EP remained relatively constant (100’ drop compared to 50-70% drop without the pump), and the EP continued to clean the pyrolysis vapours. In addition the quantity of pyrolysis liquids that were captured by the EP increased significantly. The failUte of the electronic devices is attributed to the electromagnetic field created by the high voltage of the EP. A possible, easy way to resolve this problem is the isolation of the EP with a metalic construction that will absorb the main part of the electromagnetic field. Further experiments should be performed to prove the effectiveness of this solution. INVESTIGATION OF SCALE-UP POSSIBILITIES
During the implementation of the JOW-CT96-0099 Contract (3, the potential possibilities for scaling-up the CFB biomass fist pyrolysis reactor technology have been studied. Provision of technical information to the Technical University of Vienna led to the construction of three Cold Flow Models (CFMs) with 10, 135 and 1000 kg/h nominal biomass feed rate, in order to derive scale-up guidelines for the CRES Circulating Fluidised Bed (CFS) biomass pyrolysis reactor. The methodology followed was based on the similarity rules theory (6) and the consideration of nondimensional numbers derived from the fundamental equations of two-phase flow. Extensive experiments to study the fluid mechanic behaviour have been camed out at these CFMs corresponding to CFB fast pyrolysers (5). Satisfactory agreement was reached between the 10 kg/h CRES CFB pymlyser and the CFMs,leading to the
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conclusion that appropriately scaled and designed CFMs are well suited for the investigation of three basic fluid mechanic scale-up criteria. These criteria can be used for the design of a commercial scale CFB biomass fast pyrolysis reactor (5). The first two criteria are: (1) Constant ratio of the specific circulation rate of bed material to the specific feed rate. (2) Constant gas residence time in the riser.
-
The third criterion constant rate of feed rate to bed inventory - is not that strict, All models tested retain the minimum necessary specific circulation rate of bed material (mt.Bc,,Jmb,o,>lO) to ensure adequate heat transfer to biomass particles and high feedstock ablation rates. Investigations with Werent geometric and operating parameters have aided to optimize reactor design and process operating conditions. The measurements obtained indicate sufficient specific solids circulation rates to achieve adequate heat tranfer to biomass incoming feed and close to the optimum required gas residence time in the riser of the corresponding fast pyrolyser even with a 1000 kg/h biomass feed (7). According to the acquired results, a pilot scale (approx. 100 kg maf biomassh) CFB pyrolyser would have a riser diameter of approximately 0.16 m, while a demonstration scale one (approx. 1000 kg mafbiomassh), a riser diameter of 0.40 m. The specific feed rates for these two units correspond to specific feed rates of 1.3 and 2.6 kg/m2s respectively, which are the highest for all units investigated so far in the literature (2).
POTENTIAL APPLICATIONS Biomass fast pyrolysis is apprehended to offer sigmficant economic advantages over other thermal conversion processes, as it decouples fuel production from energy generation, since the end products are liquids. Interesting challenges in further developing and modtfyng biomass fast pyrolysis technology, include the liquids upgrading and the adaptation of applications to accept the unusual behaviour and characteristics of the liquid products (suspended char, alkali metals, low pH, high viscosity, etc.)(l,8). Bio-oil end use applications have been extensively reported in the literature (8). Among the Werent end-use applications considered, small and medium scale Stirling CHP systems fuelled with bio-oil have been reported to achieve higher thermal and electric efficiencies than any other options with the same fuel in similar power range (9). In the framework of JOR3-CT984310 Contract, co-ordinated by CRES, very promising results have been initially obtained from the operation of a Stirling engine (30 kW& equipped with a newly developed Flax@ (flameless oxidation) burner utilizing bio-oil as a fuel. The special features of the Flax@ burner consist of internal premixing of combustion air and the exhaust gases and avoidance of high temperature peaks in the flame. The measured particulate emissions were extremely low, at least 10 times lower than the relevant of a diesel engine for the same power range (lo), due to prolonged fuel residence time in the Stirling engine external combustion chamber provided by the internal mixing of combustion air and
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exhaust gas. Internal mixing also resulted in the efficient bum-out of suspended fuelcontained char particles, which constitutes a major remedy for high-solids content biosils. During the operation of the Stirling CHP system, very low CO and NO, emissions have been measured (CO < 6ppm, N Q c10 ppm)(lO). Besides, very good fuel atomisation performance has been reported with an air pressure atomiser, but since air preheating was not possible, the engine efficiency suffered (1 1). It is expected that these problems will be successfully overcome within the continuation of this contract.
CONCLUSIONS AND RECOMMENDATIONS The CRES CFB biomass fast pyrolysis pilot plant performance and availability was improved after amendments to the pyrolysis vapors collection and recovery train. The latest configuration adopted (DC-111), namely a scrubbing-type device followed by an electrostatic precipitator, resulted in: (1) Improved collection efficiency. Improved quality ofthe pyrolysis liquids collected. Prolonged pyrolysis plant operation. Sigmficantly simpler cleaning and maintenance procedures. Better process integration.
(2) (3) (4) (5)
The CFB reactor operation is currently performed on a fully-autothermal mode, i.e. all additional heat supply is suspended, as soon as steady-state conditions are reached. To the authors' knowledge, this constitutes a major breakthrough, since no other single-bed reactor system has been operated on a similar mode. This achievement is far more significant since it was performed on a low-throughput, pilot plant (approx. 10 kg/h), where no internal optimisation of energy supply was practised. The results obtained also ver@ in practice the long known postulation that no external heat supply for fast pyrolysis is overall required(12). Scale-up potential has been investigated through the study of the hydrodynamics for three cold flow models. The scale-up is not considered to create any severe problems in reactor ojteration, since it will be associated with design features not easily attainable to small, bench-scale equipment. Therefore, both the operability (adoption of suitable automation measures) and the performance (rational design) are expected to be substantrsll * yimproved. The only significant opearational problem remains the relatively high solids (char) content in the liquids, due to the excessive fragmentation of char and wood fines in the harsh environment incorporated in CFJ3 reactors. It is well-known that it is difficult to properly size and.manufacture small-scale (< 1") cyclones with a high performance and adequate solids removal from gaseous (pyrolysis vapors) streams. However, this is expected to be resolved during scale-up,where larger dimensions and manufacturing capabilities will allow a more precise close-up of design calculations and mandachuing specifications. The potential utilization routes for the final products have been extended with the integration of small CHP systems, namely a Stirling engine, successfully operated on bio-oil for sufficient time.
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Additional actions should be taken in order to achieve reliable performance of the scrubber under any conditions. A more efficient solids recovery system is essential to be incorporated in the plant. Scale-up at pilot (approx. 100 kgh) or demo-scale (approx. 1000 kgh) should also be considered.
ACKNOWLEDGMENTS Financial support from the European Commission, through JOR3-CT96-0099 and JON-CT98-0310 Contracts of the JOULE 111 Programme, is appreciated.
REFERENCES 1. Bridgwater A.V. (1999) Principles and practice of biomass fast pyrolysis processes for liquids. In: J. Anal, Appl. Pyrolysis, 51, 3-22. 2. Boukis I (1997) Fast pyrolysis of biomass in a circulatingfluidised bed reactor. PhD thesis, Aston University. 3. Boulas I. (1995) Practical implications during operation of a CFE3 air-blown pyrolyser. In: Bio-oil Production and Utilisation, Proceedings of the 2"d EUCanada Workshop on Thermal Biomass Processing, (Ed. By A.V. bridgwater &E.N. Hogan), pp. 49-65, CPL Press. 4. Bridgwater A.V.,Peamke G.V.C. (2000) Fast pyrolysis processes for biomass. In: Renewable and Sustainable EnergV Reviews, 4, 1-73. 5. Final report (1999), Contract JOR3-CT96-0099, A novel approach for the integration of biomass pyrolytic conversion processes in existing markets of liquid fuels and chemicals. 6. Glicksman J. (1984). In: Chem.Eng.Sci., Vol. 39, No 9, p. 1373. 7. Haslinger W., Hofbauer H., Gavriil L., Boukis L(1999) Scale-up guidelines for a circulating fluidized bed biomass pyrolyzer. In: Proceedings of dhInternational Conference On Circulating Fluidized Beds, August 22-27, Wurzburg, Section: Process design and scale-up. 8. Bridgwater A.V., Meier D., Radlein D. (1999) An Overview of Fast Pyrolysis of Biomass. In: Organic Geochemishy, 30, 1479-1493. 9. Band A., Personal Comunication on 25-5-2000. 10. Bandi A., and F. Baumgart (2000), Fast pyrolysis liquid feed to a FLOX@ burner. In: Proceedings of Progress in Thermo-chemical Biomass Conversion, Tyrol, Austria, 17-20 September. 11. Third progress report (2000), Contract JOR3-CT98-03 10, Small - scale combined heat andpower from bio-crude oil fuelled to a Stirling engine. 12. Bogley W.J., et.al., (1977) Solid Waste Utilisation-Pyrolysis, Oak Ridge National Laboratory, Oak Ridge, TN.
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Rotating Cone Applications
Bio-Oil
Production
and
B.M. Wagenaar', R.H. Venderbosch', J. Carrasco2, R. Strenziok3, B.J. van der Aa4 BTG Biomass Technology Group BV, Enschede, The Netherlands. Centro de Investigaciones Energe'ticas Medioambientales y Technolbgicas, Madrid, EspaZa. Universitat Rostock, Rostock, Germany. 4 URAEngineering Almelo B.V.,Almelo, The Netherlands.
'
ABSTRACT: A biomass pyrolysis processing chain is considered which starts with the raw feedstock acquisition and ends with electricity production. The following results are discussed: Biomass pretreatment:Poplar and straw have been selected as typical feedstock types. First the energy use of drymg has been investigated as well as the possibilities to obtain the heat by energy integration. Milling of these feedstocks to various particulate sizes has also been examined, together with the energy requirements. The energy requirements were obtained by monitoring the power consumption of the milling engines. Development of a 200 kg/hr pyrolysis plant: The pyrolysis process includes the biomass feeding section, the pyrolysis reactor and the liquid collection system. Various tests are discussed, demonstrating the technical feasibility of the process. For example, the yields of bio-oil, char and gas are typically 70, 15 and 15 weight percent. Finally, an economic evaluation of the pyrolysis process is presented in terms of the investment and bio-oil production costs. Prime movers on pyrolysis oil: Results of test concerned with the combustion of pyrolysis oils in a boiler are reported. And finally, preliminary attempts of combustion experiments with a gas turbine will be described. INTRODUCTION Fast pyrolysis is a technology by which biomass is decomposed into bio-oil, char and gas. When wood is used as a feedstock, bio-oil is the major product (70 wt.%) The gaseous by-product (15 wt.%) can be fired in a boiler or in a gas engine. Finally, the char (15 wt.%) may be combusted in the pyrolysis unit to drive the process autothermally. Only the ash (up to 1 wt.%) is left as a waste stream.
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Fast pyrolysis produces a clean liquid biomass energy carrier. Advantages of the technology are: Bio-oil is cheaper to transport than biomass. It has a volumetric energy density of 20 GJ/m3. Wood chips have an energy density of only 4 GJ/m3. (2) Bio-oil is cleaner than biomass. The ash content in bio-oil is a factor 100 lower than in biomass. Possible contaminants llke K, Cr, and Cu remain in the char. (3) The cost of bio-oil production is relatively low due to the mild conditions. Fast pyrolysis occurs at 500 "C and at atmospheric pressure. (1)
While demonstrating the pyrolysis technology chain (including feedstock pretreatment, bio-oil production and electricity production), problems whch have to be overcome are: (1) the selection of proper feedstock pre-treatment systems whch result in low specific costs; (2) the reliability of the pyrolysis pilot plant at the scale of 200 kg per hour; (3) retrofitting of prime movers for bio-oil utilization and demonstration of the total system.
This paper will evaluate the conditions to produce bio-oil and electricity on basis of biomass fast pyrolysis. The objective is then to report, is a sequential form, the results obtained with respect to the three topics mentioned above.
BIOMASS PRE-TREATMENT The objective of this task is to consider, for each biomass type selected in a previous sub-task, the following aspects: (1) Which pre-treatment steps are required to produce a suitable feedstock for the rotating cone reactor. (2) Which type of pre-treatment equipment should be used. (3) What are the energy requirements and the associated costs to pre-treatment operations.
The particle size reduction tests for various selected biomass types have been executed in the pilot installation of CIEMAT. A flow diagram of the installation is shown in Figure 1. Basically, the equipment consists of two hammer mills connected in series, the first one has a nominal power of 11 kW and is used for pre-grinding (crushing). The second hammer mill has a nominal power of 15 kW and is used for refining the pre-ground biomass. Figure 2 shows the effect of hammer milling with sieves of different mesh sizes. The solid curve 1 depicts the size distribution of the original poplar chps. The chips have a size varying from 5 to 30 mm. Grinding of the poplar chps with a coarse sieve of 8 mm requires 44 kWh per oven dry ton (odt) as shown by the curve: ground (8 mm).
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Grinding of the poplar chips first with the 8 mm sieve and then with a 2.5 mm sieve requires a total of 132 kwh per oven dry ton. Finally, grinding of the chips first with the 8 mm sieve and then with the 2.0 mm sieve requires a total of 179 kwh per oven dry ton as shown by the dotted curve: ground (2 mm).
LEGEND 1.2.3.4.-
RECEPTION HOPPER WEIGHT SYSTEM VIBRATING SCREEN SCREW CONVEYOR
8.- CYCLONE 9.- ROTARY VALVE 10.-INTERMEDIATE BIN ll.-GRINDER (15 kW) 12.DYNAMICAIR SEPARATOR 13.-BAG FILTER 14.-CHAIN CONVEYOR
PNEUMATICTRANSPORT SOLID TRANSPORT
-
Figure I : The biomass pre-treatment plant. 100 E
90
2 80
#
70
$
60
f
50 40
E, 2
30 20
st
10
-poplar chips ground (8 mm) -..- ground (2.5 mm) ..._.. ground (2 mm)
---
0 0.01
0.1
1
10
100
Mesh size [mm]
Figure 2: Weight distribution versus the mesh sue. These results allow an assessment of the pre-treatment costs of biomass on a larger scale. The installation envisioned is a pre-treatment station as shown in Figure 1 with a throughput of 4 ton per hour. It contains a biomass dryer, a coarse grinder with an 8 mm sieve and a fine grinder with a 2 mm sieve. The following basic assumptions have been introduced in the economic evaluation:
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Heat of the pyrolysis process is used for drymg. Cost of electricity: 0.066ECUkWh for Spain. Electricity consumption in drymg: 4.7k W t o n . Moisture content of biomass after drying: 12 wt% Table I presents the investment costs for a 4 ton per hour pre-treatment station. The variable costs are presented in Table 11. The electricity cost of grinding accounts for more than 80% of the variable costs.
Site preparation Buildings Subtotal Tntal cnat
120 30
120 30 I50
120 30 1.50
150
1242
884
367
As can be observed fiom Table 11, the pretreatment cost varies between 7 and 20 Euro's/ton for a fully equipped pretreatment station which can gnnd to less than 2
rnm. Table 11: Added costs for pre-treatment. Biomass type Wheat straw
Item Poplar
Fixed cost Depreciation (1 5 yr) Labour costs Totalfuced costs Variable cost Electricity Maintenance Total variable costs Overheads (3% of fwed costs) PRE-TREATMENT COSTS
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Rice husks
275 4,3 68,
198
077
473
493
671,
58
10.8 1.5 12,4 072 19,4
6.4 1.7 871 0,2 14,3
1.5 0.3 178 092 7,o
Recent operational data from the pyrolysis plant from April 2000 show that biomass with a particulate size of 3 mm can be easily processed with could lead to a pre-treatment station with only 1 coarse grinder equipped with a sieve of 8 mrn mesh size. Such a stripped down pretreatment station would lead to pre-treatment costs, which vary from 5 to 10 Eurohon depending on the type raw material, which is treated. DEVELOPMENT OF A 200 KG/HR PLANT
The development of the rotating cone reactor started a decade ago at the University of Twente. Since then the technology has evolved from laboratory scale to a pilot scale unit which is characterized by a biomass throughput of 200 kg per hour. In this section first the pyrolysis reactor is described as the core of the technology. Then the entire pilot plant is described and finally, characterization results of the plant are presented. THE ROTATING CONE REACTOR
The core of the pyrolysis pilot plant is the rotating cone reactor which is a compact high intensity reactor in whch biomass of ambient temperature is mixed with hot sand. Upon mixing with the hot sand of 550 degrees Celsius biomass decomposes into 70 weight percent condensable vapours, 15 weight percent non-condensable gases and 15 weight percent char. An important characteristic of this reactor type is the absence of carrier gas since it is the rotating action of the cone which propels the solids from the reactor entrance to its exit. Because of the absence of carrier gas, the vapour products are not diluted and their flow is minimal. An undiluted and concentrated product flow from the reactor leads to small downstream equipment with related minimal investment costs. Finally, the 200 k g h pilot plant contains a single cone with a top diameter of 0.65 m and a bottom diameter of 0.21 rn The top angle of the cone is pi/2 radians, as is shown in Figure 3.
Pyrolysis vapouro
Biomass Hot sand
'\:.y
.
I'
.p1
Figure 3: The rotating cone reactor.
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In mechanical terms, the reactor technology is remarkably simple and robust. The rotational speed of the cone is only 300 rpm and after more than 1000 hours of operation signs of abrasion or wear are absent. Thls is due to the combination of a relatively slow moving solids phase, directed parallel to the wall, and the selection of SS3 10 as the cone wall material. The extended axis, which supports the cone, leads to minimal heat leakage and allows the bearings to operate at 30 degrees Celsius (90 OF). For reasons of robustness, heavy-duty truck bearings have been installed, which operate at 30% of their design speed in a vibration free environment. They are expected to last longer than the depreciation period of the plant. Scaling-up of the rotating cone reactor is possible by increasing its diameter. For capacities which require a cone diameter larger than 2 meters, stacking of multiple cones on a single axis leads to the lowest investment costs. This conventional approach is also encountered in centrifbgal disk separators or rotating disk contactors. With these options all pilot plant capacities between 2 and 100 tonlhr can be served. DESCRIPTION OF THE PILOT PLANT BTG 200P
The starting point of the design of the pilot plant was a specification of its three major design goals: (1) Design and operational data of the pilot plant must provide a sufficient amount of know how to keep the investor risk for a commercial size plant at an acceptable level. (2) The pilot plant know how must allow for scaling up in a single step to a commercial size plant of 2 tih. (3) All systems and sub-systems of an anticipated commercial plant must be present in the pilot plant.
These considerations did lead to a plant with minimal infrastructure requirements. In thermal energy terms, the pilot plant is self sustained and is a net producer of heat when the biomass throughput is larger than 70 kg per hour. Inert gases are not required, and as a consequence they are absent. The electricity demand of the BTG 200P is 23 kW (3 phase, 25 A, cos tp = 0.80). Finally, the cooling tower requires 0.6 m3 water per ton of biomass processed, to compensate for evaporation losses. In the present pilot plant BTG 200P all char is cornbusted to heat the process and the pyrolysis gas is burned in a stack. This leaves as emissions to the environment a flue gas stream and an ash stream. A process flow sheet of the pyrolysis process is presented in Figure 4. The mass and energy balances, which are related to the flowsheet of Figure 4, are presented in Table 111. The following assumptions have been used in the calculations: Wood is processed with 6 wt% moisture and 1 wt% ash. (2) Bio-fuel oil consists of a single phase mixture of water and organics. (1)
Because the data in Table 111 originate from mass and heat balance calculations, the balance closure is 100% for the mass and heat balance.
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Flue gas
Figure 4: The flowsheet of the pyrolysis process.
Table IIZ: Results from the mass and energy balance calculations for each flow of the process flowsheet. Flow nr. Stream Wood
2 3 4 5 6 cut rcr exit riser air riser air kg/h !%! !Ikgm kgk !%!! kg/h L feed
200
88 22
N2 0 2
88 22
184 46
co2
Pyro gas H20
12
Organics Char Ash Sand Total
2
L Ppar a]
214 29 1 1.03
108 30 2
4500 4500 829 1.03
1.03
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110 29 1 1.10
30 2 4500 4642 747 1.10
338 29 1 1.15
Table III continued
H20 Organics
-
115
47 105
3384 7560
152 3 10
10944 302 3.00
-
Char Ash Sand Total T [K] P[bara]
47 105
2
2 829 1.00
478 829 1.00
182 761 1.03
30 302 1 .oo
1 .oo
A composed photograph of the plant is shown in Figure 5 . Not visible in the photogaph are the two cyclones in series between the pyrolysis reactor and the condenser and bio-fuel oil recycle loop of the condenser.
Figure 5: A photograph of the pilot plant. The lock hopper feeding is located most on the left hand side. This feeder injects the biomass in the pyrolysis reactor. Hot sand is fed into the pyrolysis reactor from
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the char combustor, which is located above the reactor. With means of a riser, the char containing sand from the pyrolysis reactor exit is recycled into the char combustor. All gaseous pyrolysis products are directed to the bio-fuel oil condenser. Cooling of the pyrolysis vapours releases heat which is removed from the bio-fuel oil by means of a water circuit. This water circuit is air cooled.
440
460
540
480 500 520 Reactor temperature [C]
560
Figure 6: The oil yield versus the reactor temperature. Mixed wood sawdust, 1-3 mm; vapour residence time 1.O s.
A typical operational characteristic of the BTG 200P is shown in Figure 6. A maximum bio-oil yield of 74 wt% has been measured when processing clean wood waste from the wood processing industry. Since this type of sawdust is available in abundance, BTG envisions to contract this stream for large scale future operations. The first step of such grand scheme is to demonstrate the capability of pyrolysis plants to process this commercially available feedstock. Figure 6 was constructed from measurement data obtained during a long duration run of 4 days. Vapour residence t h e versus blooil yield T = 55OTc;Pine wood; 0.8-1.1 mm 80
10
0 0
1
3
2
4
5
Residence time [s]
Figure 7. The bio-fuel oil yield versus the vapour residence time.
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6
Wood decomposes almost completely at pyrolysis conditions and so do the biofuel oil vapours. Therefore, a proper plant design is aimed at minimizing the residence time of the pyrolysis vapours in the hot enclosure. A measurement campaign has been conducted on the BTG 200P to obtain the relationship between the bio-fuel oil yield and the vapour residence time in the hot enclosure. These data are presented in Figure 7. For these measurements, pine sawdust of a well defined size has been used. Figure 7 shows that 5 weight percent of vapour loss can only be achieved for a vapour residence time which is less than 2 seconds. For a residence time of 5 seconds, more than 20 weight percent of the bio-fuel oil is lost due to gas phase vapour cracking. The BTG 200P usually operates at a gas phase residence time which is less than 1.5 seconds. Current achievements of the pilot plant: (1) the plant capacity of 260 kg/hr exceeds its design specifications by 30%. (2) a total of 20 ton of bio-fuel oil has been produced for various clients. (3) a bio-oil yield of 70 wt% on dry basis. (4) the longest continuous run lasted for 4 days.
(5) 3 mm pine wood, sawdust residues from a wood waste supplier, poplar, beech and straw have been successfully converted to bio-fuel oil. PYROLYSIS ECONOMICS The pilot plant generated a sufficient amount of know how to satisfy the design goals which have been stated at the start of paragraph 3.2. The economics can be summarized in a nutshell: BTG anticipates selling bio-oil at a price of 6 euro/GJ (100 euro/ton) for quantities larger than 30 kton. The bio-fuel oil is produced from clean wood residues and qualifies for a green label. PRIME MOVERS ON PYROLYSIS OIL
Rostock University used the bio-he1 oil from BTG in various end-user applications like a combustion furnace and a turbine. Results from both applications are presented hereafter. COMBUSTION DEVELOPMENT
The combustion characteristics of the bio-fuel oil from BTG has been assessed in a 300 kW(th) furnace at Rostock University. This combustion facility is depicted in Figure 8 and consists of a burner head, an burn out chamber immediately under the burner. Connected to the flame chamber is a horizontal heat exchanger segment and finally the exhaust flue gases are diluted with air to be released by the stack.
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exhaust gas pipel charcoal filter
Figure 8: The 300 kW(th) combustor setup of the University of Rostock.
This flame tunnel set-up was originally equipped with a modified oil burner RZ 3.3 from MAN, Germany. The burner was then fitted with a pilot burner for ignition of the bio-oil. A fuel pump delivers the fuel flow rate and pressure for load levels varying between 70 % and 100 % via electrically controlled valves. In the original configuration, a malfunctioning of the fuel pump did occur when switching from diesel to bio-oil and vice-versa. It was assumed that residues of the different fuels mix and clog the pump and fuel delivery system. In order to avoid this clogging problem, an alcohol fuel is burnt immediately before switching over. After this modification a satisfactory and stable start-up of the combustion facility was acheved. When burning bio-oil is has been established that it is necessary to preheat the bio-oil to 60 70 "C.The pre-heating of the bio-oil leads to a decreased viscosity and enhances the atomisation of the oil. As a consequence, a stable flame was obtained. The condition of the flame tube after 20 minutes of bio-oil operation is shown in Figure 9. No coking residues were found and only a fine, light greyish layer was formed by deposition.
-
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Figure 9: The burner head when viewed from within the combustion chamber. Two spark ignition electrodes can be observed on the right hand side. Flame tunnel emissions which have been measured with the exhaust gas probes are presented in Table IV.
Table ZV. Combustion tunnel emissions.
From Table IV can be observed that the emissions from bio-fuel oil combustion are slightly higher compared to the emissions from diesel combustion. TUXBINE DEVELOPMENT A Deutz gas turbine has been retrofitted to allow for the feeding and combustion of bio-fuel oil in the turbine. This 80 kW(e1) gas turbine is characterized by a nominal rotational speed of 50000 rpm and a pressure ratio of 23. A number of test runs have been executed with the turbine rig and a steady operation could be sustained for hours. A summary of the emission measurements of the gas turbine is presented in Table V.
Table V. Turbine emission results.
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FUTURE PROSPECTS Results from the development activities lead to the following route of commercialisation: At present, bio-fuel oil for co-fuing purposes can be produced at commercial terms. (2) Power generation from bio-he1 oil in turbines or diesel engines is expected within 2 years. (1)
ACKNOWLEDGEMENT The development program of the rotating cone technology has been supported amongst others by: - EC DGXII, FAIR program, - NOVEM, EWAB program.
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CFD for the Modelling of Entrainment in Fluidised Bed Fast Pyrolysis of Biomass H.Gerhauser, S.C. Generalis, R.A. Hague, A.V. Bridgwater Bio-Energy Research Group, Aston University, Birmingham, B4 7ET, United Kingdom
ABSTRACT Bio-oil generated by fluidised bed fast pyrolysis of biomass holds considerable promise towards reducing dependence on fossil fuels. One problem still holding back progress is the efficient removal of char by entrainment from the reactor, which is impossible to model with traditional empirical correlations. A new model combining computational fluid dynamics (CFD), implemented in the commercial software package CFX, and particle entrainment theory has therefore been developed and found to give good qualitative agreement with experimental entrainment data. The CFD model uses a multiphase Eulerian-Eulerian technique for the bulk of the fluidised bed and individual particle tracking in the freeboard. Furthermore, the predictions of the CFD calculations for a number of important fluidised bed characteristics were checked, namely bubble diameter and velocity, terminal particle velocity and bed expansion. These were found to correspond closely to observations. Finally, several design modifications were tested on a laboratory cold flow model, which showed a significant improvement in char entrainment, and successfully modelled by computational simulation. INTRODUCTION Biomass is a significant renewable resource, which will contribute towards replacing fossil fuels as they become depleted and less environmentally acceptable. It is particularly important to find alternatives to crude oil, which still enjoys a virtual monopoly as a transportation fuel. Presently, there are two main routes for obtaining biomass based liquid substitutes, biological (ethanol, rape oil) and thennochemical conversion (pyrolysis oil and hydrocarbons by Fischer-Tropsch synthesis from gasification). Biological methods are established technology, but, suffer from several drawbacks, notably a low overall conversion efficiency and the need to divert food production. Fast pyrolysis on the other hand can deal with a larger variety of organic feedstocks allowing far greater volumes and the use of materials that are now wasted. The technology, however, is still young and there are a number of areas, such as liquid stability, heating value and viscosity, where improvements and further research are desirable. Aston University is one of the institutions engaged in that work and is investigating fast pyrolysis in fluidised bed reactors. The specific sub-task dealt with in this paper is the removal of char, which is an unavoidable by-product that has been found to catalyse reactions that lower the pyrolysis oil yield. It is achieved by selective and enhanced entrainment from the bed followed by separation in a series of cyclones. 1281
Entrainment may be defined as the carryover of ejected particles, while selective entrainment of finer or less dense particles is often referred to as elutriation. In most industrial processes, neither entrainment nor elutriation are desirable', which is in sharp contrast to this particular application. Consequently, there is very little research aimed specifically at enhancing the selective removal of less dense material from fluidised beds. Most research on entrainment is based on dimensional analysis applied to experimental data either with no or very limited consideration of the underlying physics2. Predictions made from these correlations are limited to very simple geometries. They may vary widely even for reactor arrangements close to the experimental conditions they are based on, and are often completely unreliable when conditions are markedly different. In several internal studies they have been found inadequate for entrainment and elutriation predictions in the fluidised bed system under investigation. The problem is too complex to be adequately represented by a small number of ordinary equations that would simply require substitution of a few parameters to obtain the rates of entrainment of the different particle size fractions. Hence more elaborate ways of modelling were considered, namely CFD and the Computational fluid dynamics was chosen, as it discrete element method (DEM)394*5. can be applied to systems with a very large number of particles, where DEM becomes far too computationally expensive. In CFD the flow domain is discretised into a number of grid points for which the fluid flow equations (Navier-Stokes) are then solved numerically. There are two ways to deal with solid particles in CFD.In particle tracking Newton's second law is applied to the particles and the interaction with the fluid is obtained via drag terms. Interactions between the particles are usually neglected rendering this approach inappropriate for modelling high particle concentrations such as occur inside fluidised beds. To deal with large particle concentrations a two-phase approach is normally taken that considers locally averaged values for the particle phase, rather than individual particles.
CFD A N D ENTRAINMENT THEORY INTRODUCTION
The computational fluid mechanics calculations were implemented using the commercial software package CFX, which allows considerable flexibility through User Fortran routines. The bulk of the particles in the fluidised bed was modelled using the standard Eulerian-Eulerian two-phase model that treats the particles not as single entities but rather as a locally averaged second phase. Ejected particles, however, were tracked individually. The starting positions and velocities of the particles were obtained by assuming a bubble ejection mechanism based on experimental observations, Eulerian CFD results and the entrainment literature'. TWO-PHASE FLUID BED MODEL A Eulerian-Eulerian method is applied that uses locally averaged variables for the particle and fluid phases. The optimum scale of averaging was obtained by running the same simulation with successively better grid and time step resolutions until the numerical solution became sufficiently insensitive to further increases in the number of grid points and time steps. The equations resulting from the averaging process are of a similar form to the standard Navier-Stokes equations. Each phase has its own set of equations that are
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solved individually, but linked through inter-phase transfer terms. They were first derived by Gidaspow6. The equations for continuity and momentum are as follows. Continuity:
Momentum:
N
N.
IS
where r refers to the volume fraction, the subscripts a and and B to body forces
to the respective phases
In order to obtain the interphase momentum transfer term in regions with a gas volume fraction less than 0.8 the Ergun equation’ is adapted:
In regions with lower particle volume ftactions a modified form of the single particle drag correlation is used instead:
The drag coefficient Cd is obtained from the particle Reynolds number:
Re-’(1+O. 15Re0.687)
Cd
= 24
or
=0.44
for Re< 1000
for Re> 1000
The equations are solved numerically using a pressure correction technique, commonly known as SIMPLE and a semi-implicit method to handle the strong coupling between the two phases similar to the IPSA algorithm by Spalding’.
’,
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The inlet boundaries are specified using the standard Dirichlet boundary condition with a constant velocity. The outlets are pressure boundaries. The initial condition is no flow and particles filling the bed up to a specified height (that is the volume fraction of particles is specified). PARTICLE TRACKING
Particle tracking in the freeboard is performed with the help of Newton's second law taking account of the drag force exerted by the gas on the particles:
The effect of the particle movement on the fluid phase, and interactions between particles are neglected. This is justified because the number of particles in the freeboard is low. Particles are allowed to bounce off walls, where a coefficient of restitution of 0.8 is applied, which is based on the elasticity of sand and metal collisions but can be varied considerably (from 0.5 to 0.9) without affecting the entrainment results noticeably. ENTRAINMENT MECHANISM
When a fluid is allowed to rise through a loose bed of particles, there will be a pressure drop across the bed that acts as a lifting force on the particles. At the point of incipient fluidisation this pressure drop has become large enough to balance the weight of the particles forming the bed. Further fluid flow then percolates through the bed in the form of bubbles. The eruption of these bubbles at the bed surface is responsible for the ejection of particles of all size classes into the freeboard. Very fine particles may even be entrained without the assistance of a bubble, if 'their terminal falling velocity is below the superficial gas velocity in both the bed and the freeboard. Linking particle tracking and the two-phase fluidised bed model is not straightforward, because individual particle velocities are subsumed in the local averages and therefore not directly available from the fluidised bed model. Consequently, the initial particle velocities for particle tracking have to be obtained from what the fluidised bed model does give, namely the extent of bed expansion, the position and velocity of the bubble, gas velocities and local averages for particle velocities. Several models in the literature' use the bubble velocity to estimate the initial particle velocity without distinguishing between different particle sizes and densities, that is heavier and bigger particles are given the same starting velocity as lighter and smaller ones. This wrong assumption implemented in particle tracking would indicate that large particles reach a greater height than small ones, completely contrary to experimental observation. The particle tracking is therefore divided into two phases. In the first the acceleration within the fast gas jets that are found in and just above the erupting bubbles is simulated and in the second the accelerated particles are followed through the freeboard. In practice this means an average gas jet velocity and the time and distance the particle spends within that jet are estimated from the fluidised bed model, and then put into the particle tracking model. This gives an initial particle velocity for tracking
1284
the particles through the freeboard. The initial position for this second phase of particle tracking is also obtained from the fluidised bed model. A final question that needs to be answered is the number of particles drawn into the gas jets during bubble eruptions. Based on the literature’ and experimental observations undertaken during the course of this study, it is assumed that a layer with a thickness equal to the mean particle diameter in the bed is involved in the ejection process. From the surface exposed to a particular gas jet the total mass of involved particles is then calculated. Size classes for char and sand particles are allocated the same percentage of the ejected particle mass as in the bed as a whole. The total entrainment is calculated by dividing it up into small classes (for example char between 180 and 300 microns, ejected from a bubble between 2 and 3 cm in diameter) and then summing the individual contributions. At this point it may also be clarified that most two-phase Eulerian calculations were performed in two dimensions, while most particle tracking operations were done in three dimensions. This was due to excessive computing times being required by the two-phase calculations. SUMMARY O F PROCEDURE FOR OBTAINING ENTRAINMENT
In order to make the procedure used for obtaining entrainment clearer, it is he1pfi.d to restate it in a simplified and shortened form: 1 . A Eulerian two phase calculation is performed in two dimensions giving locally averaged velocities and volume fractions for the particle and gashapour phases. Furthermore, bed expansion and bubble characteristics are obtained. 2. It is assumed that particles are predominantly entrained when a bubble hits the surface of the bed. Furthermore, it is assumed that the volume of particles ejected at that point is equal to the volume of a layer of average particle size thickness. The size distribution of the ejected particles is assumed to be the same as in the bed. 3. Individual particles are tracked through a constant flow field obtained by extrapolating the 2-D field at the moment of bubble eruption into 3-D. The number of particles leaving the flow field are counted by the computer. A very large number of particles needs to be tracked (of the order of 5000) to obtain acceptable results. EXPERIMENTAL WORK
Entrainment and important fluidised bed parameters were measured for a 1 kg/h and a 5 kg/h fast pyrolysis reactor. In addition, a cold-flow model of the 1 kg/h rig was built to study fluidisation aspects that were difficult to obtain from the pyrolysis reactor. The cold-flow model was subsequently modified to validate the model’s capability to deal with changes in the reactor geometry. The cold flow rig is illustrated in Figure 1. PARTICLE SIZE ANALYSIS
The bed material is sand, which is mixed with char produced by the pyrolysis of wood. Three methods were used for size analysis: sieving, microscopic examination and laser diffraction. All the methods gave similar values for the particle diameters of sand and char, though microscopic examination showed that many char particles had a needle form and therefore one significantly longer dimension that could not be captured by the other two methods. The different particles are illustrated in Figure 2 and the initial size distributions for sand and char in Figures 3 and 4. 1285
Fig. 1 Cold fluid bed model
Unsieved wood nominal size 3 mm. Sieved sand between 500 and 600 microns Fig. 2 Char, wood and sand particles (Scale to the left of each picture, units in mm).
1286
455
422
~500
Dlameter In M k m
Fig.3 Initial size distribution for sand (size classes based on British Standard).
Fig.4 Initial size distribution for char (also based on BS for sieving). C F D RESULTS COMPARED WITH EXPERIMENT The qualitative fit of the CFD work and experiment is illustrated in Figure 5 . As can be seen clearly, that fit is very good.
Fig.5 Left, actual bubble. Right, simulated bubble.
1287
TERMINAL FALLING VELOCITY
The terminal velocity that particles reach when in Free fall is an important particle characteristic. It is also fairly straightforward to measure and was therefore used to check the accuracy and validity of the particle tracking model. The measurement technique is entrainment of a small sample of a particular size class in a narrow tube taking the average of the band within which 90% of the small sample are entrained. In the computational modelling of the particle tracks the upper end of the class size is used as the particle diameter and the particles are started from rest and from a falling velocity greater than their terminal velocity. For sand the tit is very good (see Figure 6). The measured velocity is within the band indicated by particle tracking. For char (see Figure 7) the picture is slightly more complicated. For smaller diameters, the measured falling velocity is above and not within the range given by the particle tracking. The microscopic examination explains the reason.
, -
61 $ 5 c
V---c'MrY +Particle
Tracking
0
7
<251
<355
422
<500
400
-
1
C699
S i fraction in microns
Fig. 6 Terminal velocity of sand.
eE
2.5
e - 2
.-b
-$8 1.51 cp
.S 0.5
€
//-:Typ
-w- Particle Tracking
c o Size fraction in microns
Fig. 7 Terminal velocity of char.
While sand particles approximate spheres fairly closely, char particles are long and thin cylinders. Furthermore, the smaller the char particle the greater the aspect ratio. This may be explained by assuming that the char particles fall apart along the grain rather than across it. As sieving measures the diameter of the thin cylinders and not their length, the effective diameter is underestimated, particularly so for smaller char particles. Consequently, the diameters used in the particle tracking undertaken to 1288
estimate entrainment had to be adjusted. This was done by shifting the measured curve to the right until the measured and computed value for the size class corresponded. BUBBLE VELOCITY
This variable was measured experimentally using a video camera. Several bubbles were analysed for each velocity and an average taken. Likewise a number of bubbles generated by the CFD software for the same conditions were examined. These two values are also compared (see Figure 8) to a widely employed correlation for bubble velocity by Davidson and Harrison": US = (u-~,f)+ 0.7 1 1
(sD)'.~
0.9
0
.~-
1
-~
2
~
~
.--., 3
-.--.
-~
4
--
.~
5
Bubble diameter in cm Fig. 8 Bubble velocity. The CFD results are in much closer agreement with the experimental values than the above correlation, in particular for larger bubble diameters, where the bubble velocity is significantly underpredicted. There is quite a degree of variation between different bubbles. Even for exactly the same conditions they may vary considerably in terms of velocity, size and shape. This dynamic and irregular behaviour is very well captured in the CFD analysis and cannot possibly be obtained by a simple correlation that is only capable of giving an average. BUBBLE DIAMETER In comparison to the bubble velocity the diameter shows a much greater range of variation. Correlations therefore usually only give the maximum diameter rather than an average that would be dificult to apply. A correlation by Mori and Wen" has been chosen for purposes of comparison (see Figure 9) with experimental and CFD results:
1289
T---7-
0
0.05
0.1
0.15
0.2
0.25
-7
0.3
Superficial less minimum fluidising velocity mls
Fig. 9 Maximum bubble diameter.
The correlation by Mori and Wen predicts the actual maximum diameter fairly well, though it underpredicts for higher minimum fluidising velocities. The CFD calculations give lower values for the maximum at all velocities. One reason is the small number of bubbles that could be modelled within the available computing time in comparison to the number available for experimental study. Furthermore, at higher velocities convergence problems were encountered again reducing the number of bubbles available for comparison. BED EXPANSION
The surface of the fluidised bed is not static, but rather subject to dynamic movement. For validation of the model the maximum height reached by the surface is therefore plotted (Figure 10). 14 I
0.05
0.1 0.15
0.2
0.25 0.3 0.35 0.4
Superficial velocity in mls
Fig. I0 Bed expansion.
ENTRAINMENT
From Figures 6 and 7 (shown earlier), which illustrate the terminal falling velocities of the sand and char particles, it can be seen that virtually all char particles have a lower terminal velocity than the vast majority of sand particles. Based on this consideration, separation can therefore be effected within a gas stream that has a superficial velocity in between those of sand and char, that is about 1.5 to 2 m/s. Based on this
1290
consideration a number of modifications was implemented both in the CFD modelling work (see Figure 1 1-14) and experimentally. Figure 15 illustrates sand entrainment on the original cold flow rig. While the trend for sand entrainment estimated from the CFD calculations is correct, on a quantitative basis the under-prediction is significant, namely by a factor of up to three for the lower superficial velocities. The main source of error is the entrainment mechanism. The next graph (Figure 16) shows the percentage of char entrained out of the original and modified designs. At the start of the experiment the bed consists of 95% sand and 5% char. The superficial velocity is 0.45 m/s.
Fig. I I Simulations of different fluid bed configurations
Fig. I2 Simulation of original 1 kglh cold rig showing a contour plot of vapour speed.
1291
Fig. 13 Simulation of the modified 5 kglh reactor
Fig. 14 Modified cold flow rig showing contour plots of speed along cross sections of the reactor 1.4
s 1.2
0
0.2
0.4
0.6
0.8
Superficial less minimum fluidising velocity mls
Fig. 15 Sand entrainment on the original cold flow rig.
1292
Y 0
20
40
60
80
Time in minutes
Fig. 16 Comparison of mixed sand and char entrainment in original and modified cold flow rig design. The graph indicates that the initial entrainment rate for the modified design is much greater than for the original design. The analysis of the entrained material also showed that in the improved design ten times fewer sand particles were entrained and that their distribution was skewed with no sand particles greater than 500 microns. On a qualitative basis that is exactly what the CFD calculations predict (see Figure 17). The figure shows that the majority of large char particles and nearly all smaller char particles that are ejected are entrained, while sand particles with a diameter of 800 microns fall back. Finally, a comment on the different geometries should be made. For the two reactor sizes investigated so far, a simple freeboard widening was found to reduce both overall entrainment and the selectivity of that entrainment with respect to char, while a narrowing (either directly, or by using an insert) proved beneficial on both counts. It should be stressed, however, that for different scales, bed materials or other changes a new calculation is required. CONCLUSIONS Classical empirical correlations have been found to be incapable of satisfactorily modelling entrainment in fluidised bed fast pyrolysis reactors. A new model that combines CFD (for Eulerian two-phase modelling and individual particle tracking) with an ejection mechanism has been developed. This has been successfilly validated against experimental data for a number of different geometries with agreements within 10% at higher entrainment rates and within 50% at low entrainment rates. The poorer agreement at low entrainment rates is believed to be due to the fact that some of the less frequent bed conditions are not captured in the CFD model. For predictions relating to fluidised bed characteristics, the agreement is much better than for entrainment, and is typically within 5% of the measured variables. Using the cold flow rig an improvement in the initial char entrainment rate by a factor of two and enhanced separation from sand by a factor of ten have been achieved and correctly predicted by the entrainment model.
1293
Fia. 17 Parl:icle Tracks, Left large char particles (800 microns); Centre - sand (800 microns); Right - small char particles (400 microns) All particles close to maximum ejection height and velocity.
-
Y
ACKNOWLEDGEMENTS
The work reported in this paper is based on contributions to EC Contract JOR3-CT970197 and the authors would like to acknowledge the support of the European Commission JOULE Programme. NOMENCLATURE
distributor area over number of orifices body forces interphase momentum transfer term drag coefficient bubble diameter particle diameter bed diameter interphase non drag forces drag force gravitational acceleration 1294
m2 N/m3 kg/m3s m m m N/m3 N m/s2
mass flow between phases
kg/m3s m
bed height number of phases pressure volume fi-action Reynolds Number time speed velocity superficial velocity minimum fluidising velocity bubble velocity fluid velocity relative to particle subscripts for the two phases density viscosity
Pa
S
mls m/s m/s m/s m/s mls
kg/m3 kg/ms
REFERENCES I. 2.
3. 4.
5. 6. 7.
8. 9. 10.
11.
Fung A.S., Hamdullahpur F. (1993) Effect of bubble coalescence on entrainment in gas fluidized beds. Powder Technology, 77, pp 25 1-265 Zenz F.A. et al (1958) A Theoretical-Empirical Approach to the Mechanism of Particle Entrainment from Fluidized Beds. American Institute of Chemical EngineeringJournal, 4, pp. 472-479 Witt P.J. et al (1998) A numerical model for predicting bubble formation in a 3D fluidised bed. Applied Mathematical Modelling, 22, pp. 1071 1080 Tsuji Y.et al(1993) Discrete particle simulation of two-dimensional fluidised bed. Powder Technology,77, pp. 79-87 Hoomans B.P.B et a1 (1996) Discrete particle simulation of bubble and slug formation in a two-dimensional gas-fluidised bed: A hard-sphere approach. Chemical Engineering Science, 51, pp. 99- I 1 8 Gidaspow D. (1994) Multiphase Flow and Fluidization.Academic Press Ergun S. (I 952) Fluid Flow through packed columns. Chem. Eng. Prog., 48, pp. 89-94 Patankar S.V., Spalding D.B.(1972) A calculation procedure for heat, mass and momentum transfer in three-dimensional parabolic flows. International Journal of Heat and Mass Transfer, 15, pp. 1787-I 806 Spalding D.B. (I 980) Numerical computation of multiphase flow and heat transfer. In: Recent Advances in Numerical Methods in Fluid Mechanics (Ed. by C. Taylor & K. Morgan), Pineridge, Swansea, pp. 139-I68 Davidson J.F., Harrison D. (1963) FfuidisedParticles. Cambridge University Press More S. and Wen C.Y. (1975) American Institute of Chemical Engineering Journal, 21, p, 117
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1295
Modelling, Scale-up and Demonstration of a Vacuum Pyrolysis Reactor J. Yang, D. Blanchette, B. de Caumia and C. Roy Institut Pyrovac Inc., 333 Franquet Street, Sainte-Foy, Quebec GI P 4C7, Canada
ABSTRACT A novel reactor configurationhas been developed in our laboratory which addresses the heat transfer limitations usually encountered in vacuum pyrolysis technology. In order to scale-up this reactor to an industrial scale, a systematic study on the heat transfer, the chemical reactions and the movement of the bed of particles inside the reactor has been carried out over the last ten years. Two different configurations of moving and stirred bed pilot units have been used to scale-up a continuous feed vacuum pyrolysis reactor, in accordance with the principle of similarity. A dynamic model for the reactor scale-up was developed, which includes heat transfer, chemical kinetics and particle flow mechanisms. Based on the results of the experimental and theoretical studies, an industrial vacuum pyrolysis reactor, 14.6m long and 2.2 m in diameter, has been constructed and operated. The operation of the pyrolysis reactor has been successful, with the reactor capacity reaching the predicted feed rate of 3000 kg/h on a biomass feedstock anhydrous basis.
INTRODUCTION Over the last ten years, the vacuum pyrolysis process has been developed at the bench and pilot scales at the UniversitC Lava1 and Pyrovac Institute Inc. in Qdbec, Canada. This thermal decomposition process enables a large variety of solid and semi-liquid wastes to be transformed into useful products. Vacuum pyrolysis is typically carried out at a temperature of 400-500°C and a total pressure of 2-20 kPa [l]. These vacuum conditions allow the pyrolysis products to be rapidly withdrawn from the hot reaction chamber, thus preserving the primary fragments originating from the thermal decomposition reactions. A limitation of vacuum pyrolysis technology is heat transfer. Previous studies have shown that the rate of heat transfer is essentially the rate limiting step for pyrolysis reactions [2]. Conventional pyrolysis reactors such as multiple hearth furnaces, rotary kilns and screw type reactors exhibit overall heat transfer coefficients ranging from 10 to 60 W.m”K-’ [3], depending on the type of feedstock handled. The low thermal conductivity of the feedstock materials partially explains why the heat transfer fluxes are low.
1296
The new vacuum pyrolysis industrial scale reactor, described herein, uses an indirect heating system with eutectic molten salts flowing inside the heating plates. The feedstock moving on top of these plates is heated both by conduction and radiation mechanisms. In addition, the novel transport system in the reactor efficiently agitates the feedstock, thus greatly improving the heat transfer. The process is commercialized under the trademark of PyrocyclingTM. It is an owned technology [4]. In a properly designed industrial scale reactor, feedstock conversion is acheved at a certain throughput capacity. In order to scale-up the reactor, heat and mass transport phenomena must be studied. This includes heat transfer phenomena, feedstock conversion kinetics and the movement of particles inside the reactor. In this work, both experimental and theoretical studies were carried out to investigate these phenomena. Two different configurations of moving and stirred bed reactors, the batch scale rotative and a continuous feed Process Development Unit (PDU), have been used to generate the data in accordance with the principle of similarity. A dynamic model to scale-up the reactor was then tested.
DESCRIPTION OF THE PYROCYCLING~~ PROCESS The process flow chart of the new vacuum pyrolysis industrial scale reactor is presented in Figure 1. From the feedstock storage facility, separated from the plant building, the feedstock is loaded in a hopper and carried along by extracting screws into a pneumatic conveying system up to a cyclone, located on the roof of the plant building. The feedstock flows down by gravity from the cyclone into a vibrating screen separator, providing a granulometry range between US Sieves 0.5 and 40 mesh. The feedstock with the specified particle size then sluices out through the vacuum feeding system, consisting of two rotary valves set in line and between which the auxiliary liquid-ring vacuum pump suction line is connected. A positive pressure drop is maintained over the second rotary valve (between the vacuum pump suction and the reactor) in order to avoid the escape of pyrolysis gas from the reactor. The feedstock is then fed into the 14.6 m long and 2.2 m diameter reactor by screw conveyors. BIOMASS
VACUUM PUMP
1
SAL
I
HEATEI
4
CHARCOAL ms m m E
STORAGE
WATEWOlL MIXTURE m SEPARATION
Figure 1 Process Flow Chart of the Vacuum Pyrolysis PyrocyclingTMTechnology 1297
Inside the reactor, the biomass is pyrolyzed on two heating plates, whle being carried along and mixed by a raking system. The 10.4 m long by 1.2 m wide heating plates are composed of tubes in which molten salts circulate as a heat c'&er medium. The reactor is also heated by electric tracers installed on the exterior shell. The total pressure in the reactor is maintained at 20 P a . At the outlet of the reactor, the solid product is moved out of the reactor by a threestage water-cooled screw conveyor system. The pyrolytic vapours and gaseous products are withdrawn from the reactor via a 600 mm header to a two-stage condensation system, where the heavy bio-oils are condensed in the first packed tower while a mixture of light bio-oil and pyrolytic water is condensed in the second packed tower. The bio-oil/water mixture is afterwards separated. The bio-oil fractions from both packed towers are mixed together and stored (either directly in 1 m3polyethylene containers or into a 50 000 L tank). The aqueous phase is neutralized with liquid caustic prior to being disposed of (temporary solution) into the municipal sewer system. The non condensable gas drawn up from the second packed tower by the main vacuum pump is further cooled past the pump, compressed to 170 kPa pressure and fired in combination with natural gas into a molten salt heater to provide the energy required for the pyrolysis reaction. From the molten salt 9 m3 storage tank, which is equipped with heating elements as well as the piping loop circuit, the molten salt is pumped to the fiunace to be heated, circulates then into the reactor heating plates and flows back into the tank. The total surface area occupied by the process equipment, excluding the biomass and product storage rooms, is equivalent to 850 m2split into two floors.
HEAT TRANSFER FUNDAMENTALS When the biomass feedstock particles enter the vacuum pyrolysis reactor, they are immediately heated by the heating plate and pushed forward by the blades. The temperature of the feedstock increases quickly while the feedstock is moving towards the outlet of the reactor. It thus produces a temperature distribution along the longitudinal direction of the reactor, with the lowest temperature at the entrance position and the highest temperature at the reactor outlet. At a certain position in the reactor the feedstock temperature reaches a chemical decomposition temperature. While the feedstock flows forward and its temperature fiuther increases, the pyrolysis reactions extend and convert the feedstock into vapours, gases and solid residues. At the outlet of the reactor, the maximum temperature is reached and the final conversion is achieved. During the pyrolysis process, the final conversion mainly depends on three phenomena: the heat transfer fiom the reactor to the feedstock, the feedstock movement in the reactor and the kinetics of pyrolysis reactions. The heat transfer rate determines the rate of temperature increase of the feedstock. The feedstock movement behaviour determines the residence time of the feedstock particles in the reactor. In tum the heating rate and the residence time control the quantity of energy transferred and thus the temperature distribution throughout the feedstock in the reactor. Once the temperature distribution is known, the kinetic behaviour of the feedstock determines the final conversion at the reactor outlet. A systematic study has been carried out in our laboratories over the last decade on the above three phenomena with the objective to design industrial reactors (See Figure 2). Three theoretical and semi-theoretical models have been developed. The heat transfer model, namely the "Surface Renewal Model" has been developed to predict the 1298
heat transfer coefficient in the reactor. The feedstock movement model, known as the "Single Blade Volume Output Model" was developed to predict the average residence time of the feedstock in the reactor. The kinetic model, here a "Parallel Kinetic Model", predicts the conversion of the feedstock as a function of temperature. The combination of these three models gives birth to a dynamic model to scale-up the reactor by allowing the prediction of the final conversion of the feedstock.
Figure 2 Chemical and Transport Phenomena Occurring in a Vacuum Pyrolysis Reactor and Their Role in Reactor Design. THE "SURFACE RENEWAL MODEL"
The proposed Surface Renewal Model is based on the heat transfer model that Schliinder [ 5 ] developed for the indirect contact vacuum drying of solid particles. In Schliinder's model, particle movement due to the agitation was described as a series of time periods during which the bed of particles is assumed to be static, followed by a rapid mixing of the particles. Thus the average heat transfer coefficient during the static period can be calculated using the Penetration Theory, using the following equation:
k=l
where a b e d is the overall heat transfer coefficient in the bed, and B k and a k represent the amount and the heat transfer coefficient of fraction k. The variables, Bk and ak,are calculated by the following equations:
1299
In equation (3), p, C, and R are thermophysical properties lmked to the nature of the feedstock. Thus, the main parameters whch govern the reactor design are the surfaceto-particle contact heat transfer coefficient, aluS, and the bottom layer renewal Since these two parameters are directly influenced by the reactor design, efficiency, /I. they allow a relationship to be established between the reactor design and the overall heat transfer throughout the bed of particles. THE "SINGLE BLADE VOLUME OUTPUT" MODEL
In a moving and stirred bed reactor, the most important parameters representing the feedstock flowing behavior are the residence time t,, the velocity of feedstock movement vbed and the feedstock bed thickness in the reactor hbd. The "Single Blade Volume Output" model proposed in this work is based on the assumption that at the outlet of the reactor, the amount of feedstock accumulating in front of each agitation blade is determined by the value of hbedusing a power expression: 'blade = * (hbed 1" (4) where c and n are constants. Their value depends on the frictional properties of the material and the configuration of the agitation blades. The value of c and n can be determined experimentally. If a steady flow in the reactor is assumed, the mass flow in the reactor, (Qin), must be equal to the mass flow out of the reactor:
Qin
= Pbed
'blade ' 'blade
'blade
*
(5)
'-
Lehain
where Qin is the feedstock feed rate, pbd is the bulk density of the feedstock, nblade and vb[(,&are the number and the velocity of the agitation blades and L&in is the length of the chain to which the agitating blades are attached. vblade and hbed are determined by combining Eq. ( 5 ) and Eq.(6). The velocity of feedstock movement V b d can then be calculated by the following expression which is derived from a mass balance around the reactor: 'bed
=
'blade
hbed
' 'blade ' Wbed
*
'blade
' Lchain
Using equation (6), the residence time of the feedstock in the reactor can then be determined by equation (7):
t, =-Lbed
(7)
'bed
where Lbed is the total length of the two heating plates.
1300
THE "PARALLEL IUNETIC MODEL" For many materials that can be treated in the vacuum pyrolysis reactor, the weight loss of the pyrolyzed material can be expressed by the sum of the conversion of each group component in the material:
xi is the ratio of each component. For where m is expressed as (rn-mo)l(m~,,-mo). example, in the case of wood bark, the major components are assumed to be lignin, cellulose, hemicellulose and water. The conversion of each component can be described by the Arrhenius Equation:
dC . -- - Zip'' R T ( 1 J
dt
Ci)"'
(9)
where E, Z and n are the activation energy, the pre-exponential factor and the reaction order. In this work, the parameters were determined by thermogravimetricanalysis. A DYNAMIC MODEL TO SCALE-UP THE VACUUM PYROCYCLING~~ PYROLYSIS REACTOR
Based on the above three models, a dynamic model to scale-up the vacuum pyrolysis process was developed, which correlates the temperature and the mass of feedstock at any position on top of the heating plates inside the reactor, as a function of heat transfer, particle flow and pyrolysis kinetics phenomena. The energy conservation in the reactor is the foundation of the model. It assumes i) steady flow, ii) one dimensional temperature variation and, iii) feedstock thermal properties vary as a h c t i o n of temperature T.
The term on the left side of the equation represents the flow of internal energy in and out of the system, where m is determined by the lunetic Equations (1-9). At is determined by the total number of nodes generated, N, and the feedstock residence time in the reactor, tl whch can be calculated by equations 4-7. The first term on the right side represents the heat transfer from the bottom heating plates; Tw, is the temperature of the heating plate and abed is the heat transfer coefficient which is determined by the heat transfer equations (Eq. 1-3). The second term is the radiation heat transfer contribution from the reactor wall. The last term represents the kinetic energy released during the pyrolysis reaction, which is assumed to be proportional to the rate of pyrolysis reaction (Eq.8-9).
1301
EXPERIMENTAL
Two different configurations of moving and stirred bed reactors, the batch scale rotative and PDU reactors, have been used for the tests in accordance with the principle of similarity. THE I'BATCH SCALE ROTATIVE REACTOR"
The batch scale rotative reactor was developed and used as a tool to validate the new heat transfer model [4,61. The batch reactor consists of a well insulated tank which contains molten salt and is equipped with heating elements in order to be able to heat the salt to the set point temperature. A second well insulated tank can be placed in the salt bath. The feedstock enters the reactor through the feedpipe. The same agitation blades as those found in the industrial reactor are used. The stirring mechanism transports and agitates the feedstock in a circular manner. The center of the reservoir is kept free of feedstock by a scraping mechanism. The diameter of the feedstock tank is 107 cm and the effective heat transfer area is 0.82 m2 [6]. Both cold and hot tests were carried out in the batch scale rotative reactor. In the cold tests, the bottom layer renewal efficiency, j.3 (see Eq. 1-3), has been measured as a function of the agitation speed [7]. During the hot tests, a certain quantity of feedstock is dropped into the reactor when the reactor has been first heated up to 520°C by the molten salt. Several thermocouples record the temperature change in the molten salt and in the bed of particles. The heat transfer coeecient from the molten salt and the feedstock in the reactor is determined as a function of the temperature variation [6]. THE PDU VACUUM PYROLYSIS REACTOR
The PDU vacuum pyrolysis reactor is a semi-continuous horizontal pilot plant reactor 3 m long with a diameter of 0.6 m and a throughput capacity of about 50 - 200 kg/h, depending on the feedstock treated. The configuration of the PDU reactor is almost the same as that of the industrial reactor, except that the PDU has smaller agitation blades. Two types of tests have been conducted with this reactor, the cold and the hot runs.In the cold tests, the particle flow behaviour is studied by a stimulus-response techmque, under different agitation speeds and feed rates. The hot tests enable the conversion to be determined as a function of the feed rate and the agitation speed. THERMOGRAUMETRIC ANAL YSIS
Thermogravimetric analysis (TGA) has been used to study the kinetics of feedstock pyrolysis. In this work, a Seiko 220 TGIDTA thermal balance system was used. The sample mass was 10.0 fO.l rng.The heating rate was 10 "Clmin. For each test, the nitrogen gas was first introduced in the timace to remove the air. Then the pressure in the fUrnace was lowered by a vacuum pump to 13 kPa absolute. THE FEEDSTOCK
The feedstock was air-dry softwood bark obtained from a wood shredding plant. It was composed of approximately 31% vlv of balsam fir (Abies balsamea), 55% vlv white spruce (Picea glauca) and 14% v/v black spruce (Picea mariana). The moisture content of the feedstock at the inlet of the reactor averaged 10.0% by wt. The
1302
proximate analysis gave: 74.8 volatile matter (V.M.),22.3% fmed carbon and 2.9% ash. The sample particle size is comprised between 0.5 and 40 mesh US sieves.
RESULTS AND DISCUSSION THE TGA RESULTS The furnace of a TGA instrument can be considered as a small scale reactor, in which a linear heating rate and a uniform temperature and pressure environment exist. The weight loss measured under such conditions is only controlled by the hetics of the feedstock pyrolysis. Figure 3 shows the weight loss (the TG curve) and the rate of weight loss (the DTG curve) of softwood bark thermal decomposition measured by TGA. As shown in Figure 3, the softwood bark pyrolysis starts at 160 "C. At 358°C it reaches the maximum rate of weight loss. At 400 "C, one observes a turning point, as the extensive pyrolysis phase is terminated and a slow pyrolysis process continues until 900°C. At 900"C, the pyrolysis is completely terminated. The weight loss then reaches 74% of the initial weight.
T8 7
100 6 80
5
40
2 1
20
' 0
0
-1
0
200
400
600 Temperature "C
800
1000
Figure 3 Weight Loss (the TG curve) and the Rate of Weight Loss (the DTG curve) of Softwood Bark as Measured by TGA Similar to most lignocellulosic materials, the s o h o o d bark used in this work basically contains three organic chemical components: cellulose, hemicellulose and lignin. Using the "Parallel Kinetic Model" (Eq. 8-9) to fit the DTG curve in Figure 3 (the solid line), the activation energy E, the reactor order n and the ratio of each of the three components can be determined. The simulated DTG curve is drawn in Figure 3 (the dashed line). It shows a very good agreement with the measured DTG for the temperature range of 30°C to 520°C. As the temperature in the vacuum pyrolysis 1303
reactor is less than 52OoC,the simulation is satisfying. The kinetic parameters obtained from the DTG curve simulation are presented in Table 1. The s u m of the ratio of each component gives 95%, while the other 5% weight loss occurs in the temperature range from 520°C to 9OO"C, which is not of interest in this study. Table 1. Kinetic Parameters Obtained by Fitting the TGA Results Components
E (kJ/mol)
Z (l/min)
n
Cellulose Hemicellulose Lignin Water
180 80 30 60
1 .ox 1 oi5 1 . 5 81~ 0' 2.4~10'
1 1 1 2
5.0~10~
Component Content (%) 35 20 38 2
THE HEAT TRANSFER COEFFICIENT IN THE ROTATIVE REACTOR
The heat transfer coeficient in the batch rotative reactor has been determined both theoretically and experimentally. As indicated by the heat transfer model (Eq. 1-3), the heat transfer coefficient between the reactor and the feedstock is a function of the thermophysical properties of the feedstock C, , p and A, the contact heat transfer and the bottom layer renewal efficiency p. When the feedstock material coefficient G~ is known, the thermal properties and a,,,scan be determined [ 5 ] . The heat transfer coefficient becomes a function of p and tmu. The thermal properties of the softwood bark are given in Table 3. The values were determined experimentally in this work [8]. Table 2. Physical and Thermal Properties of the Softwood Bark Feedstock Bulk density p, kg/m3 Specific heat capacity C,, Jkg K Thermal conductivity A, W/m K Contact heat transfer coefficient G , W/mz K
250 2720 0.074 640
The main parameter influencing the value of /3 is the speed of the agitation blades [7]. Through cold tests in the rotative reactor, the experimental values of p can be determined. The results are shown in Table 3. Table 3. Bottom Layer Renewal Efficiencies for the Circular Agitation Device in the Batch Rotative Reactor Tuming speed (rpm) 10 19 26 38
Characteristic time, tmir (s) 6.1 3.1 2.0 1.6
Bottom layer renewal efficiency, /3 0.56 0.51 0.46 0.47
When the thermal properties and the p values listed in Tables 2 and 3 are introduced into the heat transfer equations, the heat transfer coefficients can be determined as a 1304
function of the agitation speed (see Figure 4, the solid line). Figure 4 shows a quick a increase with the agitation speed. The maximum a is 105 W/mzK whch is reached at fm&=2.1.
c
120
. +
U
Q
2
40
1
I
+ Measured
3 4- Predicted
2o 0 0
5
10
15
Characteristictime of the agitation device tmlx ( 5 )
Figure 4 Predicted and Measured Heat Transfer Coefficients in the Rotative Batch Reactor.
The heat transfer coefficient in the rotative reactor has been determined experimentally. The tests measured the temperatures of the feedstock in the reactor and that of the heating plate. Then the heat transfer coefficient is calculated based on the variation of the feedstock internal energy and the temperature difference between the feedstock and the heating plate [ 6 ] . The maximum a is reached at tmk=2.1, and corresponds to 180 W/mZK. This value incorporates both the heat transfer from the heating plate and from the reactor wall. It has been estimated that the heat transfer by radiation fiom the reactor wall provides approximately 1/3 of the overall heat transfer. Figure 4 presents the measured heat transfer fiom the heating plate to the feedstock as a function of tmir,which is equal to 2/3 of the measured overall heat transfer coefficient (see the diamond marks). Comparison between the experimental and the theoretical results indicates that the experimental measurements and the model predictions are in good agreement. Even though the predicted a seems to be closer to the experimental value at h g h turning speeds, the difference between the experimental and theoretical values is less than 21%. This difference may be due to the fact that the model only considers the heat transfer from the heating plate to the feedstock. In reality, there is also some convective heat transfer caused by the evolving pyrolysis vapor products. THE MOVEMENT OF PARTICLES IN THE PDUREACTOR
As shown by the "Single Blade Volume Output Model", the correlation between the single blade volume V b d and the bed height hbed is the key to determine the feedstock flowing behaviour in the vacuum pyrolysis reactor. Figure 5 presents the Vbed measured as a function of hbed. The square marks in Figure 5 are obtained using the large blades designed for the industrial reactor. The round marks are llnked to the small blades in
1305
the PDU reactor. Through curve fitting, the c and n in equation ( 5 ) were determined to be 110 and 1.2 (the solid line) and 23 and 1.2 (the dashed line) for the large and the small blades, respectively. When c and n are known, the feedstock residence time t, can be calculated using the Equations 4-7. Figure 6 shows the predicted residence time in the industrial reactor. The design parameters used for the calculations are presented in Table 4. Figure 4 shows that t, significantly varies with the speed of the agitation blades and moderately varies with the feedstock feed rates. 140
120 100 A
i,$A
80 60
1
>
;
4oo
i
40
P Industrial D 4U reactor
20
200
0
0 0
2
4
6
10
6
Thickness of the bed of particles in the reactor hb.d (em)
Figure 5 Observed Values of Vbed as a Function of bed
b
0
500
1000
1500
2000
2500
3000
3500
4000
Q (kglh)
Figure 6 Predicted Residence Time in the Industrial Reactor as a Function of the Velocity of the Agitation Blades. 1306
Table 4. Design Parameters for the PDU Industrial Scale Reactors Parameters wbed wblade Lchain
Lbed
PDU 0.4m 6.4~ 10 -3 m 5.24 m 4.04 m 96
nblude
Industrial 0.96 m 0.038 rn 26.7 m 20 m 385
To validate the "Single Blade Volume Output Model", the feedstock residence time was also studied experimentally by using the stimulus-response techmque. Using scrap tires as feedstock, the residence time in the PDU reactor has been measured, which showed a very good agreement with the model prediction [8]. The experiments for measuring the tr of wood bark feedstock in the PDU reactor is in preparation. The results will be presented later.
BIOMASS CONVERSION IN AN INDUSTRIAL SCALE PYROLYSIS REACTOR The mathematical model to scale-up the vacuum pyrolysis reactor predicts the temperature distribution in the reactor and the final conversion of the feedstock under varying feed rates and agitation speeds, once the reactor size and configuration are fixed. The dimension of the industrial scale reactor is presented in Table 4. The values of wbed and Lbed presented are the effective dimension for heat transfer study, which are smaller than the true values (the true values of WM and L w are 2m and 20.8m, respectively). When one uses as inputs the thermal properties (Table 3), the flow of particles characteristics (Table 2 and Figure 5) and kinetic (Table 1) parameters in the mathematical model (equations 1-10), the temperature and the conversion distributions can be predicted. The speed of the agitation blades is assumed to be 25 d m i n in the calculation. The temperature of the heating plate is 500°C. The average tmkin the reactor is about 11 s. The heat transfer coefficient is thus estimated to be 100 W/m2K, whch includes both the heat transfer from the heating plate and from the reactor wall. The feedstock moisture is assumed to be 10%. To calculate the conversion behaviour, it is assumed that the conversion is 100% when the volatile matter content in the residual wood charcoal is 35%. As depicted in Figure 7,the temperature in the reactor increases more rapidly at the beginning due to the large temperature difference between the reactor and the feedstock. The maximum temperature is reached at the end of the reactor, before the product flowing out of the reactor. In Figure 7 the temperature distribution changes greatly with different feed rates: the lower the feed rate, the higher the final temperature of the feedstock. Figure 8 illustrates the predicted feedstock conversion in the reactor as a fiction of the reactor zone and the feed rate. The conversion is low at X < 2 m In the case of a feed rate of 1000 kg/h, a rapid conversion is observed between 2 c: X < 10. Here the conversion reaches 100% even before reaclung the end of the reactor. At feed rates of 2000-3500 kg/h, the conversion occurs much later, at X=5 m The final conversion is lower and reaches loo%, 96% and 92% at the reactor outlet, for feed rates of 2000,3000 and 3500 kgh, respectively. Figure 9 illustrates the predicted final temperature (dashed line) and the f m l conversion (solid line) in the reactor, as a function of the feed rate. It shows that when the feed rate is 5 2600 kg/h, the final temperature is higher than 450°C, and the final conversion reaches 100%.
1307
When the feed rate is higher, the final conversion decreases due to a decrease of the charcoal outlet temperature.
500
e-
400
Es 300 0)
n
!i
-1000 kglh 2000 kglh 3000 kglh
I- 200
--
100
5
0
10
15
-
20
Distance X (rn)
Figure 7 Predicted Temperature Distributions in the Reactor as a Function of the Distance along the Reactor and the Feed Rate. 1.2
-1000 kglh - 2000 kglh
1
0.8
.-E0
0.6
0
0.4
I! Q) > E
0
0.2 0 0
5
10
15
-0.21
Figure 8 : Predicted Feedstock Conversion in the Reactor as a Function of the Distance along the Reactor and the Feed Rate.
1308
540
r-
-I 100%
520
95% 500
90%
1
%
480
.EI!? 0
>
c
85%
C
s
Predicted T PredictedConversion Observed Conversion 420 1000
70% 1500
2000
2500
3000
Feed Rate (kglh)
Figure 9 The Predicted Feedstock Temperature Outlet and Ultimate Conversion, and the Observed Conversion in the Industrial Reactor, as a Function of Various Feed Rates.
Figure 9 shows the final conversion measured as a h c t i o n of the feed rate during experimental runs in the industrial PyrocyclingTMreactor described earlier (diamond marks). When the feed rates are less than or equal to 2600 kgh, the final conversion is always higher than 100%. When feed rate reaches 3200 kgh, the conversion is 92%. Comparing the measured and the predicted conversions shown in Fig.9, a very similar trend can be observed for the conversion. For most of the points in Fig.9, the predicted conversion is very close to the experimental results, although the predicted values are a little bit higher. The most llkely reasons for the slight discrepancies are: i) the mathematical model considers no mass transfer resistance when the pyrolysis vapour or gas products are evolving in the reactor and ii), there may be vapour products recondensing in the reactor particularly when the vapour products are flowing from the hot zone to the colder zone, a factor whch is not taken into account in the mathematical model. CONCLUSIONS
The objective of scaling-up a vacuum pyrolysis reactor is to achieve the desired capacity and conversion of feedstock by determining the dimension of the reactor. The feedstock conversion in a vacuum pyrolysis reactor mainly depends on three factors: the heat transfer coefficient from the reactor to the feedstock; the residence time of the feedstock in the reactor and the kinetics of the feedstock pyrolysis reactions. The heat transfer coefficient and the residence time determine the quantity of energy transferred and thus the temperature distribution throughout the feedstock in the reactor. The temperature distribution and the kinetics determine the fmal conversion acheved at the reactor outlet.
1309
A heat transfer model, namely the "Surface Renewal Model" has been developed. It predicted that the heat transfer coefficient a vacuum pyrolysis reactor fed with biomass is 80-140 W/m2 K, depending on the process conditions. A model for the movement of particles, known as the "Single Blade Volume Output Model" was developed to predict the average residence time of the feedstock in the reactor. It showed that the residence time of the biomass in a 20 m long reactor ranges between 5 and 9 min, with a speed of the agitation blades varying from 15-25 dmin. The biomass conversion starts at 180 "C. At 520 "C, 95% conversion is achieved. A kinetic model, called the "Parallel Kinetic Model", simulated the TGA results and enabled the determination of the activation energy and the fraction of each component in the feedstock. Finally the combination of these three models gives birth to a dynamic model to scale-up the reactor by allowing the prediction of the final conversion of the feedstock. An industrial vacuum pyrolysis reactor, 14.6 m long and 2.2 m in diameter, has been constructed and operated. The operation of the pyrolysis reactor has been successful, with the reactor capacity reaching the predicted feed rate of 3000 k g h on a biomass feedstock anhydrous basis. Using the dynamic model, the conversion of the biomass feedstock in the reactor has been predicted as a function of the biomass feed rates. Comparison of the predicted and measured conversions in the reactor shows a fairly good agreement. NOTATION A B C CP
d E
hbed
i
i
k
Lbed Lchuin
m n nblude
N Qin
41,
R T t
tmix
tr
At vbed vblade
surface area, m2 the amount of the agitated fraction k of the first layer of particles conversion specific heat capacity, J.kg-'K-' diameter, m activation energy k~ moP thickness of the feedstock bed of particles in the reactor, m grid location on the longitude axis index of each component in the pyrolysis material index for the first layer of particles having different residence times the length of the heating plate, m the length of the chain to which the agitation blades are installed, m mass, kg the reaction order number of agitation blades installed in the reactor the total number of nodes generated in numerical calculations the feedstock feed rate, kgah-' heat of reaction, W kg-' ideal gas constant, J K ' mol-' temperature, K time, s characteristic time of the agitation device, s residence time of feedstock in the reactor, s the time interval for the feedstock flowing from one grid to the next, s velocity of feedstock in the reactor, m s-' the velocity of the agitation blade, m s-'
1310
vb/,,de w&d Wb/&
X Z
the volume of the particle accumulating in front of an agitation blade, m3 the width of the heating plate, m the width of the agitation blade, m longitudinal distance along the reactor, m pre-exponential factor, d'
GREEK LETTERS a
P /1 P 0
heat transfer coefficient, W.m-*K-' bottom layer renewal efficiency thermal conductivity, W . m k ' density, kg.m-3 Stephan-Boltzmann constant, W.m-2K-4
REFERENCES 1. Roy, C., Labrecque, B. and de Caumia, B. (1990) Recycling of scrap tires to oil and carbon black by vacuum pyrolysis. Resour. Conserv. Recycl., Vol. 4, pp. 203-213. 2. Yang, J., Tanguy, P. A. and Roy, C. (1995) Heat transfer, mass transfer and kmetics study of the vacuum pyrolysis of a large used tire particle. Chem. Eng. Sci. Vol. 50, NO. 12, pp. 1909-1922. 3. Perry, R.H. and Green, D. (1984) Perry's Chemical Engineers' Handbook, Sixth edition, McGraw-Hill Book Co., New York, pp. 11.48 11.49. 4. Roy, C., D. Blanchette and B. de Caumia. Horizontal Moving and Stirred Bed Reactor. Canadian Patent Claim Number 2,196,841. US Patent Number 8,811,172. International Patent Number 98,902,153.0. 5 . Schlhder, E. U. (1984) Heat transfer to packed and stirred beds from immersed bodies. Chem. Eng. Process., Vol. 18, pp. 31-53. 6. C. Malendorna, Yang, J. and C. Roy. (2000) Determination of the Overall Heat Transfer Coefficient in a Vacuum Pyrolysis Moving and Stirred Bed Reactor. Chem. Eng. Res. Des. Vol. 78, pp. 633-642. 7. Roy, C., D. Blanchette, L. Korving, J. Yang and B. de Caumia. Developments in Thermochemical Biomass Conversion. A.V. Bridgwater and D.G.B. Boocock, Eds. Blackie Academic and Professional, London, UK, 1997 pp. 35 1-367. 8. M. Gupta, J. Yang, and C. Roy, Density of Softwood Bark and Softwood Char: Procedural Calibration and Measurement by Water Soaking and Kerosene Immersion Method. Manuscript in preparation. 9. Yang, J., C. Roy, M. Gupta and X. Roy, Study of Tire Particle Flowing Behaviour in an Agitated and Moving Bed Reactor Manuscript in preparation.
-
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Thermal efficiency of the HTU@ Process for Biomass Liquefaction F. Goudnaan", B. van de Beldd,F.R. Boerefijnb, G.M. Bos', J.E. Naber", S. van der Walb, and J.A. Zeevalkink' a Biofuel B. V., Rendorppark 30, 1963 AM Heemskerk, The Netherlands b
d
Jacobs Comprimo Nederland, Postbox 58026, 1040 HA Amsterdam,, The Netherlands TNO-MEP , Postbox 342, 7300 AH Apeldoorn, The Netherlands Biomass Technology Group (BTG) B. V. , Postbox 21 7, 7500 AE Enschede, The Netherlands
ABSTRACT The development of the HTU@Process is now well under way. A pilot plant with an intake capacity of some 10 k g h (drybasis) is in operation. Biomass is converted by treatment in liquid water at temperatures from 300 to 350°C and pressures from 100 to 180 bar. The product is 'Biocrude', a heavy organic liquid with a lower heating value of 30-35 MJkg. A case study is presented for the HTU conversion of sugar beet pulp (130 kton/a on a dry basis). A process description is given, and the heat economy is discussed. For the generation of process heat some external fuel is required (2% of the heating value of the feedstock). The thermal eficiency is defined as the ratio of heating values of biocrude product and feedstock plus external heat input It has a theoretical maximum of 79%. The process designed here has a thermal efficiency of 75%. The total capital expediture is 30 M$. At a zero cost of feedstock the price of the biocrude product is $ 153/ton, or $4.6/GJ. A number of items for M e r improvement of the thermal efficiency is discussed. It requires a systematic exergy analysis. However, any further gain in heat economy has to be traded off against increased capital expenditure and reduced operability.
INTRODUCTION Hydrothermal liquefaction is a process for the conversion of biomass into an organic product oil. It has been widely studied in the early 1980's (see the recent extensive literature survey in [l]). Starting in 1983 Shell Research performed process development based on experiments in autoclaves and a continuous bench scale unit. This resulted in a conceptual design for the HTU process including cost estimate [2,3]. The work was discontinued in 1988. 1312
In 1997 Stork Engineers and Contractors (now Jacobs Comprimo Nederland), Shell Netherlands, TNO-MEP, BTG and Biofuel formed a consortium for further development and commercialisation of the HTU process. With financial support fiom the Dutch Government (the EET programme) a 5.5 M$ project is under way fiom 1 November 1997 to 1 September 2000. It encompasses, experiments in autoclaves, product research, process design, and the construction and operation of a pilot plant with a throughput of 10-20 k g h (db), [See list of abbreviations at the end of this paper] see Figure 1. The business plan shows favourable commercial prospects [4]. A feasibility study is being carried out which aims at the construction of a demonstration plant, capacity 10-20 ktoda (db), in 2002.
Fig. I HTU pilot plant, located at TNO-MEP, Apeldoorn, The Netherlands.
The HTUprocess In the HTU process the feedstock is treated in liquid water at temperatures ranging from 300 to 360°C. The pressure is 100-180 bar and the residence time 5-20 minutes. Oxygen is removed fiom the biomass, mainly as COz. This results in a product designated as “biocrude” with an oxygen content as low as 10-18 %w. The biocrude is not miscible with water and it has a relatively high heating value (LHV is 30-35 MJkg). Its main application as such is power generation, e.g., co-combustion in a coalfired power station. Alternatively it can be upgraded by catalytic hydrodeoxygenation. Scouting experiments have demonstrated that in this way a diesel fuel with excellent properties can be produced. By its nature, the HTU process is very well suited for the conversion of wet biomass feedstocks. These are generally low in price, such as residues from agriculture, forestry, food processing industry, etc. However, at a longer term, large-scale processing of crops fiom energy plantations is considered as well.
1313
Thermal efficiency Based on the results of the process development work the design of a commercial-scale HTU process unit is now under way. Heat integration and energy economy are of prime importance to keep the processing cost low. A useful criterion for the energy economy is the thermal efficiency. This paper discusses a case study of such a process design. HTU PROCESS DESCRIPTION CHEMISTRY OF OXYGEN REMOVAL
In chemical terms the key to biomass liquefaction is the removal of oxygen. Biomass contains typically 40-45%w (DAF basis) of oxygen. Oxygen removal increases the heating value and it leads to a product with more hydrocarbon-like properties causing it to be immiscible with water. In thermochemical liquefaction the oxygen is either removed as water or as carbon dioxide. Removal as water leads ultimately to carbon as the remaining product (examples: charcoal manufacture, pyrolysis). Removal of carbon dioxide tends to leave a product with a higher WC ratio and therefore a higher LHV. At HTU reaction conditions a significant proportion of the oxygen is removed as C 0 2 .The decarboxylation selectivity is defined as the ratio (oxygen removed as COz) : (total of oxygen removed as C 0 2 and H20). Table 1 shows decarboxylation selectivities for a number of biomass HTU conversion cases with different feedstocks and reaction conditions. It appears that the selectivity is almost constant.
Table 1 Deoxygenation selectivity. % of oxygen ending up in product fraction
Product fraction Biocrude Light organic product (OD) c02
H20 Decarboxylation selectivity
case 1
case2
case3
case 4
14 20 35 31
21 12 36 31
15 9 37 38
12 9 40 40
0,53
0,54
0,49
0,50
The chemical reaction mechanism is not quite understood [ 5 ] . Apart from water and carbon dioxide the reaction products are: Biocrude. This is the main product. It has an oxygen content of 12-20 %w and the atomic WC ratio is 1.0-1.3. Average molecular weight is about 300. Organics dissolved (OD). This is the organic fraction dissolved in the process water after cooling the reaction mixture to ambient temperature. It consists mainly of acetic acid, acetone, and components like methyl-cyclopentenone and hydroxypyridine. Gases. Apart from carbon dioxide, some carbon monoxide is formed, and traces of methane and hydrogen.
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As a feedstock for the case under study sugar beet pulp is selected. Its properties are given in Table 2. Table 2 Sugar beet pulp feedstock composition SBP Feed Composition, %w DAF feed, C H
dry
wet
45.75 6.42 46.17 1.66
42.32 5.94 42.71 1.54 7.49
100.00 20.35
100.00 22.00
9.31 1.31 9.40 0.34 1.65 78.00 100.00
DAF
0
N Minerals Water Total %w of wet feed
100.00
A typical product distribution is shown in Table 3. With the SBP feedstock less biocrude and more OD are produced than with wood [3].
Table 3 Product distribution and LHV balance Product DAF feed Products Biocrude Gas,
%w on feed (DAF)
c02
co
CH4 H2
Organics Dissolved (OD) Water Total products
LHV, MJkg
LHV, MJkgfeed
100.00
17.5
17.5
4l.27 24.25 1.20 0.26 0.03 12.35 20.64 100.00
33.32
13.75 0.00 0.12 0.13 0.03 3.21 0.00 17.25
0 10 50 121 26 0
PHASE EQUILIBRIA Phase equilibrium data are essential for process design. Prof. J. de Swaan Arons and Dr. Chongli Zhong (Presently Professor at Beijing University of Chemical Technology.) of Delft University of Technology are developing a thermodynamic model, which also predicts phase equilibria. Main uncertainties are in the modeling of the biocrude properties. Experimental verification of phase equilibrium data with autoclave and pilot plant equipment is ongoing. An alternative method is to realise that carbon dioxide and water constitute over 90 mole % of the reaction mixture. Literature data [6] for the C02-H20phase system are used, and for the distribution of biocrude and OD over the phases, distribution coefficients are derived from experiments and literature.
1315
Thirdly, flowsheeting programmes like AspenPlus@ also provide data on thermodynamics and phase equilibria. Here again, biocrude and OD have to be introduced, e.g., as pseudo-components. THERMOCHEMICAL DA TA Specijic heat The specific heats of various biomass samples were measured. The specific heat of biocrude was estimated from a group contribution method. LHV data Table 3 gives for each of the products its LHV and its contribution to the total LHV of the product mixture. The LHV of biocrude is given as 33.32 MJkg. This value is derived from experimental determination and from calculation on the basis of elemental composition. The two methods show a good agreement. The Table shows that the LHV of the biocrude is 79% of the LHV input of the feedstock. The LHV of the combustible by-products (gases and OD) is some 20% of the LHV input. It is, in principle, available for the generation of process heat. Standard enthalpy of reaction In theory the reaction enthalpy at 298 K can be derived from the LHV balance as given in Table 3. This method gives AH,,298= 17.25 - 17.50 = - 0.25 MJkg feed (DAF). However, the accuracy of the most important LHV data (feed and biocrude) is considered to be at best *lo%. So, this would yield m r , 2 9 8 = -0.25 f 3 MJkg (DAF). On this basis, it cannot be determined whether the reaction is endothermic or exothermic. For the present purpose we derive the reaction enthalpy from autoclave experiments. In a standard experiment, a 2 litre autoclave is filled with some 350 g of water and heated to 330°C. Then an amount of SBP feed is injected with a specially constructed “high-pressure syringe”. The injection of the (cold) feed makes the temperature drop sharply, and then it increases again, the main driving force being the heat content of the autoclave wall. The standard experiment was then repeated with the injection of water in stead of SBP feed. By comparing the course of the temperature of the contents of the autoclave for the blank and the standard experiments, one can estimate the reaction enthalpy. The temperature curves are shown in Figure 2. Calculations take into account the heat capacity of the wall and the contents of the autoclave, enthalpy of vaporisation, and other relevant factors. The resulting values of the reaction enthalpy are: Initially endothermic, = + 0.40 MJkg (db), and then exothermic, m r , 2 9 8 = -1.00 MJkg.
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...-
.Blanc
-HTU
330"
Time Fig. 2 Temperature recordings in blank and HTU experiments. Indications for the reaction enthalpy are also derived from the continuous pilot installations, both at Shell Research (1984-1986) and the present one at TNO-MEP. Here one measures the steady state electric energy supply required for heating the reaction mixture to the desired temperature. By comparing this with a blank experiment where only water is heated one concludes that the overall reaction heat effect is exothermic, and its value is consistent with the result from the autoclave work.
CONSIDERATIONS FOR PROCESS DESIGN In t h s section we discuss a case study for the design of a commercial scale HTU installation, based on the data given above. The intake capacity of SBP feed is 130 ktoda, or 4.51 kg/s db assuming 8OOOWy. A simplified process flow scheme for the main process loop is shown in Figure 3. Reactor conditions On the basis of autoclave and pilot plant experiments a HTU reactor temperature of 330°C is selected, with a liquid residence time of 10 minutes. The product distribution in Table 3 pertains to these conditions. The reaction pressure does not have much of an effect, provided the water is in the liquid phase. The decisive factor in choosing the reactor pressure is the extent of vapour phase formation in the reactor. A lower pressure causes a larger proportion of the reaction mixture to evaporate. As a consequence more heat has to be supplied in E l to maintain the reactor temperature at the desired level. This tends to be disadvantageous in view of the temperature level of this heat supply. A concomitant advantage of the lower pressure is the lower equipment cost in all of the high-pressure section of the process. These factors, along with the possible advantage to have a larger part of the reaction mixture in the vapour phase, have to be traded off. In the present case a pressure of 165 bar was chosen as the optimum. 1317
system boundary
3.9 A
350
2309
260
I
E2
El
I
Gas
.I R ]
9 4sl
1.0 260.
Biocrude
260
330
E3
4
17.4
Recycle water ~
a
Water
Legend: 80 = temperature in OC 5.8 = Enthalpy, MW
-
Fig. 3 Flow scheme for main HTU process loop.
Feed pump A major accomplishment of the present project is the cooperation with a pump manufacturer who has constructed a pump for the pilot plant installation, for bringing the feedstock to the HTU process pressure of 100-200 bar in one step. On the basis of this, and their experience with commercial-size pumps, the manufacturer has made a design and cost quotation for a commercial-scalepump. In the process scheme of Fig. 3 the SBP feedstock is first heated to 8OoC, using heat available in the process (see below). For heat economy reasons it would be desirable to choose this temperature considerably higher. For the time being this is not done because it would require pressurized equipment upstream of the feed pump. The pump(s) P bring the feedstock to the HTU pressure.
Preheating of the feedstock At the outlet of the pump the feedstock flows through a duct with a diameter of, say, 200 mm. Supply of heat through the wall of the duct would require virtually all of the heat transport to occur by conduction. The penetration of heat to the core of the "solid" cylinder of feedstock would require many hours. It is therefore not a practical solution. Hence, alternative ways of preheating are called for. Almost all feedstocks of interest have such properties that they do not flow readily through either pipes or shell of a conventional heat exchanger without distribution problems. After a number of alternatives had been considered it was concluded that the most adequate way of preheating is direct injection of hot water. In Fig. 3 this is done in vessel V where a stream of hot recycle water is mixed with the feed. A special way of contacting achieves efficient heat and mass transfer. It has been verified in our experiments that at 200-250°C biomass feedstocks of different origin, including wood chips, are softened and form a paste-like substance [7,8]. The 1318
residence time required for this is 2-100 minutes, depending on the temperature. Some carbon dioxide forms as well, but owing to the combination of high pressure and relatively low temperature, it remains in solution and does not form a separate vapour phase. In our case study it was decided to have the softener vessel V at a temperature of 230OC. The temperature of the recycle water has to be high in order to have the amount of recycle water as small as possible. This is essential to limit the size of all equipment in the hgh-pressure section. For this reason the recycle water is heated to 350OC in heater E3. HTU reactor
In the reactor R the HTU reactions take place. At the prevailing temperature and pressure a vapour phase forms. The design of the reactor vessel is such that it allows for sufficient stageing to prevent backmixing, and for efficient separation of the vapour phase. Its volume is in the order of 40 m3. Product cooling and separation The liquid effluent of the reactor is cooled in cooler E2 to 260°C. This is sufficiently low to stop any undesired reactions leading to degradation of the biocrude. In separator section S the two liquid phases, water and biocrude, are separated. Experiments in autoclaves have shown that at this temperature the two liquid phases separate well and that their mutual solubility is sufficiently low to avoid excessive loss or contamination of the product. From the water stream from the separator the recycle water is routed to the recycle pump and then to heater E3.The remainder is the stream “water” leaving the main section in Figure 3.
Downstream process sections The products leaving the main process loop shown in Figure 3 are further handled in sections not shown in the Figure. The gas phase stream is expanded. The possibility of recuperation of work by means of an expansion turbine is under study. Then, the stream is further cooled and treated in such a way that biocrude and water are routed to the liquid streams as much as possible. For the OD contained in the gas stream, it is desired to leave them in the gaseous phase. The handling of the biocrude stream depends on the eventual application of the product. In general it will run off at some 80°C to storage and transport. The water stream is passed through flasher(s), coolers etc. with the aim of having as much as possible of the OD stripped to the gas phase. The remaining liquid water stream is treated by anaerobic digestion. Here biogas is formed which contains the major part of the LHV originally contained in the OD present in the waste water stream. An after-treatment by aerobic digestion is required in most cases to comply with water discharge regulations. Also, the minerals dissolved in the effluent of the water treater are removed to a level corresponding to effluent regulations.
1319
Furnace The gas streams originating from all downstream process section are combined. They contain all of the non condensable gases CO, CH,and H2, and part of the OD. This gas is combined with the biogas and combusted in the process furnace. High-temperature heat is generated and transferred to a hot-oil or high-pressure steam circuit to provide heat to the process. If necessary (see below) additional external fuel is burnt in the furnace so as to match the heat requirement. If regulations require so, a catalytic deNOxer will be installed to reduce the NO, content of the flue gases. For reasons of heat economy the flue gas temperature should be as low as possible and preheating of the combustionair is applied. HEAT ECONOMY AND THERMAL EFFICIENCY
HEAT BALANCE
Basis of calculations For the purpose of a scouting study like the one under review it is convenient to use a simplified method of calculating the enthalpy of process streams. Only the enthalpy of water (liquid and/or vapour), biomass and biocrude are taken into account. The required data are readily available, the calculation is simple and the error made by this simplification does not exceed some 10%. The reference level for H=O is liquid water at 0°C. The enthalpies of the streams entering and leaving the main process loop have been calculated in this way. They are indicated in Fig. 3. The mass balance for the main process loop is obtained in a straightfonward fashion from the product distribution given in Table 3, the recycle water flow rate being obtained fiom an enthalpy balance over vessel V (see below).
Main heaters and coolers The enthalpy balance around softener V considers the sensible heats and the endothermic reaction enthalpy of 4.51 kg/s(db) * 0.4 MJ/kg(db) = 1.80 M W . The flow rate of the recycle water stream is found to be 1.20 * the feed stream (on the basis of water). The enthalpy to be supplied in heater E3 is 10.4 MW. The enthalpy balance around the reactor R deals with the exothermic reaction enthalpy of 4.51 kg/s(db) * -1.0 MJ/kg(db) = - 4.51 M W . The duty of heater El is 17.5 MW. Cooler E2 serves to rapidly cool down the liquid product mixture and to recover high-temperature enthalpy. The outlet temperature of E2 is important for the overall enthalpy economy since it determines the enthalpy content of the streams exiting at the system boundary. By that it determines the required overall heat input to the main process loop shown in Figure 3. For overall enthalpy economy reasons it should be as low as possible. This temperature is found fiom optimisation of the hot oil (or highpressure steam) system that transmits heat from the process h a c e to the main process loop. The temperature of the hot oil leaving E l cannot be lower than 240°C. This stream picks up the heat fiom E2,where it is heated to at most 320°C.With these data
1320
the enthalpy balance over E2 gives the outlet temperature (process side) to be 260°C. The duty of E2 is 14.0 MW.
Enthalpy balance, main process loop From the optimisation of the hot oil system in the preceding paragraph it is found that the process furnace has to supply 13.9 MW to the hot oil. The overall enthalpy balance over the main process loop then reads (see Fig. 3 for system boundary):
Enthalpy In: Stream out of P: Reaction enthalpy Net from timace Total In Enthalpy Out: Gas stream fkom R Biocrude from S Waste water Total Out
5.8 MW 2.7 = (-0.4 + 1.O)* 4.5 1
13.9 22.3 MW
3.9 MW 1.o
17.4 22.3 MW
Overall enthalpy balance, entire process The enthalpy of the streams leaving the main loop is 22.3 MW. Part of it is utilised for heating the feedstock in Ef , duty 4.7 MW, and for providing heat to flashers and strippers in the downstream process sections. The remainder is some 10-15 MW. It is cooled away in air or water coolers. In the downstream process sections some 40% of the OD ends up in the gas to the furnace. The remainder of OD has to be removed in the waste water treating section. Here it is assumed that 80% of the LHV is conserved as LHV of the biogas. In downstream processing some 3% of the biocrude product is lost to gas and water streams. It is assumed that 2% of the biocrude LHV becomes available in the furnace. So, the LHV input to the furnace from the various process streams is (in MJkg feed DAF,see Table 3): Combustible gases, 100% OD via gas, 40% of 3.21 Biogas, 80% of 60% of 3.2 1 Biocrude lost, 2 % of 13.75 Total
0.28 1.28 1.54 0.28
3.38 MJkg DAF feed
This corresponds to 3.38 (MJkg DAF feed) * 4.17 (kg DAF feeds) = 14.09 MW The required enthalpy output of the furnace into the hot oil is 13.9 MW. If we assume the furnace to have an efficiency of 90% the required enthalpy input to the fiunace is 15.4 MW. Availablility from process streams is 14.1 M W . So an external input of (15.4 - 14.1) = I .3 MW is required from additional fuel.
1321
THERMAL EFFICIENCY Definition For the present purpose we define “thermal efficiency” as: tlth =
(LHV of biocrude output) ( L W of feed) + (LHV from external fuel)
* 100%
Input of electric energy is disregarded on the presumption that electricity will be purchased, and as such it figures in the economic analysis. From the figures in Table 3 it follows that the theoretical maximum of the thermal efficiency for the present case is (13.75/17.5) = 78.6 %.
Result In the process design presented above the (LHV from external fuel) is 1.3 MW. The feedstock LHV is 17.5 (MJkg DAF feed) * 4.17 (kg DAF feeds) = 72.98 MW. The LHV of biocrude output is 97% of 13.75 (MJ/kg DAF feed) * 4.17 (kg DAF feeds) = 55.62 MW. So the thermal efficiency found here is: qth=
55.62 72.98 + 1.3
*
100% = 74.9%
When electricity consumption is considered (as fuel consumed in generating the required electricity with an efficiency of 50%), thermal efficiency drops to 72.5 %.
PROCESS ECONOMY For the HTU installation described above an estimate has been made for the total capital expenditure, including, a.o., a 15% development contingency. For the intake capacity of 130 ktoda (db) the capital expenditure is 30 M$. This is significantly lower than what would have obtained from cost estimates based on an earlier process design [3]. This reflects the cost efficiency of the process improvements applied since 1994. In Table 4 an economic analysis of the process is given. A capital charge of 12% per year is assumed. Maintenance, overhead, insurance, etc. are covered with a 6% per year on the capital. Labour costs are based on two operators per shift plus the appropriate day shift personnel. Electricity is purchased for 0.04 $ikWh.External fuel costs 5 $/GJ. In this analysis it is assumed that the feedstock cost is zero. This may be true for certain residues from agriculture, food industry, etc. Equally well the feedstock price may be negative or positive. It is beyond the scope of this paper to discuss t h s any further.
DISCUSSION By successful heat integration in the above process setup a thermal efficiency of 75% has been achieved, which is quite satisfactory. Further improvement may be obtained 1322
by having the outlet temperature of the main process loop at a value lower than 260°C. This requires a different heat integration in the hot oil system. Furthermore a different choice of the temperature of softener vessel V might lead to a higher efficiency. A more detailed exergy analysis would seem required for this. Table 4 Process economy Capacity intake, ktoda (db) Net production biocrude, ktoda Installed capital, M! Yearly cost Capital charge (12%) Maintenance, overhead. etc (6%) Labour Feedstock Electricity External he1 Total $/GJ
130 48.1 30 M$la 3.60 1.80 1.40 0.00 0.39 0.19 7.38
29
4
However, an unproved heat efficiency leads, in general, to more pieces of equipment. Table 4 has shown that about 75% of the process cost is related to the capital investment. Therefore a gain in thermal efficiency has to be carehlly traded off against a higher capital investment and a more complicated operability. A fundamental improvement of the process must be sought in the chemistry. If biocrude could be obtained in a higher yield and with a lower oxygen content, and if the yield of OD would be lower this would raise the thermal efficiency. Supporting studies into the chemistry of the HTU reactions are about to be started. A lower water content of the feedstock would obviously be beneficial since less heat has to be supplied and the equipment size diminishes. T h s has to be balanced against the cost of the feedstock. In the process scheme attention should be paid to lower loss of biocrude and a larger proportion of OD routed to the gas stream. These factors lead directly to a hgher thermal efficiency and a smaller waste water treatment section. It must be stressed that the present case study is to a large extent a simplified one. Examples of simplifications are: 0
0
0
The build-up of minerals, OD, COz etc. in the recycle loop has not been taken into account. The method of calculation of the enthalpy of the process streams is a simplified one. The estimation of capital expenditure was done on the basis of a conceptual design. A more thorough basic engineering study is presently being conducted.
1323
CONCLUSIONS A case study has been made for a conceptual process desin for the HTU conversion of sugar beet pulp with an intake capacity of 130 ktoda (db). The main conclusions are: Optimisation of the heat economy is essential. In the present case some additional external fuel is required for process heat. It amounts to about 2% of the LHV of the feedstock. Thermal efficiency as defined here has a theoretical maximum of 79 % for the present conversion. The process designed here has a thermal efficiency of 75 %. The total capital expediture is 30 M$. At a zero cost of feedstock the cost of producing the biocrude product is $ 153/ton, or $4.6/GJ. A number of items for M e r improvement of the thermal efficiency is discussed. It requires a systematic exergy analysis. However, any fUrther gain in heat economy has to be traded off against increased capital expenditure and reduced operability. ABBREVIATIONS
DAF db LHV OD SBP
Dry and Ash Free dry basis Lower Heating Value at 298 K Organics Dissolved in process water phase. Sugar beet Pulp
REFERENCES 1.
2.
3.
4.
5.
6.
Venderbosch R.H. & Sander C. (2000) Hydroconversion of wet biomass: a review. GAVE rapport 99 19 (order at [email protected]). Annee J.H.J. & Ruyter H.P. (1986) Process for producing hydrocarboncontaining liquids jiom biomass. European Patent 0204354, to Shell Inernationale Research Maatschappij B.V. Goudriaan F., Grinsven P.F.A. van & Naber J.E. (1994) Electricity generation ftom biomass via the HTU process, In: DGMK Conference "Energetische und stofliche Nutzung von Reststoffen und nachwachsenden Rohstoffen ", Velen; DGMK Tagungsbericht 9401 (ISBN 3-928164-70-8) 149-162 Naber J.E., Goudriaan F., Wal S. van der, Zeevalkink J. & Beld B. van de (1999) The HTU process for biomass liquefaction: R&D strategy and potential business development. In: Proc. of the 4Ih Biomass Conference of theAmericas, Oakland (Calg),(Ed. by R.P. Overend & E. Chornet, ISBN 0 08 043019 8), 1, 789-95. Luijkx G.C.A. (1994) Hydrothermal conversion of carbohydrates and related compounds. Ph.D. Thesis, Delft University of Technology. Tadheide K. 8z Franck E.U. (1963) 2. Physik. Chemie (Frankfurt), 394, 387401; also in: Gmelin Handbuch der anorganischen Chemie. 6* edn. (1973) 3 pp 54-7. Verlag Chemie, WeinheindBergstr.
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7. 8.
Mok WS-L. & Antal M.J. (1992) Uncatalysed solvolysis of whole biomass hemicellulose by hot compressed liquid water. Ind. Eng. Chem. Rex , 31, 115761. Jollez P., Chornet E. & Overend R.P. (1993) Steam-aqueous fractionation of sugar cane bagasse: an optimisation study of process conditions at the pilot plant level. In: Advances in Thermochemical biomass conversion (Ed. by A.V. Bridgwater) pp. 1659-69. Blackie, London.
1325
Thermochemical Treatment of Radiata Pine using Hot Compressed Water T. OGI,S. INOUE National Institute for Resources and Environment, 16-3 Onogawa, Tsukuba, Ibaraki 305-8569, Japan Y. YAZAKI Monash University, Dep. Chem. Eng., Clayton, Victoria 3 800, Australia
ABSTRACT: Radiata pine bark contains a large amount of polyflavanoids (radiata tannin) and other valuable materials suitable for feedstock of adhesives and resins. We tried to extract radiata tannin effectively and to partially degrade the bark to water soluble materials using hot compressed water. Radiata pine bark was extracted using hot compressed water either with or without 1 % NaOH at temperatures of 100°C and higher under different pressures for various holding times. When 1% NaOH was added to the compressed water system, the yields of the extractives and polyflavanoids increased remarkably, and the highest yield (3 1.3%) of the extractives was obtained at 140°C and at 10 atm. The hot compressed water extraction requires less than 1/6the heating energy required for the conventional extraction. In order to obtain partially degraded water-soluble materials for producing wood adhesive and/or resin feedstocks, Radiata pine bark was treated with 015% H202 solutions at temperatures higher than 100°C. The highest yields of water-soluble materials (47.4and 47.9%) were obtained with 15% H202 solution at 120°C. The water-soluble fraction was analyzed by GC-MSand other methods. It was found that the water-soluble fraction mainly consisted of low molecular weight acidic organic compounds, and succinic acid and malonic acid were major compounds. These compounds can be used as feedstocks for isocyanate adhesives. 1. INTRODUCTION
In Australia, Radiata pine (Pinus Radiata) is widely silvicultured and a large amount of bark is generated unused. This bark contains large amounts of polyphenols called “polyflavanoids” (radiata tannin) which react with formaldehyde to give high-quality water-proof wood adhesives. An effective tannin extraction is one of the most important key technologies to produce water-proof and structural wood adhesives. It is well-known that the
1326
extractives yield increases with increase of water temperature for bark extraction up to 100°C. Furthermore, extraction of radiata pine bark with aqueous sodium hydroxide (NaOH) gives a higher yield of extractives than extraction with water (1). However, the amount of NaOH used for extraction should be very carefully monitored, because highly alkaline conditions cause the rearrangement of the A-rings structure of polyflavanoids, resulting in the loss of reactivity of polyflavanoids with formaldehyde (2,3). In our laboratory we have studied thermochemical treatment of biomass using hot compressed water and found solid-like biomass wastes were converted to liquidized materials by the release of water soluble compounds inside cells due to thermal rupture of cell around 175°C at 40 atm (4). Applying this liquidization process, we tried to extract radiata tannin from the bark. The bark also contains other materials suitable for resin feedstock. We tried to partially degrade the bark to water-soluble materials using hydrogen peroxide (H202). Hydrogen 'peroxide (H202)yields only water andor oxygen upon decomposition, so it is an ecologically and environmentally desirable pollution control agent. Alkaline Hz02 solutions are also used to chemimechanically bleach ground wood, and to chemically treat pulps. The use of H202 for pulping and bleaching is based on delignification and further degradation of compounds produced during delignification (5). In this paper, we report the results of effective tannin extraction from the bark and degradation of the bark for resin feedstocks using a hot compressed water process (6). 2. EXPERIMENTAL
Radiata pine bark sample The bark samples were collected from a plywood and lumber mill that was using logs from 35-year-old Radiata pine trees at Mt. Gambier in South Australia. The bark samples were dried to 12% moisture content and then ground to pass a 6.2 mm screen in a Wiley mill. The resulting particle size distribution was: 6.2 mm to 1.4 mm (3 1.6%), 1.4 mm to 1.O mm (25.2%), 1.O mm to 0.6 mm (19.8%) and less than 0.6 mm (23.4%). Furthermore, for a comparative study, some of the bark samples were also ground to pass through a 0.6 mm screen. The chemical composition of the Radiata pine bark sample was analyzed according to the Japan Industrial Standard method (7). Elemental and chemical composition of Radiata pine bark are as follows; Composition of Radiata Pine Bark Moisture: 9.3 wt% Organics: 89.9 wt% Elemental Composition (wt%) C: 5 1.9, H: 5.7, Chemical Composition (wt??) Lignin (polyflavanoids) 56.6 12.4 Cellulose 17.8 Hemicellulose
1327
Ash: 0.8 wt% 0: 42.1 *
*by difference
Hot compressed water and 1 % NaOH aqueous extractions at temperatures higher than 100°C All experiments were carried out using a stainless steel autoclave of 300 mL capacity. The bark samples (20 g) and water or 1% aqueous NaOH (based on the weight of the oven-dried bark) solution (70 mL), in which the ratio of bark to solvent was 1:4 (w/w), were added to the autoclave. The autoclave was flushed and then pressurized (4 atm or 10 atm) with nitrogen gas. The temperature was increased to 100°C (or 120, 140, 160, 180OC) and either immediately air-cooled (0 hour holding time) or maintained at a designated temperature for 30 minutes (0.5 hour holding time). After heating, the autoclave was cooled to room temperature and then opened. The bark mixtures were separated into water soluble fraction and insoluble residues by filtering and washing with water. Extractivesyield, Sttasny values and polyJlavanoidsyield The yield of extractives was calculated on the basis of the weight loss of the oven-dried extracted bark, compared with the weight of the oven-dried unextracted bark. The Stiasny reaction was carried out using the freeze-dried extractives. The dried sample (100 mg) was dissolved in water (10 mL), to which 36% formaldehyde solution (2 mL) and 10M HCl (1 mL) were added, and heated under reflux for 30 minutes. The Stiasny precipitates were collected using a no. 3 sintered glass filter and were dried at 105°C for 16 hours. The Stiasny value (original) was calculated on the following equation:
The adjusted Stiasny value was then obtained by dividing the Stiasny value (original) by 1.112,since the value of 1 1 1.2 is the Stiasny value for catechin being used as a model compound for radiata tannin.
Treatment of bark using H202at temperatures higher than 100°C The ground bark (Particle size 1.25 mm to 0.65 mm; 3.0 g) and 0, 5, 10, or 15% H20z Solution (30 mL) was added to an electrically heated 100 mL stainless steel autoclave. Afier this was purged with nitrogen gas, a pressure of 10 atm nitrogen gas was applied and then the autoclave was heated. The temperature was increased to 100°C (or 1 10, 120, 130, 140°C), maintained at the designated temperature for 30 min, and then reduced to room temperature. Fractionation and ana&ses of products aper treatment with H202 Fig. 1 is a schematic diagram showing the procedures for the fractionation of products after the bark was treated with H2O2.The evolved gas was collected in a sample bag and then analyzed using gas chromatography (GC). The GC analysis was done using a WG-100column (Shimadzu GC-8A). The temperatures of the oven, injector and detector were 50°C and 7OoC, respectively. The TCD current was 120 mA. The carrier gas was He at a flow rate of 33 mL min-'. 1328
I
I
Bark
H201treatmenr
Reaction mixture Filtration and washing with HzO
Filtrate and washings Frezze-d ying Water soluble marcrials
H:O insoluble fraction Drying ac IOS'C Water insoluble materials
Fig. 1 Procedures for the Fractionation of products from Radiata pine bark The products obtained from the treatment of the bark with H202 were separated into water-soluble and insoluble fractions by filtering and washing with water (60 mL) on a No.3 glass filter. The water-soluble fraction was diluted with water exactly to 200 mL and was measured for Total Organic Carbon (TOC), pH (using Duotest pH indicator paper), and residual H202 (titrimetric permanganate method), then freeze-dried. The TOC was measured using a TOC meter (Shimadzu TOC-SOOOA). Elemental composition was analyzed by a NA-1500 Elemental Analyzer (Carlow Erba NA-1500). GCMS analysis was done using a Hewlett-Packard 5890 Series 2GC with H p 5971 MS Detector. GC was done on a DB-5 column using He carrier gas at 0.9 mL min-'. Injector and detector temperatures were 320 and 28OoC, respectively. The sample was injected into the GC column at 4OoC, left for 3 min, and then heated to 320°C at 5°C mid'. 3. RESULTS AND DISCUSSION
Hot compressed water and 1 % NaOH aqueous extractions Table 1 shows the results of extraction of Radiata pine under hot compressed water system in two cases; that is, the case using compressed water only and the case with 1% NaOH aqueous solution, together with the conventional reflux extraction case. The extraction using hot compressed water at temperatures higher than 100°C gave lower yields of extractives and polyflavanoids than those extracted using conventional hot water extraction at 100°C under ambient pressure. However, when a small amount of NaOH (1% of the weight) was added to the hot compressed water for the extraction of Radiata pine bark, the extractives yields increased remarkably. The highest extractives yield (3 1.3%) was obtained from the extraction at 140°C and 10 atm with no time at temperature. The second highest one (25.9%) was obtained also at 140°C but 1329
Table 1. Extractives yield, Stiasny value and polyflavanoids yield from the bark Extraction condition Temp.
Pres.
Holding time
Yield
01)
(%)
0
14.4 12.9 14.8 15.6 14.4 11.7 12.2 9.9 13.3 5.2 6.9 3.8
(OC) (am) Compressedwater only 100 100 100 100 140 140 140 140 180 180 180 180
Extractives
4 4 10 10 4 4 10 10 4 4 10 10
0.5 0 0.5 0 0.5 0 0.5 0 0.5 0 0.5
pH
Compressed 1 % NaOH aqueous solution 100 4 0 23.4 100
I00 100 120 140 140 140 160 180 180 180 180 180
4 10 10 10 4 4 10 10 10 4 4 10 10
0.5 0 0.5 0 0 0.5 0 0.5 0 0 0.5 0 0.5
Stiasny value
22.3 21.0 21.8 23.8 25.9 23.9 31.3 13.7 22.9 20.6 13.7 16.8 14.9
Original
Adjusted
Yield
(%)
(%)
(%)
3.8 3.7 3.8 3.8 3.8 3.8 3.8 3.8 3.8 3.7 3.7 3.7
85.0 73.4 82.0 70.2 74.6 32.2 48.4 27.3 43.3 22.5 15.1 27.4
76.4 66.0 73.7 63.1 67. I 29.0 43.5 24.6 38.9 20.2 13.6 24.6
11.0 8.5 10.9 9.8 9.7 3.4 5.3 2.4 5.2 1.1 0.9 0.9
5.6 5.3 5.9 5.0 5.6 5.9 5.3 5.9 5.9 5.9 5.9 4.7 5.9 4.7
85.6 83.9 84.2 80.3 84.0 77.4 72.3 77.2 13.9 71.1 68.9 8.4 25.5 11.1
77.0 75.4 75.7 72.2 75.5 69.6 65.0 69.4 12.5 63.9 62.0 7.6 22.9 10.0
18.0 16.8 15.9 15.8 18.0 18.0 15.8 21.7 I .7 14.6 12.8
Conventional extractionat ambient pressure Extraction condition Extractives Temp.
Reflux
time
(T)
(h)
time
0
1 1
100 100 100 100
100 100
0 0.5 0.5 0.5 0.5
1
Ext. Solvent H20 l%NaOH H2O l%NaOH
1
H20
2 2
H20 l%NaOH
Yield
pH
(%)
13.3 17.4 18.2 22.6 19.6
30.5
Polyflavanoids
4.4 6.5 4.4 6.2 4.4 6.2
1330
Stiasny value
I .o
3.8 I .5
Polytlavanoids
Original
Adjusted
Yield
(%) 88.0
(%)
19.1
(%) 10.5
83.3 88.8 84.1 88.7 82.8
14.9 19.9 15.6 19.8 14.5
13.0 14.5 11.1 15.6 22.1
at 4 atm pressure, while the lowest (13.7%) was obtained after extraction for 30 min at 140°C and 10 atm pressure and after extraction for 30 rnin at 180°C and 4 atm pressure. When treated at lower pressure (4 atm) with no time at temperature, the extractives yields were very similar (23.4% and 25.9%) at 100°C and 140°C and slightly decreased to 20.6% at 180°C. When the bark was treated for 30 rnin at temperature, the extractives yields were less than those of the bark extracted with no time at temperature. The greatest decrease (17.6%) from 31.3% to 13.7% was found in the bark extracted at 10 atm and 140OC between with no time at temperature and with 30 min holding time at temperature. The extractives obtained fiom the extraction at 4 atm pressure with no time at temperature gave the highest adjusted Stiasny value (77.0%) at 100°C which decreased to 69.6% and 62.0% at 140°C and 180"C, respectively. The highest polyflavanoids yield (21.7%) was obtained from the bark extracted at 10 atm and 140°C with no time at temperature, followed by the yield of 18.0% from the extractions at 100°C and 4 atm, 140°C and 4 atm, and 120°C and 10 atm pressure with no time at temperature. On the other hand, the smallest polyflavanoids yield (1 .O%) was obtained after treated for 30 min at 180°C under 4 atm pressure.
Comparisons between the Conventional and the hot compressed exfraction process Ambient pressure has been used for the extraction of Radiata pine bark, although different conditions have been used for the tannin extraction of wattle (Acaciamearnsii) bark. This is largely due to our previous experience that the application of temperatures higher than 100°C decreases the extracted yield from the bark. The yields and Stiasny values of extractives obtained from the bark samples using the conventional extraction method at 100°C are shown in Table 1. At ambient pressure, the extractives yields increased with the increase of extraction time (reflux time) and also with the addition of NaOH. The extractives yields (14.4% and 14.8%) obtained fiom the extraction with hot compressed water at 4 and 10 atm were slightly higher than that (13.3%) obtained fiom the conventional extraction method at ambient pressure with no time at temperature. However, after extraction for 30 min, the reverse result was obtained in which the yields (12.9% and 15.6%) from the extraction with hot compressed water were much lower than that (1 8.2%) for the conventional extraction. In addition, the extractives obtained from the conventional extraction method gave much higher Stiasny values than those of the extractives obtained from the extraction with hot compressed water. When using water only, there is not much advantage in hot compressed water extraction. However, when using 1% NaOH hot compressed solution, the highest extractives yield (31.3%) was obtained at 14OoC, 10 atm, with no time at temperature. This extractive had a high Stiasny value, therefore high polyflavanoids yield (21.7%). Since the highest extractives yield was obtained at 140°C under 10 atm pressure with no time at temperature, two additional extractions were carried out at 120°C and 160°C. The results are included in Table 1. The extractives yields were 23.8% and 22.9% and their adjusted Stiasny values were 75.5% and 63.9% at 120°C and 16OoC, 1331
respectively. Therefore 140°C was the best temperature which gave the highest yield. It is interesting to note that 140°C is the temperature, at which lignin has been reported to soften (8,9, 10). When the polyflavanoids are subjected to highly alkaline conditions, rearrangement of the A-ring structure of the polyflavanoids takes place so that the polyflavanoids are likely to have lost their reactivity resulting in poor glue bond quality (11, 12). Accordingly, the pH of the extractives solution obtained from the 1% aqueous NaOH solution at higher temperatures and pressures is a very critical factor for obtaining wood adhesives of good quality. However, the experimental results showed that the pH's of the 1% NaOH extractives were between 4.7 and 5.9, which being acidic are suitable pH's for the formulation of wood adhesives. The conventional extraction using 1% aqueous NaOH solution gave almost the same yields of extractives (30.5%) and polyflavanoids (22.7%) as the hot compressed extraction. Concerning extractive and polyflavanoids yields, there was not so much difference between conventional and hot compressed water process. However, there is much difference for the energy consumed for extraction process.
Energy consumption and C02 emissionfrom the extractions The conventional process is based on the extraction with hot water, which is 20 times the bark weight, heated under reflux for 30 min and repeated once. However, the hot compressed water extraction uses water amounting to 4 times the bark weight, heated up to 140"C, cooled immediately, and the extraction is not repeated. Based on these conditions and assuming that the extractives yields are 30.5% and 31.3% from the conventional (1% NaOH) and the hot compressed water (1% NaOH) extractions, respectively, the energy required for heating the water from room temperature (20°C) to the designated temperatures (either 100°C or 140°C) was calculated on the basis of the production of It of the tannin extracts from the bark. The specific heat for water is 4.18 [MJ t-I. C ' ] and the amount of bark to obtain ton of tannin extracts was U0.305 [t bark / t tannin extracts] for the conventional and 110.313 [t bark / t tannin extracts] for the hot compressed water extractions, the energy requirement for heating the water to either 100°C or 140°C under these conditions can be calculated according to the following equations; Conventional hot water extraction 1/0.305 x 20 x 4.18 x (100 - 20) x 2 = 43,900 MJ t-' Hot compressed water extraction U0.303 x 4 x 4.18 x (140 - 20) = 6,600 MJ t-' Consequently, the hot compressed water extraction required less than 1/6 the heating energy required for the conventional extraction method. Furthermore, since the conventional extraction method requires additional energy to keep the water at 100°C for 30 minutes twice, the new hot compressed water extraction may be a preferred method. 1332
Assuming that in order to generate l[kWh] of electricity, 0.547[kg-C02] is emitted (13) and 1 [J] is equivalent to 2.8 x lo-' [kWh], the amounts of COz emission for these methods would be; Conventional hot water extraction 43,900 MJ = 12,300 [kWh] = 6,700 [kg-COz] Hot compressed water extraction 6,600 MJ = 1,900 [kWh] = 1,000 [kg-COz] Therefore, assuming that 7,000 t of tannin extracts are produced using the hot compressed water method in a year, the COz emission would be 7,000 t/year, whilst that from the conventional method would be 46,900 t/year. Consequently, since the calculations described above did not consider the energy to keep the temperature at 100 "C during the extraction time (2 x 0.5 hours), the conventional extraction method requires at least six times the amount of energy required for the extraction with hot compressed water at 140 "C with no time at temperature. At the same time, the conventional extraction method emits at least six times the amount of COzthan does the hot compressed water extraction method.
Decomposition of Radktu pine bark treated with H202. The yields of the water-soluble and insoluble materials obtained from Radiata pine bark after treatment with H202solutionsat temperatures higher than 100 "C are summarized in Fig. 2.
loo
r-
A Soluble, reflux - - -
Soluble, 130°C
--
I Soluble, 100°C
Soluble. 140%
- - A - - Insoluble, 120'C
~
-*-
Soluble. IIOT - - 0 - - Insoluble. nflux - - V - - Insoluble, 130°C
80
40
20
I
0
- - Insoluble, 110°C
5
10
15
HI02concentration %
Fig. 2 The yields of water soluble and insoluble materials
1333
At temperatures higher than 100 "C, the yield of water-soluble materials increased with increasing HzOzconcentrationfrom 0 to 15%. The highest yield of water-soluble material (47.4%) was obtained when the bark was treated with 15% HzOzsolution at 120 "C for 30 min. Very similar yields were obtained with the same H20tcharge at 100 "C and 110 "C. Over 130 "C, the yield of the watersoluble materials decreased due to the production of gases such as COzand CO. Various amounts of H2O2remained in the aqueous phase after the bark treatments. The highest concentration of residual HzOz(80%) was obtained when the bark was treated with 15%HzOl under reflux conditions. The decomposition of Hz02was dependent on the temperature applied for the treatments, and there was no H202after treatment at 140 "C. Residual Hz02 limits the use of the water-soluble materials, so it is important to find conditions in which there is little residual HzOz in the aqueous phase. When the bark with HtOZsolution was heated higher than 140 "C, an exothermic reaction took place, which was to vigorous to control, and during the reaction all Hz02 was decomposed and consumed. The suitable degradation condition was the following: temperature 120 "C, 30 min, 10-15% HzOz.
Chemical characterization of products More than half of the bark was obtained as water soluble materials by the hotcompressed H z 0 2 solution process at 120 "C. The water soluble materials were fractionated further into ethyl acetate and aqueous fractions and analyzed by NMR and GC-MS. In Fig.3, the GC-MS spectrum of ethyl acetate fractions and the compounds identified are shown. The compounds with asterisk (*) were also identified in aqueous fractions. The results of the GC-MS for water-soluble materials showed the presence of many low molecular weight organic acidic compounds. The pHs of the water layers after treatments were between 2.8 and 5.0. Kadla et al. studied the reactions of lignin model compounds treated with HZO2at 90 "C.In addition to the reactions of the phenolic representatives, the non-phenolic compounds are shown to react at these higher temperatures via an S N mechanism ~ to ultimately give the corresponding benzoic acids and dimeric ether compound. (14) Krochta et a]. reported the thermochemical degradation of cellulose into organic acids (15). Consequently, it is expected that lignin and / or cellulose contained in the bark react to give acids during the treatments with H202 solutions at temperatures higher than 100 "C.Radiata pine bark mainly consists of cellulose, hemicellulose, and lignin (polyflavanoids). It is a complex feedstock so that many organic compounds, other than acids, might be contained in the water soluble materials. Russel et a]., recognized the formation of aromatic compounds, when cellulose was converted by a thermochemical treatment. (16) The results of NMR and GC-MS analysis also showed phenolic and non-phenolic compounds were contained in the water soluble materials, which maintain the structure of lignin and polyflavanoids after the treatment. The main compounds in the water soluble fraction were succinic acid and malonic acid. It is presumed that these compounds are found by degradation and oxidation of lignin. The solubulization process of the bark consists of a very complicated combination of various reactions such as oxidation, 1334
dehydration, and cyclization. The organic acidic compounds contained in the water-soluble materials can be used in isocyanate adhesives.
b
I& I
2 I 3
1 L
I_lr RmP- 2'
Peak # 1 2 3 4 5 6 7 8 9 12 13
00
25.00 --
30.00
35.00
40.00
-.
Compounds Peak # Compounds Glycolic acid 14 DLmalic acid * Levulinic acid 15 2-Hydroxypentanedioic acid 3-Hydroxy propanoic acid * 17 meso-Tartaric acid * Malonic acid * 18 Xylonic acid * Maleic acid * 19 2,3-Dihydroxy succinic acid * Succinic acid * 22 Tricarballyic acid Methyl succinic acid 23 Aconitic acid 2,3-Dihydroxy propanoic acid * 24 Vannilic acid Fumaric acid 27 3,CDihydroxy knzoic acid Pentanedioic acid 2- Pentanedioic acid
Fig. 3 GC-MS spectrum of water soluble fraction and compounds identified
I335
4. CONCLUSION
The extraction of Radiata pine bark using hot compressed water at temperatures higher than 100 "Cgave lower yields of extractives and polyflavanoids than those extracted using the conventional hot water extraction at 100 "C under ambient pressure. However, when a small amount of NaOH was added to the compressed water, the extractives yields from the bark increased remarkably. The extraction of the Radiata pine bark at 140 "C and 10 atm pressure with no time at temperature provided the highest yields of extractives (31.3%) as well as polyflavanoids (21.7%). The hot compressed water extraction requires less the 1/6 the heating energy for the conventional extraction. Treatment using hot compressed H202solution: When Radiata pine bark was treated with 0-15% H202 solution, the highest yield (47.4%) of water soluble materials was obtained at 120 "C with 15% H202. The water-soluble materials contain many organic acidic compounds in which succinic acid and malonic acid were major compounds. These obtained compounds could be used in wood adhesives, because they contain hydroxyl groups.
ACKNOWLEDGEMENTS: This project was financially supported by the Agency of Industrial Science and Technology (AIST) in Japan and also by the MFP Branch, Department of Industry, Science and Technology (DIST) in Australia. We thank Ms.Y.Fukuda, an assistant staff member at NIRE, for her help during the experiments and Dr.A.F.A.Wallis and Mr.R.Wearne at CSIRO for their GC-MS analysis.
REFERENCES: 1. Yazaki Y. (1985) , Holzforshung, 39,267 2. Sears K.D. Casebier R.L. Hergert H.L. Stout G. H. and McCandish L.E. (1974) J. Org. Chem., 39,3244 3. Yazaki Y. and Aung T. (1989) Holzforshung, 43,281 4. Sawayama S. Inoue S. Yagishita T. Ogi T. and Yokoyama S. (1995) J. Fermentation and Bioengineering, 79,300 5. Hosoya S . (1992) Kizmi-parupu gijutu kyoukaishi, 46,1344 6. Some parts of this paper have been published in the following papers; Inoue S. Asaga M. Ogi T. and Yazaki Y. (1998) Holzforshung, 52,139, and Inoue S.0gi T. and Yazaki Y. (1999) Bull. Chem. SOC.Jpn., 72,2135 7. Migita N. Yonezawa Y. and Kondo T. (1981) Wood Analysis (Chapter lo), Vol2, Kyoritu Shuppan Inc. Tokyo, Japan 1 8. Atack D. (1972) Svensk Papperstidn, 75,89 9. Goring D.A.I. (1963) Pulp Paper Mug. Can. 64,517 10. Hillis W.E. and Rozsa A.N. (1978) Holzforshung, 32,68 11. Yazaki Y. and Aung T. (1989) Holzforshung, 43,281 12. Yazaki Y.and Collins P. (1994) Holz Roh- Werkstoff 52,307 13. NED0 (New Energy Development Organization) report (NEDO-GET 9410-1) (1995) Chap.l1,28p 1336
14. Kadla J. Chang H. and Jameel H. (1997) Holzforshung, 51,428 15. Krochta J.M. Hudson J.S. and Drake C.W.(1984) Biotechnol. Bioeng. Symp., 14,37 16. Russel J.A. Miller R.K. and Molton P.M. (1983) Biomass, 3,43
1337
Chemical conversion of biomass resources to useful chemicals and fuels by supercritical water treatment S . Saka = R. Konishi Department of Socio-Environmental Energy Science, Graduate School of Energy Science, Kyoto University, Yoshih Honmachi, Sakyo-ku, Kyoto 606-8501, Japan
ABSTRACT: A batch-type supercritical water system was used to study the chemical conversion of sugi (CrypforneriujuponicaD. Don) and buna (Fagus crenuru Blume) woods to useful chemicals and fuels. Cellulose, hemicelluloses and milled wood lignin (MWL) as their cell wall components were also studied to understand the reaction behaviors of wood in supercritical water. High performance liquid chromatography (HPLC) and gel permeation chromatography (GPC) have been performed on the reaction mixtures of these samples. The results clearly indicated that the woods can be almost fully liquefied in supercritical water for 5 sec. In addition, the water-soluble fraction was mainly derived from cellulose and hemicelluloses with average molecular weights in a range between 100 and 200, whereas the methanol-soluble fraction derived mainly from the lignin and hydrophobic in nature with an average molecular weight of 530. From these fractions, useful chemicals such as glucose, levoglucosan, 5hydroxymethyl furfural, furfural and some phenolic compounds could be obtained.
INTORODUCTION Biomass resources will become more important in the future as alternatives to fossil resources, whch will be exhausted sooner or later. The features of biomass are renewable, carbon-neutral, and abundant. However, these resources have not been utilized efficiently and their unused portions are wasted in the world. Human beings are, thus, urged to develop efficient utilization technology of biomass, especially cellulosic biomass resources. For the conversion of biomass resources into useful chemicals and bio-energy, three major process types, direct combustion, gaslfication and liquefaction, are effective. Among them, liquefaction for liquid transpiration fuels is, in particular, important because it is difficult to replace them with other energy forms. Thus, a study of the liquefaction of biomass is of interest.”’ There exist two major conversion methods, biochemical and thermochemical, of wood into useful chemicals and liquid fuel. Enzymatic saccarification is one of the biochemical conversion method3, whereas acid-, base-, and metal-catalyzed hydroly ses are the thermochemical conversion m e t h ~ d . However, ~.~ the hydrolysis reaction does not proceed at a sufficiently high reaction rate. In addition, corrosion of the reactor may occur in the latter and it requires purification process of the waste water.
1338
Recently, some researchers have suggested that ether or ester linkages are hydrolyzed in supercritical water even without any catalyst.6 Sakaki et al.'.' have investigated non-catalytic decomposition of cellulose in subcritical water, and cellulose was found to be decomposed to water solubles, which were decomposed further after their yield reached nearly 80%. On entering the second decomposition process, the water solubles were converted into gaseous products and methanol-soluble products, and solid products like char were formed from the liquid phase. The hydrolysate of cellulose obtained in this process was subjected to a fermentation test, and the formed glucose was successfully converted into ethanol. In our previous study," a supercritical biomass conversion system was developed with a bath-type reaction vessel to hydrolyze woody biomass resources and to convert them into useful chemicals and fuel. This system can cover a range of up to 280MPa in pressure and up to 500°C in temperature, measured with the pressure gauge and thermocouple installed. Water used in this supercritical state behaves very differently from water under normal pressure and temperare." In such a supercritical state, the water can be expected to act as an acid or base,'* but by returning the system to ordinary conditions before pyrolysis occurs, glucose and its derivatives could be obtained in water from cellulose. Therefore, supercritical water treatment can be superior to enzymatic sacchanfication or ordinary acid hydrolysis mentioned above, for the chemical conversion of biomass to useful chemicals. Sasaki et al. have also studied cellulose hydrolysis in subcritical and supercritical water13 and cellulose was found to be rapidly hydrolyzed in supercritical water to sugars. This experiment was conducted with a flow-type reactor in a range of temperature from 290 to 400°C at 25ME'a with a high pressure slurry feeder developed to feed the cellulose-water surries. As a result, hydrolysis products of about 75% in their yield were found in supercritical water, much higher than those in subcritical water. This is because around the critical point of water, the hydrolysis rate of cellulose jumps over one order of magnitude larger in supercritical condition than the subcritical state of water. In addition, Townsend et a1.6 reported that the aromatic model compounds of coals were hydrolyzed in supercritical water and that ether bonds of the compounds can be cleaved during the treatment. It can be, therefore, expected that biomass components can be hydrolyzed in supercritical water. In this study, thus, we selected sugi as softwood and buna as hardwood species. Just for comparison, cellulose and hemicellulose (glucomannan, xylan) and lignin (milled wood lignin) were also studled for their reaction in supercritical water.
MATERIALS AND METHODS MATERL4LS As biomass samples, sugi (Cyptomeria japonica D. Don) and buna (Fagus crenata Blume) woods were selected for supercritical water treatment. In addition, microcrystalline cellulose (avicel), glucomannan extracted from softwood dissolving pulp, larch xylan (Shigma Co., Ltd.), and milled wood lignin (MWL) from buna wood were studied to understand reaction behaviors of wood components during supercritical water treatment.
1339
SUPERCRITICAL WATER BIOMASS CONVERSION SYSTEM The supercritical water biomass conversion system used in this study was revised to be a batch-type reaction vessel from the semibatch-type employed in a previous work." This system can cover a range of up to 280MPa in pressure and up to 500°C in temperature, A schematic diagram of this batch-type reaction system is shown in Fig. 1 in which a 5 ml volume reaction vessel is made of inconel-625. To this reaction vessel, about 5 ml of water was fed with 150 mg of the biomass sample, and then it was attached to the system. To start a treatment of the sample in supercritical water, the reaction vessel was quickly heated by immersing it in the tin bath preheated at 500°C and maintained at supercritical temperature between 380 and 400°C and pressure between 140 and 200MPa for about 5 sec unless otherwise indicated. To quench the reaction, the reaction vessel was moved in the water bath. During this treatment, the temperature of the reaction vessel was monitored by a thermocouple installed into the reaction vessel. At the same time, the pressure was also measured by the pressure gauge attached to reaction vessel.
+I
tfi.3 Water bath
Tin bath
Fig. 1. Biomass conversion system in supercritical water.
P: Pressure monitor, T: Temperature monitor
Fig. 2. Schematic flowchart of the experimental procedure for lignocellulosics as treated in supercritical water.
1340
SEPARATION OF THE REACTION MIXTUREAND THEIRANALYSES According to the scheme in Fig. 2, the obtained reaction mixture of 150 mg sample in 5 ml water after supercritical treatment (>Tc (critical temperature) = 374 "c , >Pc (critical pressure) = 22.1MPa) was filtrated with 0.2 LL m membrane filter (ADVANTEC) to separate water-soluble and water-insoluble portions. The water-insoluble portion was then washed with 10 mL of MeOH to separate MeOH-soluble and MeOH-insoluble portions. The obtained water-soluble and MeOHsoluble portions, the sum of whch nearly corresponds to the supercritical water-soluble portion, were then analyzed by the high performance liquid chromatograph (HPLC) (Shimadzu LC-IOA) which consists of a high pressure pump (Shimadzu Co., Model LC-IOAT). HPLC analysis conditions were as follows: [Column: STR ODs- II, Column temperature: 40"c, Camer solvent: CH,0H/H20 = 20/80 (0-10 min), 20/80 -1OO/O (10-20 min), 100/0 (20-30 min), Flow rate: 0.7 mumin, Detector: a spectrophotometric detector (SPD) ( A=254 nm) or a refractive index detector (RID)], or [Column: ULTRON PS-gOP, Column temperature; S o t , Camer solvent: HzO,Flow rate: 1.0 mumin, Detector: SPD ( A=254 nm) or IUD]. To obtain the yields of monosaccharides and their decomposed products, their standard samples (Nacalai tesque, extra pure reagent) with known concentrations in water were analyzed by HPLC as a standard in a similar manner. To study the molecular distribution of water-soluble and MeOH-soluble portions, gel permeation chromatography (GPC) (Shimadzu LC-IOA) which consisted of a high pressure pump (Shimadzu Co., Model LC-IOAT) was performed with a column of KF803L, operated at 40°C with flow rate of 1.0 mL/min of tetrahydrofuran. The molecular weight of the fractions was determined from calibration curve obtained with polystyrene as a standard. Water-insoluble and MeOH-insoluble portions were also studied by light microscopy (Nikon OPTIPHOT) and the latter portion which corresponds to the supercritical water-insoluble portion was weighed to obtain its yield, and studied by X-ray diffractometry with a hgaku RINT 2000V (Cu-Ka, A=1.542A) at 40 kV and 30 mA to examine its crystallographic nature.
RESULTS AND DISCUSSION SUPERCRITICAL CONDITIONS OF WATER IN R4TCY-TYPE REACTION SYSTEM Figure 3 shows the changes of the temperature and the pressure in the reaction vessel as it was immersed into the tin bath and moved into water bath. From this figure, the water reaches Tc after 15 sec, while for Pc, it takes only 6 sec. Therefore, supercritical condition of water can be achieved after 15 sec. Under such conditions, supercritical water treatment was made for about 5 sec against biomass samples and their cell wall components. SEPARATION OF SUPERCRITICAL WATER-SOLUBLE AND SUPERCRITICAL WATER-INSOLUBLEPORTIONS OF WOOD According to the scheme in Fig. 2, the biomass sample was treated with water under
1341
supercritical conditions, and the reaction mixtures of sugi and buna woods and their cell wall components were separated into water-soluble and water-insoluble portions by centrifugation. However, the reaction mixture always contains oily substances. Therefore, the latter was further treated with methanol (MeOH) to remove oily substances from water-insoluble residues. Because a dielectric constant of water at an ordinary condition being about 80 is reduced down below 10 in supercritical state,I4 oily substances collected as the MeOH-soluble portion must be solvated with supercritical water during the treatment. Therefore, under supercritical condition of water, water-soluble portion plus MeOH-soluble portion would roughly conespond to a supercritical water-soluble portion. The MeOH-insoluble portion, thus, conesponds to the supercritical water-insoluble portion. Table 1 shows the yields of these fractions for sugi and buna woods as treated in supercritical water. Since the supercritical water-insoluble residues were in a small quantity in both woods, it is evident that both sugi and buna woods were effectively converted into supercritical water-soluble fraction in about 5 sec treatment. However, the water-soluble portion was more in buna than sugi wood. In tum,the MeOH-soluble portion was more in sugi wood. This would be due to the different chemical compositions of these two woods as described later. Table I The yields of the water-soluble and MeOH-soluble portions and MeOH-insoluble residues.
Species Sugi Buna
*
Lignin* Supercritical water-soluble(%) (%) Water-soluble MeOH-soluble 39 32.7 57 24.0 76 22 Klason lignin plus acid-soluble lignin
Tin Bath
Supercritical water-insoluble(%) MeOH-insoluble 4 2
Water bath
Immersion time (sec)
Fig. 3. Changes of the temperature and the pressure in the reaction vessel as it was immersed in the tin bath and moved into water bath. The supercritical treatment of water was made for 5 sec. Tc, critical temperature of water=374C; Pc, critical pressure of water=22.1MPa
1342
SUPERCRITICAL WATERSOLUBLE PORTION OF WOOD Water-solubleportion HPLC analysis has been performed with RID detector on the water-soluble portion of the sample treated in supercritical water for about 5 sec for sugi and buna woods and their cell wall components (Fig.4). For cellulose, it is apparent that glucose appears predominantly as a hydrolysate product. In addition, levoglucosan, S-hydroxymethylfurfural (5-HMF) and furfural can be observed. As a minor component, cellobiose and fructose appeared sometimes. On the other hand, for hemicelluloses, glucomannan and xylan, the expected hydrolysate products such as glucose and mannose from glucomannan, and xylose from xylan appear in a very small quantity. Instead, an unknown product appears at 16 min in a retention time in Fig. 4. It seems reasonable that amorphous hemicelluloses are decomposed to a greater extent than crystalline cellulose, under the same supercritical condition of water. Therefore, the supercritical treatment of water for 5 sec is too severe for hemicelluloses. As reported by Mok et al.”, 90% of hemicelluloses is hydrolyzed to be monomeric sugars, under hot compressed liquid water (2 min, 230“C, 34.5MPa). Thus, more appropriate treatment condition must be found for hemicelluloses. Fig.5 shows the HPLC chromatograms with RID detector for buna wood as studied in various tin bath temperatures. It is obvious that oligomers and monomeric sugars such as xylose and glucose from hemicelluloses appear at 300 and 350°C tin bath temperatures and that much higher sugar yields can be obtain in this subcritical state of water. For MWL from buna wood, lignin-derived products sensitive to the ultraviolet light ( A =254nm) could be observed in the water-soluble portion as studied by the HPLC with SPD detector (Fig. 6).
Sugi
0
20
10
30
Retention time (rnin)
Fig. 4. HPLC chromatograms of the water-soiuble portion from woods and their components as treated in supercritical water. (Tin bath 500°C) Detector : RID,Column : ULTRON PS-gOP, Carrier solvent : H,O
1343
Retention time (mid
Fig. 5. HPLC chromatograms of the water-soluble portion from buna woods as treated in various temperatures of the tin bath. Detector : RID,Column : ULTRON PS-BOP,Carrier solvent : H,O
Clucomannan A
L
1 .
0
-
Xylan
L
I
I
10
20
Retentiontime (mid
Fig. 6. HPLC chromatograms of the water-soluble portion from woods and their components as treated in supercritical water. Detector : SPD( A =254nm), Column : STR ODs- II , Carrier solvent : H2O/MeOH=2O/80(O-t 1 Omin)-+01100(20rnin)+O/100(20-+30min)
Based on these lines of evidence, compounds appeared in sugi and buna woods in Figs. 4 through 6 seem to be derived mainly fmm cellulose and hemicelluloses with some from lignin. As shown in Table 1, buna wood contains less lignin than sugi. It is, therefore, reasonable that the water-soluble portion derived mainly from cellulose and hemicellulosesis more in buna than sugi wood.
1344
MeOH-soluble portion The reaction mixture after supercritical treatment of water consists of the water-soluble portion and water-insoluble residues with oily substances. The oily substances seem to be solvated with water in its supercritical state due to their hydrophobic nature, but separated in ordinary condition of water. Such oily substances shown by the arrows in Fig. 7 are recovered with water-insoluble residues, but can be washed with methanol. Therefore, they may be collected as MeOH-soluble fraction. On this MeOH-soluble portion HPLC analysis has been performed with SPD detectoor for sugi and buna woods and their cell wall components.It is quite apparent in Fig. 8 that such oily substances collected as MeOH-soluble fraction is only detected in MWL among cell wall components. Therefore, those from sugi and buna woods can be concluded to be lignin-derived products. Such MeOH-soluble portion derived from lignin is more in sugi wood than buna, as in Table 1. This is consistent with the fact that sugi wood has the higher lignin content than buna as shown in Table 1. Fig. 9 shows GPC chromatograms of water-soluble and MeOH-soluble portions of sugi wood. It is apparent that, compared with the water-soluble portion, the MeOHsoluble fraction has a higher molecular weight. The former has an average molecular weight in a range between 100 and 200, whereas the latter has on average of 530
Me0H-i nsoluble
Water-insoluble
rL-)
ash with
5 0 pm
Fig 7. Light micrographs of water-insoluble residue of sugi wood (left) and MeOH-insoluble residue after washing with MeOH (right). The arrows show the oily substances present with MeOH-insoluble residue.
Glucomannan
Sugi
Buna I
I
0
10 20 Retention time (mid
30
Fig, 8. HPLC chromatogramsof MeOH-soluble portions from woods and their components as treated in supercritical water. Detector : SPD( A =254nrn), Column : STR ODs- Carrier solvent : H20MeOH=20/80(0-t 1 Omin)-+0/100(20min)-t0/100(20+30min)
n,
1345
M
530 170 110
J/ :!
!: {;
{
j ;
i
MeOH-soluble:
Water-soluble
6
I
I
1
I
8
10
12
14
16
Retention time (min)
Fig. 9. GPC chromatograms ofwater-soluble and MeOH-soluble portions of sugi wood as treated in supercritical water. Detector : SPD( ;I=254nm), Column : KF-803L, Carrier solvent : THF
SUPERCRITICAL WATER-INSOLUBLE PORTION OF WOOD The obtained MeOH-insoluble residues, which are supercritical water-insoluble portion of wood, were only 2 to 4 % as shown in Table 1. Therefore, these can be ignored. However, they were studied by X-ray diffractometry as in Fig. 10. For comparison, untreated wood and cellulose were also studied. It is obvious that the MeOH-insoluble residues are amorphous. Perhaps, these residues would be liquefied in supercritical water if the treatment condition is more appropriate.
28(' )
Fig. 10. X-ray diffractograms of supercritical water-insoluble portion (MeOH-insoluble residue) from sugi wood as treated in supercritical.
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CONCLUDING REMARKS Biomass resources will become more important in the future as alternative chemicals and fuel from fossil resources. Such alternatives can be carbon-neutral and renewable. Therefore, the use of biomass is of benefit for our global environments. The present study has clearly demonstrated a potential of woody biomass resources to be converted into useful chemicals in supercritical water. In addition, the products derived from cellulose and hemicelluloses were effectively separated from lignin-derived products. This raises the possibility of converting whole biomass substrates to alternative chemicals, and provides ample opportunities for human beings to produce useful biomass-based products as chemicals and bio-fuels without using fossil resources. ACKNOWLEGEMENT The authors would like to thank TOY0 KOATSU Co. for manufacturing the supercritical water biomass conversion system. This work has been done under the program of the Research for the Future (RFTF) of The Japan Society for the Promotion of Science (JSPS-RFTF 99 P01002). This research was also supported by a Grant-in-Aid for ExploratoIy Research (11876039, 1999), and a Great-in-Aid for Scientific Research (B) (2) (12460144, 2000) from the Ministry of Education, Science, Sports and Culture, Japan. REFERENCES 1 . Boocock D.G.B. & Shennan K.M.(1985) Further aspects of powdered poplar wood liquefaction by aqueous pyrolysis. The Can. J. of Chem. Eng. 63,627-633. 2. Meier D. Larimer D.R. & Faix 0. (1986) Direct liquefaction of different lignocellulosics and their constituents, 1. Fractionation, elemental composition, FueI65,910-905. 3. Boussaid A.&Saddler N.J. (1999) Adsorption and activity profiles of cellulase during the hydrolysis of two douglas fir pulps. Enzyme Microbial Technol. 15, 138143. 4. Mok S.W.& Antal J.M. Jr. (1992) Productive and parasitic pathways in dilute acidcatalyzed hydrolysis of cellulose. Ind. Eng. Res. 31, 94-100. 5. Mnowa T. Zhen F. & Ogi T. (1998) Cellulose decomposition in hotcompressed water with alkali or nickel catalyst. J. of Supercritical Fluids 13, 253-259. 6. Townsend H.S. Abraham A.M. Huppert L.G. Klein T.M. & Paspek C.S. (1988) Solvent effects during reactions in supercritical water. Ind. Eng. Chem. Res. 27, 143-149. 7. Sakaki T. Shibata M. Miki T. & Hirosue H. (1996) Decomposition characteristics of woody biomass in hot compressed liquid water. The Second International Conference on Solvothermal Reactions,' December 18-20, pp 180- 183. 8. Sakaki T. Shibata M. Mki T. 8z Hirosue H. (19%) Decomposition of cellulose in near-critical water and fermentability of the products. Energy and Fuels 10, 684-688. 9. Sakalu T. Shibata M. Uki T. & Hirosue H. (1996) Reaction model of cellulose decomposition in near-critical water and fermentation of products. Bioresource Technol. 58, 197-202. 10. Saka S. & Ueno T. (1999) Chemical conversion of various celluloses to
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glucose and its derivatives in supercritical water. Cellulose 6, 177-191. 11. Fukuzato R. (1998) Chemical recycling process for waste plastics using supercritical water (in Japanese). Sigenkan&youtaisah 34, 1165-117 1. 12.Ramayya S . Brittain A. DeAlmeida C. Mok W.S.&Antal M.J. Jr. (1987) Acidcatalyzed dehydration of alcohols in supercritical water. Fuel 66, 1364-1371. 13. Sasaki M. Kabyemela B. Malaluan R. Hirose S. Takeda N. Adschiri T.&Arai K. (1998) Cellulose hydrolysis in subcritical and supercritical water. J. of Supercritical Fluids 13, 261 -268. 14. Tester J.W. Holgate H.R. Armellini F.J. Webley P.A. Killilea W.R. Hong G.T. & Bamner H.E. (1993) Emerging technologies in hazardous waste management III. In : Supercritical water oxidation techno1ogv:Process development and jhndamental Reseach.@d. by D.W.Tedder&F.GPohland), pp.35 ACS Symposium series 518, Amer. Chem. Soc.,Washington, D.C. 15.Mok W.S. &Antal M.J. Jr. (1992) Uncatalyzed solvolysis of whole biomass hemicellulose by hot compressed liquid water. Znd. Eng. Chem. Res. 31, 1357-1361.
1348
Co-pyrolysis Under Vacuum of Bagasse and Petroleum Residue A. Chaala, M. Garcia and C . Roy Department of Chemical Engineering, Universitd Laval, Sainte-Foy, Qudbec GIK 7P4, Canada
ABSTRACT:Pyrolysis under vacuum of sugarcane bagasse, petroleum residues (PR) and mixtures thereof was performed. Upgraded liquid and solid products were obtained by thermal decomposition of bagasse combined with petroleum residue. It has been found that vacuum pyrolysis of bagasse alone yielded a large amount of bio-oil(34.2% by wt., bagasse anhydrous basis). Although h s bio-oil meets several of the gas turbine fuel specifications, some combustion properties such as the ignition point, the susceptibilityto carbonisation and the atomization can be improved by co-pyrolyzing the bagasse with heavy crude oils or petroleum residues. Blending pyrolytic oils separately produced by vacuum pyrolysis of bagasse and PR resulted in an unstable emulsion. In order to make this emulsion more stable, direct incorporation of PR at different proportions in the bagasse feedstock was performed. The oil blends obtained from the 5 wt.%, 15 wt.% and 50 wt.% PIUbagasse mixtures were stable emulsions. However, the oil from the 30 wt.% PR / bagasse mix was unstable. The physicochemical and rheological characterization of the bio-oils are reported. Ageing tests of bagasse oil were performed at 80°C. The co-pyrolysis process also affected the properties of the pyrolytic charcoal. A surface analysis by ESCA performed on the charcoal samples showed that the bagasse charcoal surface was covered with coke formed during PR decomposition. This increased the ignition point of the bagasse charcoal. 1
INTRODUCTION
Reserves of light crude oils have drastically decreased over the last few decades. If no new crude oil deposits are discovered in the near future, it will be difficult to meet the demand for the energy consumed in fossil fuel power stations. Alternative solutions are to find new extracting and processing technologies for heavy crude oils or to make use of environmentally friendly renewable resources. Many approaches have been investigated which ought to take into account environmental and economical concerns. Thermal decomposition of biomass into transportable, storable and workable fuels is a promising approach to find a solution to the foreseeable shortage of fossil fuels [1-41.
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Several studies have been conducted on co-pyrolysis processes in order to elucidate any synergetic effect on the quality of the products generated. Co-pyrolysis of coal and biomass has been studied by themogravimetry [ 5 ] . No strong interactions between the components of the coal and the biomass feed have been observed. McGhee et al. [6] investigated the co-pyrolysis of polyvinyl chloride (PVC) and wood mixture with straw to simulate municipal solid waste. They found that the carbon obtained has a reduced reactivity. Co-pyrolysis in packed-bed pyrolyser (PBP) of mixtures of coals and heavy petroleum residues [7] and of Australian oil shale and lignite [8] has revealed the prevalence of synergetic effects. These studies also showed that the initial composition of the mixed feedstock influences the product distribution and properties. Vacuum pyrolysis of biomass yields a high amount of oils (biofuels) whch meet several of the heavy fuel performance specifications [9,10]. Some properties of these bio-oils might be improved by upgrading processes such as co-pyrolysis of biomass with heavy crude oils or petroleum residues. Co-pyrolysis can also affect the charcoal properties. The carbonaceous material formed by PR thermal decomposition is usually called coke. The solid product obtained by pyrolysis of biomass is called charcoal. The surface chemistry, the bulk composition and the structure of these two materials are different. 2
EXPERIMENTAL
2. I
MATERIALS
The sugarcane bagasse was provided by United States Sugar Corporation, Clewiston, Florida. The feedstock was air-dried to 8 wt.% moisture content and then sieved in order to remove the particles smaller than 0.450 mm. The petroleum residue supplied by Shell Canada, MontrBal, was a grade 150-200 penetrability petroleum bitumen.
2.2
SAMPLE PREPARATION
Petroleum residue was dissolved in toluene in order to well mix it with the bagasse fibres. The mixture obtained was then carefully shaken and exposed to evaporation during 12 h at room temperature in order to remove the solvent. Residual toluene was eliminated by heating the sample at a temperature of 106°C for 10 h.
2.3
PYROLYSIS RUN
In the present work, bagasse has been pyrolyzed in a pilot plant and a bench scale retort. PR and bagassePR mixtures have been pyrolyzed in the retort. 2.3.1.
Bench scale retort
A sample of 80 g was introduced into the batch retort. All the pyrolysis tests were performed at a temperature of 500"C,a total pressure of about 8 kPa and a heating rate of 1ZoC/min.The holding time of the solid residue after completion of the test was one hour. Vapours formed in the reactor were removed and condensed in three traps connected in series and maintained at -3O"C, -78°C and -78"C, respectively. The non condensable gas was removed by the vacuum pump and stored in a vessel previously depressurized. A detailed description of the apparatus is available elsewhere [ 111
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2.3.2.
Pilot plant reactor
A sample of 20 kg of air-dried bagasse (8 wt.% moisture) was introduced into a cylindrical reactor 3 m long and 0.6 m in diameter equipped with two internal heating plates. The test was conducted in a batch mode at a temperature of 530°C and a total pressure of about 12 kPa. The condensing system included four traps connected in series and maintained at 25"C, O"C, -30°C and -8O"C, respectively. The equipment has been described by Roy et al. [ 121 In order to simulate the condensing system of an industrial vacuum pyrolysis plant which consists of two condensing packed towers continuously operating, the liquids collected in each trap were mixed and then evaporated at 45 "C during half an hour in a rota-vapour (Buchi, RE 111). The heavy fraction which remained in the flask corresponds to the oil from the first condensing tower and is called "bio-oil", while the evaporated fraction whch consists of water and light organic compounds corresponds to the aqueous phase of the second tower and is called "aqueous phase". 2.4
CHARACTERISATION OF THE PYROLYSIS PRODUCTS
2.4.1
Bio-oil
The physico-chemical properties of the oil samples were measured according to ASTM methods: density D 369; h e m a t i c viscosity- modified D 445-88; [9], flash point D-93; gross calorific value D4809; water content D-1744; and ash content D 482. The apparent viscosity was measured using a Brookfield viscometer, LVDV III+. The content of the methanol insoluble materials (MIM) was determined according to the method described by Oasmaa et al. [ 131. The molecular weight distribution was determined by gel permeation chromatography (GPC) using a Waters 510 pump with refractive index detector. The separation was performed using two columns in series: StyragelRHR 4Eand HRI with 5pm particle diameter. THF was used as the eluent at a flow rate of 1 ml/min. The biooil samples were dissolved in THF at a concentration of 2 g/l. The GPC columns were calibrated using polystyrene standards of 400, 600, 760, 800, 1200, 2000, 2430 and 20700 molecular weights. For the stability test, 60 g of bio-oil was poured in tightly closed 60 ml volume bottles. The ageing effect was measured in terms of viscosity, MIM content, water content and molecular weight distribution. 2.4.2.
Charcoal
Proximate analysis of the charcoal sample was performed in a MAC-400 instrument from LECO, while the C, H and N concentrations were determined in a LECO CHN600 equipment following the ASTM D 5291-92 method. A SSC/5200 microbalance from Seko was used for the thermogravimetric tests. Samples of 3-4 mg were heated from room temperature to 600"C, under air flow of 100 d m i n , at a constant heating rate of 5"C/min.
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3
RESULTS AND DISCUSSION
3.1
BAGASSE AND PETROLEUM RESIDUE CHARACTERISATION
The particle size distribution of the bagasse and the proximate analysis of the fractions obtained are shown in Table 1. The fine particles contained in the two last fractions which represented 7 wt.% of the total bagasse, were removed due to their high ash content. All the data reported in this paper refer to the fraction > 0.450 mm. Table 1. Particle sue distribution of the air-dried bagasse and proximate analysis of the fractions. Sieve Denomination Amount Volatile Ash Fixed Carbon (mm) (w%) > 4.75 34 82.3 1.3 16.4 0.85 - 4.75 44 82.1 1.6 16.3 0.45 - 0.85 15 81.8 2.3 15.9 0.25 - 0.45 4 72.5 13.0 14.5 c 0.25 3 60.0 27.7 12.3 Proximate analysis as well as generic and elemental compositions of the sieved bagasse feedstock are presented in Table 2. Some physico-chemical properties of the petroleum residue are also presented in Table 2. Petroleum Residue (PR) Penetration (1/10 mm) Flash point (" C) Viscosity at 135OC (cSt)
163 338 185
Bagasse Volatile matter (wt.%) Ash(*.%) Fixed carbon (wt.%)
82.1 1.6 16.3
High heating value* (MJkg)
41.2
High heating value* (MJkg)
16.7
Carbon (wt.%) Hydrogen (wt.%) Nitrogen (wt.%) Ash (wt.%) Sulphur + oxygen, by diff. (wt.%)
Carbon (wt.%) Hydrogen (wt.%) Nitrogen (wt.%) Trace Sulphur (wt.%) 6.5 Oxygen + ash, by diff. (wt.%)
Asphaltenes (wt.%) Maltenes (wt.%) Toluene insolubles (wt.%)
76.9 23.0
82.8 10.0 0.7
0.1
Hemicellulose (wt.%) Cellulose (wt.%) Lignin (wt.%) **
48.8 5.9 0.5
Trace 43.2 35.2 42.4 20.8
PR is composed of the residues obtained from crude oil vacuum distillation unit, fluid catalytic cracking (FCC) unit, lubricant oil deasphalting unit and other refinery operating units. PR is frequently used as binder for road pavement. PR is rich in high molecular weight aromatic hydrocarbons (asphaltenes). Its carbon content is very high which results in a high heating value. However, it is not a good heating fuel. Copyrolysis with bagasse might be a valuable upgrading process for PR. 1352
3.2.
PYROLYSIS AND CO-PYROLYSIS PRODUCT YIELDS
Vacuum pyrolysis of 20 kg of bagasse camed out in the pilot reactor in a batch mode yielded 30.2 wt.% oils, 25.6 wt.% charcoal, 21.2 wt.% aqueous phase and 22.0 wt.% gas (bagasse anhydrous basis). Meanwhile, vacuum pyrolysis of bagasse camed out at the laboratory scale yelded, on the same basis, 34.2 wt.% oil 19.4 wt.% charcoal, 28.1 wt.% aqueous phase, and 17.6 wt.% gas (Table 3). Differences observed between laboratory and pilot plant tests can be explained by the varying heating rates and the thickness of the feedstock bed inside the two reactors. Heating rate was higher in the retort and the bed thickness was hgher in the pilot reactor. In comparison with other biomass feedstocks which have been previously pyrolysed in our laboratory [14], the charcoal yield is low. Table 3 also indicates the variation in the product velds with the addition of PR in the bagasse feedstock
Table 3. Yields on vacuum pyrolysis of bagasse, PR and mixtures thereof (wt.%, anhvdrous feed basis). Concentration of PR (wt.%) Product 0 50 100 5 15 30 53.3 85.3 31.2 41.3 Pyrolysis oil 34.2 33.9 20.6 10.7 10.3 0.5 28.1 21.1 Aqueous phase 23.0 10.0 29.8 27.8 Charcoal 19.4 28.5 17.5 13.0 3.4 17.6 15.8 19.4 Gas 0.9 0.8 0.6 0.8 0.7 0.7 Loss 100 100 100 100 100 100 Total 3.3.
CHARACTERISATION OF THE BAGASSE BIO-OIL
3.3.1
Physico-chemical properties
The physico-chemical properties of the bagasse-derived bio-oil obtained in the large batch reactor are summarized in Table 4. The bio-oil obtained after evaporation contains 13.8 wt.% water. Like the bio-oils originating from diverse biomasses using various pyrolysis techniques, the oil from vacuum pyrolysis of bagasse is heavier (d20= 1.21 1 g/ml) than water. The low viscosity at 20°C (1 16.5 cSt ) means a flowability of the bio-oil with less energy requirement, while the viscosity at 80°C (5.4 cSt) will ensure a good atomization and consequently a complete combustion. The content in methanol insoluble materials (MIM) which represents not only solid compounds such as charcoal and mineral particles but also waxy-ldce substances, does not meet the usual gas turbine specifications [ 151. The carbon Conradson residue which indicates the susceptibility of the bio-oil to carbonisation is hgh. The h g h heating value ( 22.4 MJkg) makes this bio-oil suitable for combustion in boilers or gas turbines. Its flash point is high enough ( > 9OOC) from a safety point of view. However, it can be corrosive for certain metals (pH = 2.7). Elemental composition showed that the oil contains a very low sulphur content, a relatively low carbon content and a high oxygen content.
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Table 4. Physico-chemicalproperties of the bagasse bio-oil Properties Value Moisture, wt.% 13.8 1211 Density at 20°C, kg/m3 16.4 Kinematic viscosity at 5OoC,cSt MIM, wt.% 0.38 CCR, wt.% 18.6
PH High heating value (anhydrous basis), MJkg Flash Point, "C Carbon (wt.%) Hydrogen (wt.%) Nitrogen (wt.%) Sulphur (wt.%) Ash, (wt.%) Oxygen*, (wt.%) * Determined by difference 3.3.2
2.7 22.4
>90 47.0 5.6 0.6
traces 0.1 46.8
Metal content in bio-oil
In comparison with the bio-oils reported in the literature [2, 31, the investigated bio-oil exhlbits relatively low contents of sodium (Na), phosphorous (P), calcium (Ca) and potassium (K) (Table 5). These concentrations are far from the gas turbine normal specifications [ 151. One should take into account however that the crude bio-oil has not been submitted to any post-treatment. It is important to note the presence of Mg , Fe and Al. In spite of their low concentrations, these metals have potential to inhibit the corrosion caused by the vanadium contained in PR-derived oil.
Table 5. Metals present in the bagasse bio-oil. P Se V Zn Ca Fe K Mg
Metals Na mdkg 21.5 3.3 0.2 3.3.3.
0.02
43.8
33.8
105.9 5.3
2.2
Mn
A1
1.1
10.1
Rheological properties
It has been found that oils obtained from pyrolysis of bagasse and PR, on an individual basis, exhibit a non-Newtonian flow behavior. The viscosity decreasing with increasing shear rates (Figures 1, 2) shows that these oils have a pseudo-plastic behaviour. The non-Newtonian character of the bio-oil from bagasse is due to the partial miscibility of the two liquids (organic compounds and water) which compose it, as well as the presence of solid particles. The behaviour of the oil from PR is due to the threedimensional constituents of the oil (large molecules, colloidal particles and other suspended materials such as asphaltenes and gritty materials). The force required to move these oils is determined by the size, the shape and the cohesiveness of the oil constituents. At any shear rate, the alignment of the constituents is different and more or less force will be required to maintain motion.
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0
0.25
0.5
I
0.75
1.25
1.5
74
76
12
I0 8
6
!z4
4
c
m
2
2
0
0 0
10
5
15
20
30
25
Shear rate (Us) Figure I . - Shear rate vs shear stress for the pyrolysis oils fi-om bagasse and PR @ 30°C
0
0.25
0.5
a. 75
7
1.25
1.5
70
1700 1600
60
1500 50 1400
1 % 40
1300
30
1200 I100
20 1000 I0
900
0
800 0
5
10
15
20
25
30
Shear rate ( 1 / s) Figure 2.- Apparent viscosity vs. shear rate for the pyrolysis oils fiom bagasse and PR @ 30°C
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3.3.4.
Ageing of the bagasse-derived oil
A sample of 60 ml of oil was poured into 60 ml tight glass bottles. The bottles were retightened a few times during the heating-up period. The bottles were rapidly cooled under cold water and weighed before the analyses are performed. The possible difference in the weights before and after the test must not exceed 0.1 wt.% of the original weight. The variation in viscosity, water content and methanol insoluble materials (MIM) was determined. The ageing tests were conducted at 80°C during 6 h, 12 h, 24 h, 48 h and 168 h periods. The variations in the viscosity, water content, MIM content and molecular weight distribution were determined as follows: AY (%) = [(Yz-
Y1)~Yll.
100
y1,yz:
viscosity of original and aged samples, respectively [ cSt] methanol insolubles of original and aged samples, respectively [ wt.%]
A1 (%)= [(I2 - II)/I1].100
11, 12:
AM (%) = [(Mz - Ml)/M1].100
MI, M2: molecular weight of original and aged samples, respectively [ a.mu.1
AW (%) = [(Wz- W1)/ Wl] . 100
WI,W2 :water content of original and aged samples, respectively [wt.%]
The variation in the viscosity, MIM content and water content shown in Table 6 revealed that the increase in viscosity determined at 40°C is very high, particularly for the oil aged during 168 h (134.4%). The viscosity increase is due to diverse reactions, which can occur during heating, generating high molecular weight compounds which may form three-dimensional colloids and miscelles. The water content variation is a parameter indicating the occurrence of these reactions.
Table 6. Variation in viscosity, water content and methanol insoluble materials (MIM) of the bagasse-derived oil after ageing at 80°C Heating time (h) A1 AY (%I AW 40°C 80°C (%I 0 0 0 0 0 12 -26.3 29.6 16.9 9.4 10.1 24 -26.3 47.2 24.5 48 -52.6 73.8 49.0 15.2 168 -63.2 134.4 94.3 15.9
(“w
The changes in viscosity and water content are confirmed by the molecular weight distribution variation (Fig. 3). It is important to note that the changes are particularly intensive in the queue of the oil fraction ( i e . due to the formation of heavy compounds M,, M, and M,+& The changes observed in each fraction of the bio-oil drastically increased from time 0 up to 24 h, then slightly grew up from 24 to 48 h ageing. The increase slowed down during an ageing period longer than 48 h, except for the lighter fraction where the increase had stopped. One can conclude that the ageing process provided more changes in the queue and the middle fractions of the bio-oil than in the head fraction.
1356
100 n
s
3 a v
& B 5 3
.-M
80-
1
AW = [(W, - WJWJ
100
60 -
40
-
B
-aa
20 -
9
O-
Q)
Figure 3 - Molecular weight distribution after ageing at 80°C
3.4 CHARQCTERISATION OF THE BAGASSE CHARCOAL The high content of fmed carbon, and the low volatile matters in the charcoal from bagasse (Table 7) make this material suitable for various uses including production of carbides, reduction of minerals, etc. The heating value (36 MJkg) also allows the charcoal to be used as a solid fuel. Due to its relatively high specific area (529 m2/g), the charcoal might be used directly as an adsorbent material or as a feedstock for activated carbon production. Table 7. ProDerties of the bagasse charcoal Properties Value Moisture, wt.% 3.2 High heating value (anh drous basis), MJkg 36 Specific surface area, m1/g . 529 18.9 Volatile matters, wt.% Ash content, wt.% 6.7 Fixed carbon, wt.% 74.4 Carbon, wt.% Hydrogen, wt.% Nitrogen, wt.% (Oxygen + sulk)*, wt.% * Determined by difference, ash not included
85.6 2.9 1.3 3.5
The metals contained in the bagasse charcoal are presented in Table 8. The elements detected in the bio-oils are also present in the charcoal, but at a much higher concentration. In terms of elemental composition, the charcoal obtained meets several specifications whch are required for carbon-rich materials used in the production of carbides and for the reduction of minerals.
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Metals mgikg
Table 8. Metals present in the bagasse charcoal. Se V Zn Ca Fe K Mg
Na
P
449
1324 6.5 0.4- 27.6 5384 950
Mn Al 5537 2257 73.5 281.5
3.5 CHARQCTERISATION OF THE PRBAGASSE MLYED OILS The properties of the pyrolysis oils obtained during the pyrolysis of the PRhagasse mixtures are summarized in Table 9. CCR decreases with the increase of PR concentration in the bagasse feedstock. This means that PR reduces the susceptibilityof the bio-oil to carbonisation. The content in pentane soluble compounds is indicative of the amount of PR products present in the oil. The oil calorific value increases almost proportionally with the content of pentane soluble compounds. As petroleum residue and bagasse oils have approximately the same ash content, the ash content of their mixtures was almost constant.
Table 9.ProDerties of the mixed oils. PR Concentration (wt.%) 0 5 15 30. 50 CCR (wt.%) 18.6 19.7 19.4 13.8 8.8 Ash (wt.%) 0.05 0.05 0.04 0.05 0.06 22.4 23.0 24.7 36.5 33.2 H H v (MJkg) 4.6 10.0 47.1 62.8 Pentane solubles (wt. %) 2.7 Water content (wt. %) 13.8 9.5 10.3 6.9 11.7
100 9.3 0.05 43.7 89.9 0.2
The viscosity of the mixed oils is higher than that of the bagasse and the PR individually. This is due to the formation of complex three component emulsions (biooil, PR-derived hydrocarbons and water) with dispersed solid particles. As expected, the mixed oils exhibit non-Newtonian flow behaviour (herein not shown). The complex emulsion obtained seems to be more stable than the one obtained by mixing the oils produced separately fiom bagasse and PR. The oils fkom bagasse, PR and the mixed oils were also observed by microscopy. The existence of three liquid emulsions was confirmed by microscopic analysis (Figure 4). The photograph presented in Figure 4 revealed that the mixed oil obtained fiom the 30% by wt PR mixture is not an homogeneous liquid; it forms emulsion. The mixed oils obtained with the 5 wt.%, 15 wt.% and 50 wt. YOPR mixtures seem to be very stable (here not shown). However, the oil from the 30 wt.% PR mixture is not stable. A clear phase separation was observed during the optical test. As an emulsion is generally composed of a dispersed phase and a continuous medium, one can say that at a defined concentration of PR, a phase inversion occurs. Considering the high density of the bagasse oil and its low specific volume, it is postulated that the mixed oils obtained from the mixtures of 5 wt.% PR and 15 wt.% PR are composed of bagasse-derived oil as continuous medium and PRderived oil as dispersed phase. At 30 wt.% PR content, no concrete boundaries between continuous medium and dispersed phase were observed; a competition between phases occurred leading to the formation of an unstable oil blend. At 50 wt.% PR a phase inversion was observed; the PR-derived oil became a continuous medium and the bagasse-derived oil a dispersed phase.
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Figure 4.- Bio-oil from bagasse + 30 % PR
Addition of PR in the bagasse feedstock increased the viscosity of the oil. However, the emulsion obtained exhibited a high storage stability. It has been shown that the incorporation of petroleum fuels (5 wt??) into the bio-oil prior to injecting it into the combustion area of the gas turbine considerably reduced the ignition point of the oil and improved the colour of the flame [16]. Plugging frequency of gas turbine nozzle fed with bio-oils represents a challenge which can be resolved by a special feeding system or by the use of a mixed oil as feedstock. Plugging fiequency of the nozzles is due to the high susceptibility of the oil to carbonisation. This property is evaluated by the carbon Conradson residue (CCR) value, which is high for oils very susceptible to carbonisation and low for oils less susceptible. A correlation between the reduction in plugging fiequency and the CCR tests has to be investigated in the future (Table 9). Combustion of the mixed oils may generate less pollutant compounds such as N O , PAH and SO,. The hot corrosion caused by the vanadium contained in the usual gas turbine fuel might also be inhibited by the presence of some metals contained in the biofuels [ 171. The sulphur content of the mixed oils is expected to be lower than that of conventional petroleum fbels. 3.6. CHARACTERISATION OF THE PWBAGASSE MIXED CHARS
Table 10 indicates that the ash content of the chars obtained during the pyrolysis of the PRhagasse samples decreased as PR concentration increased in the feedstock. As the charcoal fiom bagasse exhibits a similar elemental composition to the PR-derived coke, no great differences in the elemental composition of the mixed chars were observed.
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Table 10. Elemental composition of the mixed chars (anhydrous basis). PR Concentration(wt.%) Element 0 5 50 100 15 30 87.1 85.6 84.5 85.4 84.2 Carbon 88.2 4.3 3.3 2.9 3.4 Hydrogen 2.9 3.5 1.5 2.1 1.3 1.o 0.9 1.2 Nitrogen 5.8 5.6 5.5 0.0 Ash 6.7 6.2 3.0 6.3 3.5 4.8 3.6 5.7 (Oxygen -+ Sulk)* *Determined by difference 3.61.
Surface composition of the mixed chars
The ESCA analysis showed that the elemental surface composition changed with the incorporation of PR in the bagasse feedstock (Table 10). In addition to carbon, oxygen, and nitrogen, potassium and calcium are present on the bagasse charcoal surface. The last two elements disappeared fiom the surface upon addition of 5 wt.% of PR to the bagasse feedstock. This indicated that the bagasse charcoal particles were covered by products of the PR decomposition. Based on the changes in element contents, one can conclude that the addition of 5 wt. % PR to the bagasse feedstock provides major effects on the surface chemistry of the charcoal particles. An extensive study performed by the authors [ 181 confiied that bagasse char surface becomes completely covered when petroleum residue concentration in the feedstock reached 15 wt.%. Table 11. Surface composition of the mixed chars [atom %]
*
N
Si
Ca
K
*
1.1
1.0
0.6
1.2
* * *
0.5
2.1
*
2.9 2.7
Sample
C
0
S
Bagasse Bagasse + 5 % PR
87.6 93.1
9.8 4.2
Bagasse + 15 % PR Bagasse + 30 % PR
93.9 95.8
2.3 1.8
2.6 2.4
Bagasse + 50 % PR PR
94.8
1.6
90.0
4.4
No peak detected
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*
*
*
*
*
*
0.7
*
*
1.3
1.6
*
* * *
3.6.2
Thermogravimetric analysis
Thermogravimetrictests have been performed in order to determine the influence of the co-pyrolysis on the bagasse charcoal reactivity. The DTG curves under oxidizing conditions of the chars obtained by vacuum pyrolysis of bagasse, PR and mixtures thereof are presented in Figure 4. It is clear that the oxidation process of the bagasse charcoal started earlier than the other investigated chars. Oxidation of the mixed chars is delayed due to the formation of coke on the bagasse charcoal surface. T h s coke is made of carbonaceous materials formed during the thermal decomposition of the hydrocarbons contained in PR. The coke is more structured and more compact (low porosity) than bagasse charcoal. The presence of coke and intermediate by-products such as carbenes, carboides and other structured compounds is confirmed by ESCA analysis.
--
I
Bagasse Bagasse + 30 % PR
I
Bagasse + 50 % PR
280
320
360
400
440
480
520
560
600
Temperature (“C)
Figure 5.- DTG ofbagasse charcoal, coke and composite chars
Covering up the bagasse charcoal with coke is beneficial from the point of view of charcoal storage. In fact, the storage of charcoal particles is hazardous as powderous char can easily catch fire when brought into contact with air. The incorporation of PR in the bagasse feedstock increased the self heating temperature of the charcoal particles. The self heating temperature of the charcoal is the temperature at whch the charcoal particles, placed in standardized conditions, warm up and catch fire. From a combustion point of view, the coke formed will increase the ignition point of the bagasse charcoal, which requires additional oil fuing and slower burning in large h a c e s to reach complete combustion. ms type of chars is used for domestic heating, where heat is transferred directly from the fuel bed. The high reactivity of bagasse charcoal on the other hand is desirable in cyclone burners which carry out rapid, intense combustion to maximise carbon utilisation and minimise smoke emission. The mixed chars obtained exhibited lower specific area than the bagasse charcoal [ 181, resulting in a reduction of their adsorption capacity.
1361
4
CONCLUSION
Co-pyrolysis under vacuum of low commercial value products such is sugarcane bagasse and petroleum residue seem to be a promising process. The reactor throughput capacity can be increased by blending petroleum (PR) residue with bagasse feedstock due to the high density of PR. The co-pyrolysis generates two main products: dark mixed oils which have a high heating value, low CCR and relatively high viscosity and mixed chars with high carbon content and low susceptibility to self heating. The mixed oils obtained are complex emulsions consisting of oxygenated organic compounds derived from bagasse, hydrocarbons derived fiom PR and water. Solid particles are also present in the emulsions. The emulsions obtained are more stable than those prepared by mixing bagasse-derived oil with PR-derived oil. The mixed oils exhibit a pseudoplastic behaviour.
5
ACKNOWLEDGEMENTS
The authors are thankful to the United States Sugar Corporation (Clewiston, Florida) and Shell Canada (Montreal, Quebec) for providing the sugarcane bagasse and petroleum residue (bitumen), respectively. Thanks are also due to Dr. H. Pakdel for the GPC analysis and to Dr.J.Yang for the TGA tests. 6 1.
2.
3.
4. 5.
6.
7. 8. 9.
REFERENCES Moses C., (1994) “Fuel-Specification Considerations for Biomass Liquids”. Proceedings Biomass pVrolysis Oil Properties and Combustion Meeting, T.A. Milne, ed., National Renewable Energy Laboratory. Golden, CO, NREL-CP430-7215, pp. 362-282. Diebold J.P. and Bridgwater A.V. (1999) “Overview of Fast Pyrolysis of Biomass for the Production of Liquid Fuels’’ Fast Pyrolysis of Biomass: A Handbook. Aston University, Bio-Energy Research Group, UK (1999). Diebold J.P.; Oasmaa, A.; Bridgwater A.V.; Piskorz J,; Huffman D.; Cuevas A.; Gust S.; Czernik S. and Milne T.A. (1996) “Proposed Specifications for Various Grades of Pyrolysis Oils” Bio-Oil Production and Utilisation. Proceedings of the ZDd EUEanada Workshop on Thermal Biomass Processing, CPL Press, Newbury, UK., pp. 66-81. Meier D. and Faix 0. (1999). Bioresource Technology. 68, pp 71-77. Klose W., Stuke V.(1993). Comparison of the Pyrolysis of Different Types of Biomass and Coals. Fuel Process. Technol. 21, pp. 283-288 McGhee B., Norton F., Snape C. E. and Hall, P.J. (1995). The Copyrolysis of Poly (vinylchloride) With Cellulose Derived Materials as a Model for Municipal Waste. Fuel 74, pp. 28 -3 1 Khan M.R., Heshieh, F. Y. and Heaclky L. (1989). Am. Chem. SOC.Div. Fuel Chem Preprints, 34, p. 1167. Saxby J.D. and Sat0 S. (1990). Liquid Products From Pyrolysis of Synthetic and Natural Blends of Australian Low Rank Oil Shales and Lignites. Fuel 69, pp. 1109-1112 Boucher M.E., Chada A. and Roy C. (2000) “Bio-Oils Obtained by Vacuum Pyrolysis of Softwood Bark as a Liquid Fuel for Gas Turbines. Part I:
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10.
11.
12.
13.
14.
15.
16. 17.
18. 19.
Properties of Bio-Oil and its Blends with Methanol and a Pyrolytic Aqueous Phase. Biomass & Bioenergy, In press. Boucher M.E., Chaala A,, Pakdel H. and Roy C. (2000). “Bio-Oils Obtained by Vacuum Pyrolysis of Softwood Bark as a Liquid Fuel for Gas Turbines. Part 11: Stability and Ageing of Bio-Oil and its Blends with Methanol and a Pyrolytic Aqueous Phase” Biomass & Bioenergy, In press. Pakdel H., Couture G. and Roy C. (1994). Vacuum Pyrolysis of Bark Residues and Primary Sludges. Tappi Journal 77, (7), p.205-211 Roy C., Yang J., Blanchette D. and de Caumia €3. (1997). Development of Novel Vacuum Pyrolysis Reactor with Improved Heat Transfer Potential. In Developments in Thermochemical Biomass Conversion. Blakie Academic and Professional, London, UK, pp. 351-367. Oasmaa A., Lepparntiki E., Koponen P, Levander J., Tapola E. (1997). “Physical Characterisation of Biomass-Based Pyrolysis Liquids : Application of Standard Fuel Oil Analyses”. VTT Energy Publication 306. Roy C. (1999). “The PyrocycIingm Process: New Developments” 4‘h Biomass Conference of the Americas, Oakland, CA, U.S.A. Orenda Aerospace Corporation. (1997). Engine Summary Report : Bio-fuel Testing and Optimisation for Gas Turbine Applications. Internal Report. SPE “Mashproekt” (2000). The GT 2500 Gas Turbine Engine Fuel Nozzle and Combustion Liner testing on Pyrvac Biofuel. Technical Report, Kiev, Ukraine. Tiwari, S.N. and Praliash, S. (1998) “Magnesium Oxide as Inhibitor of Hot Oil Ash Corrosion”. Xke Institute of Materials, pp. 467-472. Darrnstadt, H. ; Garcia-Perez, M. ; Chaala, A. Cao, N.Z. and Roy, C. CoPyrolysis Under Vacuum of Sugar Cane Bagasse and Petroleum Residue. Properties of the Char and Activated Char Products. Carbon. In press. Anonymous. Transport Canada - Rkgion du Qukbec. Transport des marchandises dangereuses. In Recommendation on the Transport of Dangerous Goods, Manual of Tests and Criteria. United Nations (1993).
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Preliminary Results on Wood Waste Pyrolysis C.Dudouit' and Y. Schenke12 ' ENSTIB - School of Wood Science and Engineering, 27, Rue du Merle Blanc F - 88051 Epinal Cedex 9 - France CRA,Agricultural Engineering Department, Chausske de Namur, 146 - B- 5030 Gembloux Belgium
ABSTRACT A safe and efficient treatment of organic solid waste is becoming more and more important, as many European countries will ban these residues from landfill. Among them, wood waste, such as demolition wood, coated wood, etc., have received little attention as far as their treatment is concerned. Pyrolysis, and more precisely thermolysis, represents a safe and efficient treatment of these contaminated woods. Preliminary experiences have been conducted on the pyrolysis of waste wood, at three pyrolysis temperatures (350,550 and 750°C), with a heating rate of 20" C/min and a residence time of 60 min. A particular attention has been given to the behaviour of heavy metals: As, Cd, Cr, Cu, Hg, Ni, Pb, Zn. These heavy metals are mainly concentrated in the residual char with the exception of As, Hg and Cd, which are volatile at low temperatures. A large fraction of these last three metals are recovered in the tars. The same phenomenon is observed with Zn at the temperature of 750" C. INTRODUCTION A promising way for the disposal of wood waste (demolition wood or residues of wood processing industries) is pyrolysis. Belgian industries produce some 1 million tons of waste wood [l], France more than 7millions and Germany 3 to 3.5 millions [2]. In 2004, the landfill of such refuse will not be authorised anymore in Belgium (2002 in France) except for the ultimate waste. Wehlte S. et al. [2] describe the hazardous potential and characteristics of several wood waste assortments in the former Federal Republic of Germany. They show that these refuse contain various wood preservatives, insecticides or coatings (lacquer, paint, vanish) which may contain heavy metals like copper, chromium, lead, andor substances like boron, arsenic, fluorine, pentachlorophenol, lindane, formol ... In fact, more than 80 % of demolition wood waste have been treated with compounds of which the toxicity degree could be variable [3].Some products as pentachlorophenol, lindane have the reputation of being teratogenic, mutagenic, carcinogenic or also neurotoxic. So, it is necessary to identify a suitable treatment respecting environment and health, and creating energy or by-products.
1364
MATERIALS AND METHODS
These preliminary tests of wood waste pyrolysis have been conducted on so called “class B y wood waste; the sample was taken in a waste wood pile of a waste management company located in Belgium. The sample has been characterised and then pyrolysed slowly at low and medium range temperatures in a thermo-balance retort kiln. WOOD WASTE CHARACTERLZATION
The particle size varies a lot. The finest particles (less than 1 cm long or thick) have been eliminated by sieving in order to not alter the pyrolysis yield results ; they represent about 2 % of the total mass of the feedstock. The bulk density is about 149 kg/m3dry basis. The particles have been classified according to their thickness and nature. The results are presented in figure 1 and 2.
Fig. 1 Classification of the particles according to their thickness (t in mm). The high standard deviation for a thickness between 20 and 25 mm can be explained as follows: the samples collected for this measure weighed 250 g; the presence of thick particles dramatically change the results and consequently thz. average. To obtain results with less standard deviation, the samples weight should be higher (that is to say at least 4 kg).
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*Varied : Hardboard / Laminated hardboard I Forestry wood (branch, leaf etc) / Wood / board including nails or staples I OSB (Oriented Stand Board) I Varied (linoleum, stone, glass wool, aluminium, cable, hinge, glass, plastics etc) I Fireproof or anti-damp particle board Fig.2 Feedstock characterisationfiom representative samples introduced in the kiln for each pyrolysis experiment
The feedstock composition shows a large variety of products resulting fiom a great variety of treatments more or less toxic. The heavy metals content of our samples was determined (table 1). Table I Heavy metals content of wood waste (mg/kg of dry matter)
Testno 1 Test n o 2 TestnO3
Cd 0.75 0.24 0.29
Cu 17 7.6 7.4
Ni 74 24 21
Pb 152 81 71
Zn 237 117 106
Hg 0.27 0.13 0.12
Cr 100 31 28
As 0.59 0.08 0.24
The other physico-chemical characteristics of the wood waste samples are given in Table 2. Table 2 Physico-chemical characteristics of wood waste samples
Proximate composition (% dry mass) Volatile content 82.67 Mineral content 2.40 “Fixed carbon” 14.93 content Calorific value (MJkg dry mass) Gross 19.63 Net 18.40
Ultimate composition (% dry mass) Carbon 45.60 Hydrogen 6.01 Nitrogen 1.77 Oxygen 44.22 Moisture content (% mass) Dry basis 24.35 Wet basis 19.58 1366
These pyrolysis experiments have been performed on actual feedstock representing the wood waste that will have to be treated in the f h r e (different particle sues, wood species, products, treatments etc).
SLOW PYROLYSIS EXPERIMENTATION PLANT Thls plant is hlly described in [4] and mainly consists o f : (a) The reactor: cube-shaped, it has a capacity of 27 litres. The inside walls are made of refractory bricks. The feedstock is put in a metallic basket and the reactor is closed tightly with a cordon of clay, whch is replaced at each experiment. Sealed up, the reactor is totally airtight. (b) The heating control system: the electrical resistances are driven by a numeric regulator which sets the temperature profile and gradient as well as the final temperature of pyrolysis. The amplitude variation of the real temperature is more or less 20°C in comparison with the temperature settings. (c) The electronic balance: the reactor is set on an electronic balance to follow continuously the loss of weight (measurement precision: 0.5 % at 10 kg). (d) The gas conditioning and analysis system: at the exhaust of the reactor, the pyrolysis gases are first condensed, filtered and continuously analysed (non condensable fraction). The liquid fraction (condensable gases) is collected and weighed. Thermocouples give the gases temperature at different levels of the plant : in the reactor and along the conditioning line. The non condensable gases are analysed by the means of a NDIR spectrometer (determination of carbon monoxide (CO), carbon dioxide (COJ and methane (CH,) content), a thermal conductivity analyser (hydrogen (H2) content) and a content). magnetomechanics analyser (oxygen (02) The pyrolysis oil is collected at the bottom of the condensation columns, weighed to obtain the wet mass before determining its moisture content (Karl Fisher method).
EXPERIMENTAL DESIGN Four lulograms of anhydrous wood waste have been pyrolysed in the thermobalance retort kiln without air supply. The first set of pyrolysis conditions were: 0
0
Heating rate: 20 "C/min. Residence time at final temperature: 60 min. Final temperature: 350,550 and 750 "C.
These conditions were followed for the first run,but some leaks at the oven were observed. As the catch of the facility could not be modified, we decided to decrease the heating rate from 20"C/min to 2"CImin hoping for a lower pressure in the hln. Unfortunately, the leaks persisted.
1367
RESULTS AND DISCUSSION PYROLYSIS TESTS
It was demonstrated that the pyrolysis begins earlier and advances faster when the particles (study carried out on cubic particles of 2, 4, 8 cm side) are disposed close by the kiln surfaces. The pyrolysis process is significantly delayed and slowed if the particles are disposed in bulk, in staggered rows and in the centre of the kiln, respectively [4]. With wood waste, the feedstock characteristics (disposition in bulk, thick particles, low bulk density), lead to a very bulky bed. These conditions favour the rate of the pyrolysis process through a enhanced heat transfer and evacuation of the pyrolysis products. The loss of weight has been followed during the experiments. There are not much differences between the loss of weight for pyrolysis temperatures of 550 "C or 750 "C. In the fiture, it would be interesting to perform such experiments with pyrolysis temperatures from 500 "C to 800 "C, in order to find the best appropriateness between the pyrolysis products quality and their heavy metals contents. PYROLYSIS PRODUCTS Gas analysis
At the beginning, the presence of nitrogen and oxygen (inert atmosphere) is constant until the gas flow become so high that air is ejected outside the kiln.From this time, the production of carbon dioxide, carbon monoxide, methane and finally ethane are successively observed. It was impossible to compare carbon dioxide, nitrogen and hydrogen contents in the different runs because of the carbon dioxide sensor breakdown . Only the gas flows and carbon monoxide, methane, oxygen contents could be compared. The different pyrolysis temperatures do not influence the carbon monoxide, methane and oxygen contents: the maximum contents and the general evolution of the composition are the same. The gas flow increases after about lh50, which corresponds to a temperature of about 320 OC, until 3h10 (equivalent to a temperature of 480 "C for the final pyrolysis temperature of 550 "C and 750 "C). So, we have a remarkable gas production during lh20. It is difficult to determine the best gas flow ; more experiments could be carried out for a finer analyse. We could also analyse the calorific value of these gas during the experiments (we have done it in order to not interfere with the pyrolysis process and in the general results). The installation for the gases conditioning and analysis could be modified. For example, we could imagine:
0 0
To increase the gas exit diameter to favour the gas evacuation and so to diminish the pressure in the kiln. To use strainers after the flow meter to capture arsenic [5] or other particles to study To install a gas-meter to know the gas volume produced The injection of a hot gas (like N2)in order to have a best control of the gas flow and simulate actual industrial retorts.
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Charcoal analysis
Mass yield of waste wood carbonisation are compared to the results of [4]:anhydrous cubes (2 cm) of Fagus silvatica were tested in a laboratory kiln in a very controlled way, the heating rate being 2 "C/min and the residence time at final temperature was 15 min. The results obtained on waste wood (Fig. 3) are lower than the mass yields obtained in a laboratory kiln. In other words, more initial matter (solid wood) has been volatilised in the case of waste wood pyrolysis. 50
35 30 25
1
I
++
I
-
0
~
A t
A
-
0
20
0 Observed mass yield
+Theoric mass yield
AsA
APublication data
0
15 -
"
I
0
200
400
800
600
Pyrolysis temperature ("C)
1000
-
Fig. 3 Comparison of mass efficiencies of waste wood (mass efficiency) and beech wood (publication) charcoal.
The results of proximate composition and calorific value have been also compared to [4](Figures 4 and 5). Figures 4 and 5 show that the proximate composition and the calorific value of the waste wood charcoal are equal to the values obtained on beech blocks in a laboratory test. The quality of waste wood charcoal does not seem to be influenced by their heavy metals content. Pyrolytic liquids analysis
Because of the leaks observed for some tests, the results of the analysis of the pyrolysis oils vary a lot and hence do not gibe any reliable in formation. Heavy metals contents
The heavy metals contents of the wood waste charcoal and pyrolytic liquids have been compared to the heavy metals contents of the wood waste (Table 3).
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100
90 80
+Volatile matter content waste wood W Fixed carbon content waste wood 0 Volatile matter content beech wood 0 Fixed carbon content beech wood
70 60
50 40
+
30 20 10
0
Pyrolysis temperature ("C)
Fig. 4 Proximate composition (volatile and fixed carbon contents) of waste wood and beech charcoals.
34 33 P 32 31 2 30 29 28 27 26 25
8
+GCV waste wood
+0
-5 2
charcoal
0 NCV waste wood charcoal
A
AGCV beech charcoal
300
500
700
Pyrolysis temperature ("C) Fig. 5 Comparison of the calorific value (gross - GCV and net - NCV) of waste wood and beech wood charcoals.
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Table 3 Average values of heavy metals content of wood waste, wood waste charcoal and pyrolytic liquids (mgkg).
Cd Wood waste 350 O C
0.43
Charcoal 0.68 Pyrolytic liquids 0.35
Cu Ni Pb Zn 10.67 39.67 101.33 153,33
Hg 0.17
Cr
As
53.00
0.30
6.17 2.47
3.27 0.98
59.67 1.23
153,67 32,60
0.02 0.03
53.50 0.67
0.63 0.46
12.87 1.13
1.33 0.28
83.33 0.73
188,67 12,27
0.07 0.04
8.67 e0.25
1.26 0.09
23.87 0.97
62.33 0.15
51.00
189,33 77,67
0.03 0.02
145.33 c0.25
8.00 0.69
550 "C
Charcoal 1.94 Pyrolytic liquids 0.20 750 "C
Charcoal <0.05 Pyrolytic liquids 0.28
0.71
Mercury, arsenic and c a h u m are metals, which are volatile at low temperatures (356.88, 614 and 767 "C respectively). Such metals must be captured at the exit of the gas after the condensation columns. Arsenic appears in treatments like CCA. It was shown that CCA (Copper, Chromium, Arsenic) treated wood could be pyrolysed with a negligible release of copper and chromium and a minimal release of arsenic (76,84 % of arsenic was found in the pyrolysis residues) at a reactor temperature of 350 "C during 20 minutes, using a nitrogen at a flow rate of 5 Nm3/h. [ 5 ] Figure 6 shows that the increase in zinc and copper content in the charcoal is due to their concentration in thls product. The pyrolytic liquids content in these two metals is very variable. But the majority of zinc and copper is found in the charcoal as well as for nickel, lead and chromium. The same conclusion was observed for CCB (Copper, Chromium, Arsenic) treated wood [2]. As the ash content of different biomass could influence the gasification reactivity of biomass chars, a study has been conducted to determine the influence of heavy metals on the gasification process [6]-[7].They show that the alkali metals increase the reactivity of wood char and that lead, copper and zinc, especially as chlorides, inhibited the gasification of the char (about 2.5 times slower that the char from untreated wood). The study of the chlorine content of our wood waste could be attractive: the presence of particles board, oriented stand board, hardboard points to the presence of chlorine usually used in glues. The combustion of such products in a boiler generates nitrogen oxide, sulphur dioxide, carbon monoxide, hydrochloric acid and formaldehyde
PI.
Finally, it would be necessary to study the heavy metals contents of the waste wood sample of each experiment, in order to increase the accuracy of the mass balance of the heavy metals.
1371
200 180 160
pE
140 120
z
100
s
80
B
60
40
3 00
400 500 600 700 Pyrolysis temperature ("C)
800
Fig. 6 Zn and Cu content in waste wood charcoal. CONCLUSION
These orientation tests enabled us to identify the problems encountered in waste wood slow pyrolysis and the modifications to make. In order to gain time and improve the analysis, it is recommended to inject hot gas in the kiln and to inject the cold gas at the completion of the test to decrease the temperature inside the kiln. The pyrolysis experiment required in fact two days to be carried out : one day for the experiment and another to empty and refill the kiln. With regard to the heavy metals contents of the pyrolysis products, we can say that the heavy metals mainly concentrate in the charcoal except mercury, arsenic and cadmium which are volatile at low temperatures. Significant quantities of these three heavy metals are found in the pyrolytic oil. In further experiments, tests should be performed in accordance with industrial conditions (high temperatures). Other studies could be conducted on the possible uses of the pyrolysis products, which were contaminated with heavy metals, on the thermochemical aspects of the wood waste pyrolysis (chemical reactions of the wood preservatives during the pyrolysis process). REFERENCES
[I] Temmerman M., Van Belle J.-F. (1998) Contribution desfilizres bois knergie au ddveloppement durable en Belgique. Activity report of the Agricultural Research Centre of Gembloux (Belgium), 195 p.
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[2] Wehlte S., Meier D., Moltran J. And Faix 0. (1997) The impact of wood preservatives on the flash pyrolysis of biomass. In : Developments in thermochemical biomass conversion (Ed. by A.V. Bridgwater & D.G.B. Boocock), pp. 206-2 19, Blackie Academic & Professional, London [3] Guide des dechets de chantiers (1998). Collection: Conna?tre pour agir. Publie par I’ADEME [4] Schenkel Y. (1999). Modelization of mass and energy jlows from wood carbonization in retort kilns (PhD thesis in French). Gembloux (Belgium), FacultC Universitaire des Sciences Agronomiques, 328 p. [ 5 ] Helsen L. And Van Den Bulck E. (1997) Release of metals during the pyrolysis of preservative impregnated wood. In : Developments in thennochemical biomass conversion (Ed. by A.V. Bridgwater & D.G.B. Boocock), pp. 220-228, Blackie Academic & Professional, London [6] Von Scala C., Struis R.,Stucki S. (1996) The influence of heavy metals on the gasification of wood. In: Biomass for energy and the environment (Ed by P. Chartier, G.L. Ferrero, U.M. Henius, S. Hultberg, J. Sachau, M. Wiinblad), pp 1388-1391, Elsevier Science Limited, Oxford. [7] Von Scala C., Struis R., Stucki S. (1997) The influence of chlorine on the gasification of wood. In : Making a business from Biomass in Energy, Environment, Chemicals, Fibres and Materials (Ed by R.P. Overend & E. Chornet), pp 415-421, Elsevier Science Limited, Oxford. [8] Falandrin B. (1999). Wood Board Combustion: the Environment Is Respected. CTBA report, Bordeaux, France.
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Fast Pyrolysis of Industrial Biomass Waste Ch. Gerdes’, D. Meier’, W. Kaminsky2 I Federal Research Centrefor Forestry and Forest Products, Institutefor Wood Chemistry and Chemical Technology of Wood, 0 - 2 1027 Hamburg, Germany ’University of Hamburg, Institute for Technical and Macromolecular Chemistry, Bundesstraljle 45,0-20146 Hamburg, Germany
ABSTRACT In cooperation with several companies different biomass waste such as wood waste, fibre sludge, cocoa shell and panel boards with a high content of phenolformaldehyde resin have been decomposed by flash pyrolysis into small molecules. The thermal conversion of biomass was performed in a fluidised bed reactor (PDUscale, throughput 5 kg biomassh) using the modified and optimised “Hamburg Fluidized Bed Process” (“Hamburger Wirbelschichtverfahren”). The main product was a liquid (bio-oil) with a maximum yield of 50 to 65 wt.%. In addition, pyrolysis gases and solids (charcoal; inorganics) were formed. In many cases pyrolysis of different real biomass waste led to a good yield of bio oil. However, also difficulties with the fluid bed reactor were observed such as large variation in particle size. The oils could be used for production of power and heat, as chemical feedstock or in some cases recycled into the production processes (resin for panel boards). INTRODUCTION Industry is not yet able to make complete and rational use of their biomass residue and waste. Intelligent waste disposal and recycling techniques are necessary to meet stronger legislative requirements in future. Due to the newer legislative requirements in Germany it will not be possible to deposit wood waste by landfill after end of year 2005 [l], [2]. Disposal companies are forced to find the best process to handle waste-wood. Wood waste can be contaminated with inorganic and organic wood preservatives such as CCB, Cu-HDO, Al-HDO, lindane, DDT etc.. Numbers of studies and papers do predict “golden years” for thermo-conversion processes, especially for pyrolysis processes for impregnated-wood [3] [4]. With help of the well known “Hamburg Fluidised Bed Process” [5] investigations were performed to improve this technology for pyrolysis of artificial “self made” waste wood with a homogeneous well defined contamination [6] [7], [8], [ 9 ] , [lo], [ l 11. Based on these earlier studies real waste wood fractions from a German wood disposal company and residual hardwood from a GermadNorth American charcoal manufacture was pyrolysed. These investigations shall prove the flash pyrolysis technique to be useful in wood waste exploitation. Looking at the wood processing industries, like furniture, panel boards or charcoal producers, it can be noticed that often a large amount of grinding grit remains from
1374
production process. Usually, this grit is burned or deposed by landfill. Going back to the argumentation of the newer legislative requirements concerning waste-wood a new way of disposal should be found for the same reasons. Therefore the pyrolytic behaviour of grinding grit from a German company was studied. Another possibility for the use of flash pyrolysis is to bring smaller scale plants into the market. Thinkable are plants with a capacity of 1-10 t/d max.. There are a lot of biomass processing companies with a more or less unusual biomass waste fraction. This waste has no value for these companies at the moment, it only entails disposal costs. It will be demonstrated that these biomass can lead to acceptable bio-oil yields. In this study cocoa shell from a German cocoa roasting company with an output of 45 t coca shell per day was investigated. Today, paper recycling is a common practice in paper business. A typical length of fibres for paper production is 0.7-0.8 mm for hardwood and 2.0-2.5 mm for softwood fibres. With each cycle, the paper fibres become shorter. After several cycles the fibres are too short to make paper of acceptable quality (below 0.1-0.2 mm). This is also valid for the production of new paper. In the year 1997 15,953,362 t of new paper, cardboard and carton were produced in Germany. As a result, German paper mills had an output of 239,300 t fibre sludge in 1997 [12]. Additionally 9,457,411 t of used paper were recycled in the same year [13]. Caused by the lower quality as a raw material for paper production, 20-30 wt.% of used paper is removed from the process due to the short fibres. Ultimately, there were at least 1.9 Mio. t. of sludge from paper recycling in 1997. Fibre sludges contain a large amount of organic compounds, inorganic fillings and make-up chemicals. Normally, these sludges are burned or deposed by landfill like grinding grit and other waste. The.fibre sludge used for pyrolysis was supplied by a German paper mill. This study wants to demonstrate an alternative way of handling them. High pressure laminates (HPL) are used as panel boards, tabletop, worktops in kitchen or laboratories, facing for buildings etc.. These panels are produced with an amount of 30 wt.% phenol-formaldehyde resole resin. The costs for the resin ground material are quite high, e.g. the phenol costs are approximately 825 Euro/t [ 141. This price and stronger legislative requirements could lead to new ways for the recycling of production waste or recycling of used up laminates from the consumer. One way could be fast pyrolysis of HPL. The aim would be monomer recovery -phenols from resin. Currently, existing pilot plants in Canada, Netherlands, UK utilize mainly well defined non contaminated biomass fractions such as wood particles, saw dust, and bark. The performed investigations in this work should broaden the knowledge of the pyrolytic behaviour of various industrial biomass waste. This will facilitate the introduction of flash pyrolysis processes into existing industrial processes. Therefore, a new way of biomass exploitation will be demonstrated. In cooperation with several companies different biomass waste such as cocoa shell, wood waste, fibre sludge and panel boards with a high content of phenol-formaldehyde resin were decomposed by flash pyrolysis into smaller molecules to use them for the production of energy andor chemicals.
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EXPERIMENTAL Thermal conversion of biomass was performed in a Pilot Demonstration Unit (PDU) scale fluidised bed reactor with a throughput of 5 kg biomassk and in a laboratory fluidised bed reactor 1 ghatch. Therefore the “Hamburg Fluidised Bed Process” (“Hamburger Wirbelschichtverfahren”) had to be modified and empirically optimised for biomass feedstock.
PDU-SCALE FLUIDIZED BED PROCESS
L
0
13
T ._.--
11
t
gas circulation
Fig. I Flash pyrolysis pilot plant for biomass (PDU-scale); (1 hopper, 2 vibration conveyor, 3 screw feeder, 4 fluidised bed reactor, 5 cyclone, 6 heat exchanger, 7 intensive cooler, 8 electrostatic precipitator, 9 flare, 10 compressor, 11 gas preheater 1, 12 gas preheater 2, 13 overflow container). The pilot plant (Fig. I) works continuously with a capacity of 5 kglh (biomass). The feedstock is transported by a vibration conveyor (2) and a screw feeder (3) into the fluidised bed reactor (4). The volatile pyrolysis products and smaller char particles leave the reactor through the reactor head together with the volatiles. Larger charcoal particles leave the reactor directly via an overflow pipe into the collection vessel (13). Micro-char particles are separated from the gas flow with a dual cyclone system (5). The higher boiling components are separated in the water-cooled heat exchangers (6). The remaining condensable components are collected in the intensive cooler (7). Aerosols are separated from the non-condensable gases with electrostatic precipitators (the second is installed for safety reasons) (8). The gas volume necessary for fluidisation is transported with a compressor through two preheaters (1 1, 12) back into the reactor. The remaining gas volume is burned in a flare (9). The plant runs continuously with recycled pyrolysis gases. The footprint is approximately 4 x 2 m. The height is ca. 2 m.
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UPDATED PDU-SCALE FLUIDIZED BED PROCESS
n
I
gas circulation
I
Fig. 2 Updated flash pyrolysis pilot plant for biomass (PDU-Scale); ( 1 hopper, 2 vibration conveyor , 3 screw feeder, 4 fluidised bed reactor, 5 cyclone system, 6 monopump, 7 quench liquid reservoir, 8 heat exchanger, 9 spray tower, 10 electrostatic precipitators, 11 heat exchanger, 12 flare, 13 compressor, 14 gas preheater 1, 15 gas preheater 2, 16 overflow pipe, 17 char collection vessel).
An updated design comprises recirculation of a pre-cooled hydrocarbon liquid with an mono-pump. The quench medium is immiscible with pyrolysis liquids due to the low polarity compared to pyrolysis liquids and can therefore be separated in the lower part of the quenching column and passed into the reservoir (7). Resisting aerosols are collected by electrostatic precipitators. Additionally, a cooler was installed to remove the fluidising gas from traces of lower volatile compounds (water etc.) (1 1).
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LABORATORY FLUIDIZED BED PROCESS For screening tests of new, unusual feedstock such as cocoa shell or fibre sludge a small fluid bed glass reactor for batch runs of 1 g biomass was designed and built (Fig. 3).
Fig. 3 Flow scheme of the glass reactor for biomass; (1 nitrogen inlet, 2 silo with screw feeder, 3 glass reactor, 4 fluidised bed, 5 heating system, 6 wire screen, 7 heat exchanger, 8 glass tube with cotton filter, 9 gas outlet).
The glass reactor works continuously with a capacity of 1 g (biomass). The feedstock is transported from the reservoir by a screw feeder (2) into the fluidised bed reactor (3). The bed (quartz sand) (4) is fluidised with nitrogen (1) and heated with an electrical heating system. The volatile pyrolysis products are leaving the reactor through a wire screen ( 6 ) to prevent the pyrolysis liquids from charcoal contamination. The volatile components are condensed in a cooling trap filled with a mixture of acetone and dry ice (7). The non condensable gases and the aerosols pass through the trap into a cotton wool filter (8). The non condensable gases are leaving the reactor through a gas outlet (9).
Experimental conditions Pyrolysis temperature has the greatest influence on the quality and quantity of pyrolysis liquids. Earlier studies indicate the optimum temperature for a high oil yield between 475 and 500 "C [ 6 ] [15]. Table I shows the experimental conditions for three different feedstock investigated by lab. scale pyrolysis (glass-pyrolysis-reactor). The experimental conditions for the different feedstock pyrolysed with the PDU-scale equipment are presented in Table 2.
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Table I Experimental conditions of laboratory scale pyrolysis. feedstock
cocoa shell fibre sludge grinding grit
feed feed, dry particle size feedstock moisture ash content pyrolysis temperature nitrogen flow exp. time
[mgl [msl [mml [wt.%] [wt.%] ["C] [Umin] [s]
1155 1040 <1 9.9 0.5 475 2.7 180
1308 1270 <1 2.9 48.0 475 2.7 120
1185 1173 <1 1 .o
7.8 475 2.7 129
Table 2 Experimental conditions of the PDU scale pyrolysis.
feedstock pyrolysis temp. particle size ash content feedstock moisture total feed total feed atro* org. feed
["CI [mml [wt.%] [wt.%] [g] [g] [wt."/]
beech fibre waste sludge 484 486 1-3 1-3 4.19 48.00 0.94 7.47 6000 10000 5552 9906 88.34 51.06
HPL 464 1-3 2.45 5.40 8124 7685 92.15
mixed waste wood 458 1-3 5.25 12.42 16602 14541 85.13
quartz quartz quartz quartz [-I sand sand sand sand bed material size of bed material [ m ] 0,3-0,5 0,3-0,5 0,5-0,6 0,5-0,6 5500 amount of bed material [g] 6000 6000 5200 experimental time [h:min] throughput [g/h] reactor residence time [SI total gas flow/cold [m'/h] total gas flowihot [m3/h]
1:57 3077 2.18 7.21 17.84
2:11 4580 3.16 4.97 12.28
1:55 4239 2.62 5.99 14.81
5:OO
3320 2.23 7.06 17.46
less feedsiock moisiure
The feedstock was prepared by milling and sieving to the suitable size Important characteristics of each feedstock such as moisture, ash content and organics are displayed in Fig 4 for the lab scale pyrolysis. The characteristics of the PDU scale pyrolysis can be found in Fig. 5
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8 5
90%
-
80%
-
70%
-
60%
1
50%
L
40% 30%
I
i1 I
20%
10% 0%
1
mixed
beech waste
fibre sludge
HPL
minorganics
7.47 4.19
.organics
88.34
0.94 48.00 51.06
5.40 2.45 92.15
0 water
waste
wood
reference (beech)
12.42
5.44 0.58 99.42
5.25 82.33
Fig. 4 Characteristics of laboratory scale pyrolysis feedstock.
loou/o
1
80% -
60%
-
40%
-
3
g
t;
Y
20% -
0% ...
coma shell
fibre sludge
grinding grit
reference (beech)
0 water
9.89 0.50 89.61
2.88 48.00 49.I2
I .oo
0 inorganics .organics
7.80 91.20
8.67 0.58 90.15
Fig. 5 Characteristics of PDU scale pyrolysis feedstock.
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Analysis
All pyrolysis products were analysed by chem./phys. methods (GC-MS, GC-FID, GCWLD a.0.). The feedstock moisture was determined by drying according to DIN 51718. The ash content was measured by annealing according to DIN 5 1719. Pyrolysis gas was investigated by GC-FID and GC-WLD using definite gas mixtures for calibration. The pyrolysis product with the highest value, the pyrolysis liquid or bio-oil was analysed by GC-MS. Each identified compound in the oil was then quantified by GCFID. The total amount of the high molecular weight fraction, so-called: pyrolytic lignin was determined by a special precipitation method [16][ 171. The water content of biooil was measured by Karl Fischer titration method.
RESULTS AND DISCUSSION The first step on the way to the decision if fast pyrolysis is a suitable technique for biomass waste exploitation is to take a serious look at the mass balance. In almost every case bio-oil is the pyrolysis product with the highest commercial value no matter weather it is used as bio-fuel for power and heat production or as chemical feedstock. The by-product charcoal can have the potential to be refined to activated carbon. The pyrolysis gas is the product with the lowest value. It can only be used as a low heating value gas for burners. Usually the inorganic fraction has no specific value. Looking at fibre sludges from paper mills it consists of cheap filling material (from the paper production process) like Ca(C03), kaoline, make-up chemicals etc.. In other cases e.g. grinding grit, HPL or waste wood the inorganic fraction consists of an inhomogeneous undefined mixture of different substances or, in worst cases, of metall and heavy metall mixtures. Existing pilot plants in Canada, Netherlands, UK utilizes mainly well defined non contaminated wood fractions for pyrolysis. For that reason non contaminated beech wood (white wood) was used as reference feedstock. The organic mass balance of different feedstock pyrolysed with laboratory and PDU scale equipment is presented in Table 3. The screening testes of cocoa shell, fibre sludge and grinding grit demonstrate a poor quantity of bio-oil between 24-30 wt.%. Contrary to that, the charcoal yields are quite high. Due to the low oil yields the amounts of calculated pyrolysis gases are nearly 1.5 to 2 times of the usual gas amount achieved in the beech wood reference experiment. The product yields of the beech wood reference experiment pyrolysed with the laboratory scale equipment are satisfactory compared to the results of the PDU scale pyrolysis.
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Table 3 Organic mass balance (bio-oil includes reaction water). laboratory scale
bio-oil* charcoal pyrolysis gas**
cocoa shell
fibre sludge
grinding grit
rwt.%]
[wt.%i
[wt.%3
reference (beech) [wt.%]
27.71 43.62 28.67
29.85 41.50 28.65
23.94 47.00 29.06
66.08 15.26 18.66
~
* less feedstock moisture;**calculated
PDU scale
bio-oil* charcoal pyrolysis gas
beech waste wood
fibre sludge
HPL
mixed waste wood
reference (beech)
[wt.%]
[wt.%]
[wt.%]
[wt.%]
[wt.%]
46.11 15.02 38.87
9.65 31.26 59.08
44.28 31.62 24.10
75.79 16.28 7.93
67.83 15.50 16.67
* less feedstock moisture The product yields of the laboratory scale pyrolysis of fibre sludge are differing badly in comparison to the PDU scale results. The oil yield of the PDU scale pyrolysis is 3 times lower than the laboratory scale pyrolysis. In fact the PDU scale data are the more reliable ones. Yields of 8.4 wt.% bio-oil and 25.8 wt.% charcoal formed by fibre sludge pyrolysis can be also found in earlier studies [ 181. The organic amount of bio-oil obtained by beech waste and HPL pyrolysis is lower than reference or mixed waste wood pyrolysis but approximately 4.5 times higher than the fibre sludge oil yield. The extremely high oil yield of mixed waste wood is caused by high reaction-water amount included (29.8 wt.%). These data could lead to the premature conclusion that, except fibre sludge, all investigated feedstocks can easily be pyrolysed. However, this is only half the truth. The total mass balances Fig. 6 and Fig. 7 are indicating the importance of the ash content. The complete separation of inorganic solids from the waste is almost impossible if the ash content is very high. Feedstock such as grinding grit with 7.8 wt.% or beech waste with 4.19 wt.% of inorganic material are difficult to handle in fluid bed.reactors. Additionally, they lead to a more or less poor yield of bio oil below 25 wt.% calculated on dry feed. Mixed waste wood is forming an exception in the whole strinx of investigated feedstock. The inorganic solids with an amount of 5.25 wt.% did not cause any problems during test runs. The oil yield is only 10 wt.% lower than the reference.
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70
i
65.75
60 48.0
50 -40 s
-30 $ 20 10
0 bio-oil*
charcoal
El cocoa shell
pyrolysis gas** inorganic solids
fibre sludge Elgrinding grit Eareference (beech)
*less feedstock moisture, **calculated
Fig. 6 Mass balance of several feedstocks investigated by laboratory scale pyrolysis. 67,4
70 60
48,s 50
-40
s
32,7
30,8
5; “30
30,5
20 10 0
1
bio-oil* beech-waste
charcoal
pyrolysis gas
fibre-sludge 0HPL
inorganic solids
waste-wood Hreference (beech)
*less feedstock moisture
Fig. 7 Mass balance of several feedstocks investigated by PDU scale pyrolysis.
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Pyrolysis plants with an integrated liquid quenching system normally operate the quencher (Fig. 2 ) at temperatures between 20-30 "C. Due to that procedure the water phase (with organic compounds included) can not be collected in the condensation system of pyrolysis plants. Therefore an intensive cooler was installed after the liquid collection system of the pilot plant displayed in Fig. 2. This phase is normally burned with the pyrolysis gas and does not enter the mass balances of other existing plants. It consists of volatile organic compounds and water. The balance of organics and water in the water phase is presented in Table 4. The organic residue in the water phase varied between 9-20 wt.%. If this water phase is not collected and ignored in the mass balance the mistake will be 10 wt.% related to the total oil-mass and 1-2 wt.% related to the organic oil mass. The total mass balance is then closed only to ca. 90 %.
Table 4 Organic residue in the water-phases obtained by pyrolysis of different feedstock
HPL
waste wood
reference (beech)
91.25 water content organic residue 8.75
86.94 13.06
82.78 17.22
feedstock
Table 5 Composition (summary) of bio-oils obtained from different feedstocks. quantified by GC-FID waste beech wt.%
ti bre sludge wt.%
acids alcohols aldehydes aromatics furans guaiacols ketones phenols PYans sugars syringols
5.46 0.18 4.44 0.00 1.49 0.83 6.00 0.20 0.13 3.19
2.20 0.40 0.13 0.00 1.14 0.82 4.29 1 1.94
1.11
2.40 0.04 0.23 0.00 1.73 4.11 1.66 2.73 0.15 0.77 2.29
water total organics total identified
54.21 23.02 77.23
11.38 16.10 27.48
feedstock compound
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HPL wt.%
waste reference wood (beech) wt.% wt.%
0.75 0.78
1.94 5.63 0.39 0.13 6.79 0.03
5.64 0.18 2.83 0.00 1.67 1.37 6.46 0.25 0.14 2.43 4.03
49.08 22.46 71.54
40.30 26.53 66.83
37.35 25.01 62.36
0.01
3.27 0.00 6.25 0.09
2.02
The mixture of organic compounds in bio oil from pyrolysis of different feedstock is shown in Table 5. It can be seen that the amount of phenols is very high in HPL-oil in comparison to others. More than 97 wt.% of these phenols originate from phenolformaldehyde resins from the HPL production process. Normally the phenol components have an amount below 1 wt.% in the oil. On the other hand Table 5 demonstrates a large variation in the total amount of each chemical class. All measured water contents are higher than the reference (except fibre sludge). The water content in the bio-oil is important for its quality and value. The amount of reaction water does not only depend on neat pyrolysis reactions but also on secondary cracking reactions which depend on the design of the pyrolyser. The wateriorganic ratio should not be lower than 0.7. An increase to a ratio of 1 will entail an oil precipitation into a water soluble and a high molecular weight fraction. The feedstock moisture has no specific influence on the bio-oil formation. Feedstock moisture can be added to the reaction water and causing high process costs and a lower oil quality. A rough calculation shows the minimum reaction water content in bio-oil oil to be nearly 13 wt.%. This was calculated based on dry hard wood assuming that all cellulose (77 wt:%) is decomposed to levoglucosan (Fig. 8). Thermal splitting
o/---
‘0
H
H
k
-
Product: levoglucosan & H 2 0
OH
OH
Fig. 8 Thermal splitting of 1-4 linked P-D-glucose with H 2 0 emission
CONCLUSIONS Pyrolysis of biomass waste with flash pyrolysis technique has been proved to be possible with all types of investigated feedstock. In comparison with the reference feedstock, beech wood (“white wood”) mixed waste wood, HPL and beech-waste are leading to more or less high bio-oil yields between 57.3 to 38.8 wt.% (based on dry feedstock). Feedstock with a high amount of inorganic solids led to a poor oil yield e.g. fibre sludge’s (5.0 wt.%), grinding grit (22.1 wt.%). Fibre sludge with 48.5 wt.% of inorganics caused technical problems during bio oil precipitation. The pyrolytic behaviour of the particle is of decisive importance. If the particle is shrinking, crumbling quickly to small sizes, the operation of a fast pyrolysis fluid bed process is extreme difficult, sometimes impossible. If these inorganics are not decomposed to dust, they can be separated with the char. Dusty pyrolysis products can only be handled in fluid bed plants with a “blow-through’’ design. These dusts are then precipitated from hot pyrolysis gases with multi cyclone systems.
1385
Beside the technical problem with dust separation, the water content and hence the value of pyrolysis products should be looked at seriously. To receive a bio-oil with a high value, the water content should be as low as possible. The feedstock moisture should therefore be minimized. The amount of reaction water does depend on the neat pyrolysis reaction (min. ca. 13 wt.%, see Fig. 8). In addition to that, reaction water content does also depend on secondary cracking reactions which depend on the design of the pyrolyser (e.g. vapour residence time). Bio oil from pyrolysis of biomass is a complex mixture of organic compounds (Table 5). At the moment there is no possibility to separate individual components economically. The first step to an application of bio oil should be the use of bio-oil as a whole. HPL-oil shows a high amount of phenolic compounds (12 wt.% phenols). The potential of this oil to be used as a basic material for “recycled” phenol resin production should therefore be high. Pyrolysis of feedstock with a high amount of inorganic solids is useful if the value for the solid product is high enough. The recovery of black liquor from pulping processes could be a practicable way. But, if the value of inorganics is to low, pyrolysis of these feedstocks will not be economic due to the high heating and transportation costs in the process. The results of the different biomass waste for flash pyrolysis are leading to some recommendations: (1) Look at the feedstock characteristics (water/organic/inorganic-content and particle size). (2) Check the pyrolytic behaviour of the particle (shrinking, crumbling to small size‘?). (3) Choose the pyrolysis process (e.g. fluidized bed with or without “blow-through’ design). (4) Look seriously at the mass balance of the chosen process and on the phydchem. analysis of pyrolysis products. (5) Calculate the production costs and the product values. (6) Decide if pyrolysis of the feedstock in question is wise or not. REFERENCES 1.
2. 3.
4. 5.
6.
7.
AbfG (Abfallgesetz) ( 1986) Gesetz zur Vermeidung und Entsorgung von Abfallen v. 27.08.1986 5undesgesetzblutt Teil I 44, pp. 1410-1420. TA Abfall ( 1993) Dritte allgemeine Verwaltungsvorschrift zum Abfallgesetz-Teil 2 v. 14.05.1993. In: Abfallrecht, Beck-Texte im dtv, Verlag C. H. Beck, Miinchen (1996), pp. 265-315. Willeitner H. & VoB A. (1994) Gesamtkonzept f i r die Entsorgung von schutzmittelhaltigen Holzern. Umweltbundesamt Abschluljbericht fur das Forschungsvorhaben 145 06 76 ,,Entsorgung“. Volj A. (1997) Aufkommen und Zusammensetzung schutzmittelbehandelter Althiilzer und zhre Entsorgung. PhD thesis, Universitat Hamburg. Kaminsky K., Sinn H. & Jannings J. (1979) Technische Protoypen f i r die Altreifen-und Kunststoffpyrolyse. Chem.-1ng.-Tech. 51,4 19. Meier D., Ollesch T., Gerdes Ch. & Kaminsky W. (1998) Herstellung von Bioolen aus Holz in einer Flash-Pyrolyseanlage. In: DGMK-Fcrchbereichstcrgzing: Energetische und stoffliche Nutzung von Abfillen und nachwachsenden Rohstojfen, 9802, pp. 83-90. Wehlte S. H. C. (1997) Untersuchungen zur Wirbelschichtpyrolyse von nicht naturbelassenem Holz. PhD thesis, Universitat Hamburg.
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8.
9.
10.
11. 12. 13.
14. 15. 16. 17. 18.
Meier D., Wehlte S. H. C., Simon C. M. & Ollesch T. (1998) Sroffliche Verwertung von nicht naturbelassenen Holzresten durch Pyrolyse in derWirbelschicht. DBU Abschluljbericht zum Projekt 03 63 1. Meier D., Ollesch T. & Faix 0. (2000) Fast Pyrolysis of Impregnated Waste Wood-the Fate of Hazardous Components. In: Progress in Thermochemical Biomass Conversion, (Ed. by T. Bridgwater), submitted for publication, Blackwell Science Ltd, Oxford, UK. Meier D., Wehlte S., Wulzinger P. & Faix 0. (1996) Upgrading of bio-oils and flash pyrolysis of CCB-treated wood waste. In: Bio-oil Production and Utilization-Proceedings of the 2nd EU-Canada Workshop on Thermal Processing. (Ed. by Bridgwater AV & Boocock DGB) CPL Press, UK, pp. 1402I412 Wehlte S., Meier D. & Faix 0. (1995) Wood waste management using flash pyrolysis in a fluidized bed. In: Proceedings of the Workshop-Frontiersof Pyrolysis. Breckenridge, CO, June, NREL, USA. Burkhart H. (2000) Personaf communication. Stora Kabel GmbH, Schwerter StraOe 263,58099 Hagen. VDP (1998) VDP-Statistiken. In: Papier'98, Ein Leistungsber-icht.(Ed. by Verband Deutscher Papierfabriken e.V.), VDP, Bonn, pp. 34-59. IS-LOR (06/2000) Price list for ground chemicals. ICIS-LOR-Group worldwide Wehlte S. (1997) Untersuchungenzur Wirbelschichtpyrolyse von nicht naturbelmsenem HoIz PhD thesis, Universitat Hamburg. Meier D. & Scholze B. (1997) Fast pyrolysis liquid characteristics. In: Biomass, Gasification cind Pyrolysis, (Ed. by M. Kaltschmitt & A.V. Bridgwater), CLP Press, UK, 43 1. Scholze B. & Meier D. (2000) Characterization of the water-insoluble fraction from pyrolysis oil (pyrolytic lignin). Part I. Py- GUMS, FTIR, and functional groups, J. Anal. Appl. Pyrolysis, in press. Y ing Y.(199 1) Pyrolyse von kommunalem und papierindustriellern Klarschkamm in der Wirbelschicht PhD thesis, Universitat Hamburg.
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Co-pyrolysis of wood biomass and plastic wastes of different origin under the pressure of argon and hydrogen B.N.Kuznetsov", V.I. Sharypov", N.G.Beregovtsova", N.Marinb and J.V.Weberb a Institute of Chemistry and Chemical Technology SB RAS, Krasnoyarsk, R USSIA The Laboratory of Industrial Chemistry, University of Metz, Sa in t-Avold, FRANCE
ABSTRACT Influence of wood biomass and plastic wastes origin, iron catalyst and co-processing parameters on the products yield and composition was studied. Experiments were carried out in a rotating autoclave in argon and hydrogen atmosphere at temperatures 340-460 "C and initial pressure 0.3-0.4m a . Pine wood, beech wood, cellulose, hydrolytic lignin, polyethylene, isotactic and atactic polypropylene were used as starting materials. Mechanically activated iron-containing ore materials were used as catalysts. Methods of GC, GC-MS,TLC, 'H and I3CNMR and chemical analysis were applied for products investigation. Some regularities describing the influence of co-pyrolysis process operating parameters, nature of wood biomass and plastics on the yield and composition of liquid products were established and discussed. Obtained data indicate that the optimum temperature of biomass/plastic mixtures conversion which corresponds to the maximum yield of liquids is 390-400"C. Hydropyrolysis process gives the higher degree of mixture conversion and higher yield of light liquids as compare to pyrolysis in an inert atmosphere. Observed in some cases non-additive effects indicate that the interaction between wood and plastic derived products takes place during mixture thermal treatment. The more pronounced synergistic effects were detected for hydropyrolysis process. Iron catalysts promote the formation of liquid hydrocarbons from biomass/plastic mixtures and influence on their composition. INTRODUCTION Utilization of wood-biomass residues as well as waste polymers is the important direction of recent research activities. It is known that direct catalytic liquefaction of plant biomass can be used to produce liquid fuels and chemicals [1,2]. Co-pyrolysis and co-hydropyrolysis processes have the potential for the environmentally friendly transformation of lignocellulosic and plastic waste to valuable chemicals. 1388
Polymers like polyethylene and polypropylene contain approximately fourteen percent of hydrogen and these materials could provide hydrogen during thermal coprocessing with other raw materials. There are some data indcating that the mixture of pine wood with polyethylene can be successfully converted into the distillation liquids [3,4]. The yield of these products depends on the process operating conditions. Liquid products obtained by biomadplastic mixtures co-processing shall contain both the products of thermal conversion of each separated component of the mixture and the products of their chemical interaction. Low cost iron-containing ore materials have been intensively studied as catalysts for the hydroprocessing of organic raw materials of different origin. Because the iron ore materials have a moderate level of catalpc activity some methods of their activation have been suggested. Natural minerals can be dispersed into small particles by grinding in tensile-energy apparatus. Such mechanical treatment not only increases the material surface area [5,6] but can change their catalyticalproperties [7]. In the present work the influence of wood biomas and plastic waste origin, iron catalysts and co-processing parameters on the products yield and their composition was studied.
METHODS The feedstock materials used in this work included medum density polyethylene (PE), atactic-polypropylene (aPP), isotactic-polypropylene (iPP), beech wood, pine wood, cellulose and hydrolytic lignin. The size of the wood biomass and the plastic particles was less 0,l mm. Some characteristics of the wood biomass used are given in Table 1 . Table I Characteristics of wood biomass. Wood biomass
Elemental composition (% weight)a
C
H
N
S
0
cellulose
Chemical composition (“A weight) lig- Hemi Water Ether nin cellu- soluble soluble lose
Dry ash-free basis Iron ore samples, containing about 40% wt. Fe in the crystal phases haematite, pyrrhotite and pyrite were used as catalysts. In order to increase the ore catalyst activity they were treated in a tensile energy planetary activator mill (PAM) in the presence of water, as described elsewhere [7]. Such treatment results in the increase of catalysts
1389
dispersion and changes their structure. Specific surface area of modified catalysts was up to 80 m2/g. Liquefaction experiments were carried out in 0,25 1 rotating autoclave. A mixture of dry wood biomass (I1 wt% moisture), plastic and in some cases catalyst (5% wt.) was loaded into autoclave, pressurised with argon up to 0.1 MPa or hydrogen up to 3.0 5.0 MF’a and then heated up 340-480OC. The duration of thermal treatment was 1 h.. After cooling the autoclave, gas products were collected and analysed by gas chromatography technique. Light liquid products (b. p. < 180OC) and water hction were distilled in vucuo. The other products were extracted with benzene. Than the solvent was evaporated, extracted products were dried to a constant weight in vacuo and their yield was evaluated. The degree of the feedstock material conversion was calculated from the difference between the weights of initial mixture and solid residue, insoluble in benzene. G.c. m.s. analyses of light liquid products was carried out using 40 m x 0,3 mm silica phased capillary column coated with SE-30 and connected with VG-70HS mass spectrometer. Heavy liquid products (b.p. > 180OC) were analysed by I3C n.m.r. spectroscopy (“Bruker” MSL-400). For characterization of the heavy liquid products, their were separated on a particular families of compounds by a colume chromatography technique. Educed fractions were analysed by g.c. - IILS. technique.
-
RESULTS AND DISCUSSION PYROLYSIS OF WOOD BIOMASS AND PLASTIC MlXTURES
The influence of pyrolysis temperature on the products yields was studied using mixtures of wood biomass with plastic (from 2:l to 2:l in weight ratio). The conversion degree did not change significantly with temperature variation as it was noted by other researches. Fig.1 shows the main effect of temperature increase for pine wood / polyethylene co - pyrolysis within the range 36OoC - 46OoC is the higher amount of gas formed, whilst water fraction yield was decreased. The maximum yield of liquids (50 wt.%) was obtained at 370-4OO0C.
60
$ 40 20 0
370
400
430
460
temperature C Fig.I Influence of temperature of co-pyrolysis of pine wood / PE mixture (2:1) in the argon atmosphere on the conversion (1) and the yields of light (2), heavy (3) liquid fractions and gas (4).
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The addition of iron ore catalyst did not increase the yield of liquid products and degree of wood / PE mixture conversion. In order to study the influence of the biomass origin on the degree of biomass I plastic mixture conversion and product yields the copyrolysis experiments were carried out at 4OOOC with fixed mixture composition (biomass I plastic ratio 1:l). Table 2 shows, that mixtures (1:l weight ratio) of beech and pine woods with atactic-polypropylene (aPP) give the similar yields of liquid and gaseous products of co-pyrolysis process. In the case of cellulose the increased yield of water and light fraction was obtained. In the runs with lignin the higher yield of char and less yields of water and light fraction were observed.
Table 2 Influence of biomass origin on the yield of products of biomass I a-PP (1:l weight ratio) co-pyrolysis in the argon atmosphere (4OO0C, lh). Biomass
Product yields, % wt.
origin
Light liquid
Heavy liquid
Water
Gas
Char
Cellulose Beech wood Pine wood Hydrolytic lignin
25.0 18.6 23.4
25.6 30.5 24.4
24.0 18.8 17.1
8.1 10.9 11.5
17.3 21.2 23.6
17.1
27.7
12.5
10.3
32.4
Some experiments have been done to study the influence of wood beach I a-PP and cellulose I a-PP mixtures composition on the degree of their conversion and products yield. In all runs the degree of conversion was increased with the growth of plastic concentration. As shown in Fig. 2 the maximum yields of light liquids were observed for the mixture consists of 20 wt.% biomass and 80 wt.% a-PP, showing the synergistic effect. The yields of light liquids fractions reached up 43.5 wt.% for the beech wood I a-PP mixture and 37.1 wt.% for the cellulose I a-PP mixture. These yields are by 2.5 and 2.2 times higher respectively, as compared to the expected ones, calculated as a s u m of light liquid fiactions produced from each separated component of the mixture. 100%
60% 80%
40% 20% 0%
(OW0 7W¶O 50150 2W80 WlOO rnlxtum cornpoaIUon (cc.lluloaahPP). wt n
toom
aono
60160
aono
wioo
mtxtura composltlon IbmachhPP). wt.96
mlight fr. Oheavy fr Owater O g a s
I l i g h t fr. Ohaavy fr. Pwater D g a s
Fig.2 Influence of cellulose I aPP (a) and beech / aPP (b) mixtures composition on the yield of co-pyrolysis products (4OO0C, 1h).
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According to the GC data gaseous products of a-PP pyrolysis contain more than 75% wt. of olefm. In the gaseous products of the beech wood I a-PP mixture (20/80 weight ratio) only approximately 8 % wt.of the unsaturated hydrocarbons were detected. Taking these data into account it is possible to suppose that olefinc products of a-PP thermal conversion react with products of cellulose and lignin depolymerization with liquids formation. This assumption shall explain the synergistic effect, observed in the pyrolysis of the biomass / a-PP mixtures. The origin of plastic has more influence on the distribution of products of biomass / plastic co-pyrolysis. According to Table 3 the yield of light liquids fiom beech-wood / plastic mixtures (1:2 weight ratio) was decreased in the following sequence: isotacticpolypropylene > atactic-polypropylene > polyethylene, whilst the yield of heavy liquids was increased in the same order. Table 3 Muence of differenttype of plastics on the yield of products and degree of conversion in the process of beech wood and plastic (1:1) co-pyrolysis (4OO0C, lh). Gas
Water
Light Liq. Heavy Liq. Char Conversion
i-PP
10,8
17,4
33,6
22,2
16
84
a-PP PE
10,9 10,8
18,7 18,9
18,5 899
30,5 39,8
21,4 21,6
78,6 78,4
HYDROPYROL YSIS OF WOOD BIOMASS AND PLASTICS MIXTURES
In co-hydropyrolysis experiments without catalysts the degree of pine wood! polyethylene mixture (1:1 weight ratio) conversion was 80% wt. and yield of the light liquid fiaction - 23% wt. The addition of iron ore catalyst activated by mechanochemical treatment increased the degree of mixture conversion by 513%. This increase was mainly due to light liquid fraction formation. The variation of catalyst nature (pyrite, pyrrhotite, haematite) influences on the product composition. Pyrrhotite catalyst yields the highest amount of the light fraction (about 40% wt.). The influence of the process temperature on catalytic hydropyrolysis of biomass/plastic mixture was studied in the range 360 - 460°C. Fig. 3 shows that the highest conversion (91% wt.) of the pine wood / polyethylene mixture (1:l weight ratio) was observed at 390°C - 430°C in the presence of activated haematite catalyst. Higher temperatures promote increased yields of char and gaseous products. At lower temperatures a reduced yield of distillate fraction was observed. In comparison with pyrolysis in inert atmosphere the increased yields of light hydrocarbon fractions (by 1.6 - 1.8 times) and increased degree of mixture conversion (by 1.2 time) were observed for hydropyrolysis process. Some hydropyrolysis experiments have been done to study the effect of biomass concentration on the degree of a biomass I plastic mixture conversion and yields of products at 39OoC. In all runs with different types of biomass the degree of conversion was increased with the growth of plastic concentration. Dotted line on Fig. 4 represents the expected degree of pine wood / polyethylene mixture conversion assuming that no interaction between wood and polyethylene derived products takes place during hydropyrolysis. But the experimental data are differed from expected ones, showing the synergistic effect. In all range of a lignite / polyethylene mixture composition the degree of conversiqn was hgher as compared to theoretical one. Maximum synergistic effect
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100
1
1
80 60
3
40
20 0
350
430
390
460
temperature C Fig. 3. Influence of temperature of pine wood / PE mixtures (1 :1) hydropyrolysis in the presence of activated hematite catalyst on the mixture conversion (1) and the yields of light (2), heavy (3) liquid fractions and gas (4).
was observed for the mixture, containing 30-70 wt.% of pine. In this range of mixture composition the observed degree of conversion in the experiment with 3.O MPa initial H2 pressure was higher by 10-15 wt% as compared to theoretical one. The increasing of pressure up 5.0 MPa results in higher conversion degree (Fzg.4). Under these experimental con&tions the mixture conversion riches up to 100% at concentrationof pine-wood more then 40% wt.
100
80
20 0 0
10
20
30
40
50
60
70
80
90
100
polyethylene content wt%
Fig.4Influence of polyethylene I pine-wood ratio on the degree of mixture conversion at 39OoCand initial pressure H25 MPa (l), 3 h4Pa (2) and on yields of light (3), heavy (4) liquid fractions and gaseous products ( 5 ) at hydropyrolysis time 1 h.
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The hydropyrolysis of pine wood at 390°C in the presence of activated hematite catalyst gives the wood conversion degree 54 wt.%. Under the same hydropyrolysis process conditions about 99.5% wt. of polyethylene was converted t o liquid and gaseous products. Brown coal, pine wood with the content of lignin 29.4% wt., pine wood after extraction with dimethyl sulfoxide (with the content of lignin 17.2 % wt.), as well as industrial cellulose were tested in co-hydropyrolysis with polyethylene at the same process conditions. Fig. 5 shows that the degree of pine wood conversion was permanently increased with the increase of polyethylene / wood ratio in the feedstock materials. The addition of 85 wt.% polyethylene to pine wood results in of wood conversion degree 80 wt.%.
80
P
70 60 50 0
10
20
30
40
50
60
70
polyetilene content wt% Fig.5 Muence of polyethylene concentration on the conversion of brown coal (l), pine wood (2), cellulose (3) and pine wood after lignin extraction with dimethyl sulfoxide (4) in the co-hydropyrolysisprocess ( initial H2 pressure 3 MPa, 39OoC, lh). As was noted early, under the action of polyethylene the conversion of pine wood was increased by 1.4 times. After removing 11% wt. of lignin from pine wood by extraction with dimethyl sulfoxide this effect was less (1.2 times). In the runs with cellulose I polyethylene mixtures no positive influence of polyethylene on pine wood conversion was defected. The maximum increase of the conversion degree (1.5 times) was observed for coal I polyethylene mixture. According to GC data the gaseous products of polyethylene catalytic hydropyrolysis contain only alkanes. In the runs with polyethylene / biomass mixture at similar conditions carbons oxides were detected in gaseous products and the lower yields of C2 C4alkanes was observed. GC-MC data show that the light liquid fractions of plastic 1 biomass co hydropyrolysis contain mainly normal paraffins C7-CI~. Their content was 75% for pine wood / i-PP mixtures. Alkyl derivatives of benzene were also detected in the light liquid fractions. The content of unidentified substances was 15%, alkylbenzenes and alkylfuranes compounds - approximately 10%relative. The difference in the composition of gaseous and liquid products of hydropyrolysis of plastic and biomass / plastic mixtures indicates the interaction between intermediate products of natural and synthetic polymers thermal conversion. The iron ore catalysts increase the yields of light liquid products and olefinic hydrocarbons. Producing of
-
1394
light liquid products is promoted by pressurised hydrogen which suppresses the char formation and facilitates the thermal cracking of heavy liquid products. CONCLUSION
Obtained data show that, the mixtures of the different types of the natural and synthetic organic polymers can be successhlly converted with a hgh yield to light distillate fraction by pyrolysis under inert atmosphere and catalytic hydropyrolysis in the autoclave conditions. The optimum temperature of biomass / plastic mixtures conversion which corresponds to the maximum yield of liquids is 390 - 400°C. In the co - liquefaction processes the interaction between products of natural and synthetic polymers thermal decomposition takes place. Iron ore catalysts, modified by mechanical treatment were found to show a catalytic activity in the process of hydropyrolysis of biomas / plastic mixtures. Some synergistic effects were observed in this process resulting in the increase of conversion degree and yield of light liquid fraction and in the decrease of hydrocarbons content in gaseous products. REFERENCES 1. Meier D., Rupp M. (1990) Biomass pyrolysis liquids upgrading and utilisation.
2. 3. 4.
5.
6.
7.
(Ed.by A.V.Bridgwater and G.Grassi), pp. 155-76. Elsevier applied science, London. Kuznetsov B.N. (1990) Catalysis of chemical conversions of coal and biomass. Novosibirsk, Nauka. Gulyurtlu F.P., Gongalves M., Cabrita I. (1994) Co-pyrolysis of plastics with biomass. In: Proc. 8th Europ. Con$ on Energy, Environment, Agriculture and Industry, p. 3 18. Vienna. Doroginskaya A.N., Sharypov V.I., Beregovtsova N.G., Parmon V.N., Kuznetsov B.N. (1997) Combined thermocatalytic conversion of plant biomass and polyethylene. In: Proc. 3th Int. Symp. on Catalysis in Coal Conversion, pp. 3 1924. Novosibirsk, Russia (rus). Pavlyukhin Y.T., Medikov Ya.Ya., Boldyrev V.V. (1984) On the consequences of mechanical activation of zink and nice1 ferrites. Solid State Chem, 53, 155 -60. Popov B.I., Shkuratova L.N., Kuznetsov V.I., Pavl-n Y.T.(1998) Effect of mechanical treatment on some properties of Fe203-Mo03catalysts. React. Kinet. and Catal. Lett., 25, No 3-4,255-9. Sharypov V.I.,Kuznetsov B.N., Beregovtsova N.G., Reshetnikov O.L., Baryshnikov S.V. (1996) Modification of iron ore catalysts for lignite hydrogenation and hydrocracking of coal-derived liquids. Fuel, 75, No 1,39-42.
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Fate of Arsenic After Fast Pyrolysis of ChromiumCopper-Arsenate (CCA) Treated Wood T. Hata*,D. Meier', T. Kajimoto*, H. Kikuchi", Y Imamura* *WoodResearch Institute, Kyoto University, Mi,Kyoto 611-0011, Japan # Institute of WoodChemistry and Chemical Technology of Wood,Federal Research Centrefor Forestry and Forest Products, Leuschnerstc 91, 0-21031 Hamburg, Federal Republic of Germany 'Industrial Technology Center of WakuyumaPrefecture, Ogum 60, Wakayaama 649-62, Japan % S S Alloy, Ltd, Techno Plaza 308, Kagamrjlama 3-13-26, Higashi Hiroshima, Hiroshima, Jupun
ABSTRQCT: The purpose of this study is to find pyrolysis conditions for chromium-copper- arsenate (CCA) treated wood in order to msvdmizethe retention of arsenic in wood charcoal along with a high oil yield. An experimental hcility was built to examine the intluence of process parameters such as pyrolysis temperature and total pyrolysis time. Milled powder was prepared from CCA treated wood. The powder was pyrolized on top of the fluidized bed at temperaturesbetween 300 to 500 "C and during total pyrolysis times of 80 to 3600 s. The concentrationof arsenic in each product was measured by AASand ICP and complete yield balance were established.The yield of gas at low temperature increases drastically at longer pyrolysis time. At higher temperatures more arsenic was volatized. When one type of chemical is used for catching arsenic compounds, there exist optimum pyrolysis conditionsfor the maximum removal of As. The concentrationof arsenic in wood charcoal increases linearly and reaches a plateau level at 400 "C. The concentrationof arsenic in pyrolysis oil drastically increases with low temperatures and longer pyrolysis times. The optimum condition in this study for the pyrolysis temperature and total pyrolysis time is 450 "C and 80 s. There will be good possibility to apply fast pyrolysis techniquesfor the treatment of CCA treated wood because of the minimum secondary reaction and maximum mass reduction. INTRODUCTION Waterborne salts such as chromium-copper-arsenale (CCA) have been used to protect wood from attack by insects or hngi for some time. Substantial amounts of CCA remain in the wood for long periods, which make CCA the preferred agent for wood treatment.Most of the arsenic is present in the treated wood as CrAs04, which is precipitated on cellulose or
1396
complexed with lignin [l]. The disposal of scrap wood is a growing problem not only in Europe and America but in Japan as well. CCA-treated wood waste includes old telephone poles, wooden playground equipment, and timber from the landscape and basement of wooden houses. Most of this waste is disposed of at non-hazardous sites, burned in conventional incinerators, or recycled to wood-based materials. Pyrolyzing the CCA-treated wood may be one solution to the problem. Pyrolysis results in three products, wood charcoal, pyrolysis oil and gas [2]. Since the first report into the volatilization of arsenic during the combustion of arsenic-treated wood in the 19503, many studies have been carried out on the burning of contaminated wood [3]. The percentage of volatilized arsenic increases with temperature and oxygen partial pressure [4]. The losses of arsenic versus range of temperatures and times are 22 % at 400 "C but rose to 77 % at 800 to 1000 "C for 6 hours. Pyrolyzing CCA-treated wood at lower temperature without any oxidizing agent, resulted in a lower loss of arsenic, and thus is a promising approach for disposal. The percentages of volatilized Cu and Cr are thought to be low enough not to be an environmental problem [5]. Although the possible production of bio-oil from CCA treated wood was already rhentioned [2], no study has so far focused on making bio-oil by pyrolysis of CCA-treated wood. Such a fuel could be used as an important source of energy. The purpose of this study is to maximize the fraction of arsenic in wood charcoal while minimizing mass reduction of the bio-oil. Lab-scale pyrolysis was conducted in order to determine mass balances of yield and percentage of arsenic over the total system. The experimental set-up was built to examine the influence of process parameters such as pyrolysis temperature and total pyrolysis time. The optimum combination of temperature and total pyrolysis time, at which the amount of arsenic retained in the wood charcoal is maximized and that in oil is minimized, was tried to be found. MATERIALS AND METHODS
DESCRIPTION OF THE SYSTEM An experimental facility for the pyrolysis of CCA-treated wood powder was built. A general schematic diagram of the pyrolysis system with a fluidized bed was shown in Fig. 1. The middle part of the glass reactor held the fluidized bed forming a cylindrical passage that kept both sand and feedstock and carried the volatile products to the exit. The glass reactor had a diameter of 30 mm and a distance of 100 mm separates the bottom of the fluidized bed and the exit. The major part of the reactor was placed inside the heating system. One end of the reactor lead to a silo through a screw feeder. The other end extended upward above the heating system and was connected with the inlet tube from a N2gas cylinder. The side arm of the glass reactor, in which a wire screen was inserted to filter dust, was connected to a cold trap and to a glass washing-bottle through a cotton filter in series. Inside the cooling trap was liquid COa acetone solution. The washing bottle had 30 ml of acetone base tetra-butyl-ammonium hydroxide (TBAH) with a volume ratio of 2 to 1, which collects arsenic (11) oxides [4]. The reaction temperature was monitored with a thermocouple connected to a temperature controller that regulates the temperature of the heating system. The total pyrolysis time was defined here as the length of time. The glass reactor is inside the
1397
Fig. 1 Schematic diagram of fluidized bed rsactor.
heating system before the target temperature was attained. The complete yield balance was determined with this set up. The calculation of a complete arsenic balance in the system is possiile. PREPaTION OF WOOD SAMPLES The feedstock used in the experiment, supplied by Koshii Preserving Ltd., was western hemlock (Tsuga heiterophylla Sarg, 100 x 100 mm cross Section) treated with type III CCA salt. Type 111 was a CCA formulation containing 47+3 parts by weight of hexavalent chromium as CIQ, 19f2 parts by weight divdent copper as CuO,and 34 k 4 parts by weight of pentavalent arsenic as &Os, of which the concentration was measured to be 7.73,6.42, and 7.18kg/m3, respectively. The CCA-treated pieces were broken down into chips and passed through a hammer mill. The resulting particles were screened through a 2 mm sieve and used for the pyrolysis experiment.
One gram of the CCA treated wood was heated in N2 atmosphere. Wood powder was fed with the screw f d e r . The wood powder entered the glass reactor above the fluidized bed and fell on top of it. Total pyrolysis time is the period h m the start of feeding up to moment when no further smoke was observed. A constant rate of N2flow (2.25 Vmin) was maintained throughout the pyrolysis reaction. N2 passed through the accumpllated wood charcoal residue and swept the
1398
volatiles to the reactor exit. The residence time of the vapors was calculated as 1.9 sec. The gas stream was passed through a wire screen in order to separate fine char particles to separate the solids and light compounds. The smoke caused by the pyrolysis can be characterized as a combination of oil vapor, micron-sized droplets and polar molecules bonded with water vapor. The gas went through a condenser to separate fine char particles with an acetone/dry ice mixture kept at 4 5 "C and a cotton filter to collect aerosols. The pyrolysis condensates were collected and the non-condensable gases went through a washing bottle filled with an acetone-based TBAH solution.
VARIABLESAND ANALYTICAL.METHOD Important variables in the pyrolysis of wood are temperature, heating rate and duration of pyrolysis. The reactor temperature was varied from 300 to 500 "C, and the duration of pyrolysis was 80, 200, 1800 and 3600 s. After the experiment, the glassware, fluidized bed including sand, wire screen and cotton filter were washed with 20 ml of distilled acetone. The yields of char, oil, and gas were calculated from the mass differences of the glassware before and after the pyrolysis. The TBAH solution was analyzed for arsenic using an inductively coupled plasma (ICP) or atomic adsorption spectroscopy (AAS). RESULTS AND DISCUSSION
MASS BALANCE AND YIELD OF PYROLYSIS PRODUCTS Fig.2 shows the mass balance of pyrolysis products. The error bars show the standard deviation. The gas yield was determined by difference. Between 300 and 400 "C, the yield of char was drastically reduced in the pyrolysis of CCA-treated wood, whereas there was a reverse trend for pyrolysis oil. The increase in yield of pyrolysis oil at higher temperatures can be related to the higher heating rates in the wood particles. More pyrolysis oil was produced from treated than untreated wood at 400 "C at 200 s. of pyrolysis time. CCA showed a significant influence on the thermal behavior of wood. The metals may have a catalytic effect on the thermal decomposition of the wood, resulting in more pyrolysis oil [6]. According to the thermal analysis study, thermal decomposition of hemicellulose is shifted to lower temperatures in the presence of CCA resulting in a higher release of volatile compounds [7]. Pyrolysis for 1800 and 3600 s at 350 "C resulted in less oil yield and almost the same amount of wood charcoal. At 450 OC, pyrolysis for 200 s produced more oil than at 80 s, but little difference in char yield. This is probably due to microstructural changes (small droplets) in charcoal occurring as the result of the thermal degradation process [5]. When the total pyrolysis time was less than 200 s, the yield of gas gradually increased with temperature, except for 450 "C. Arsenic had little influence on the yield. The gas yield at 350 "C and 1800 and 3600 s pyrolysis time increased significantly. MASS BALANCE AND PERCENTAGE OF ARSENIC IN PYROLYSIS PRODUCTS Fig.3 shows the mass balance and relative arsenic distribution in pyrolysis products.
1399
The percentage was based on the initial amount of arsenic calculated for each specimen. Any reduction in the percentage of arsenic in wood charcoal was strongly dependent on the temperature. The percentage was increased in pyrolysis oil with increasing temperature. The percentage of As at 450 "C and 80 s. was less than at higher temperatures or longer pyrolysis times. The losses of arsenic increased with higher temperatures. About 20 YOof the arsenic escaped from the TBAH solution at 500 "C.Increasing the temperature may cause secondary reactions, which make it difficult to trap all the arsenic compounds using only the TBAH solution. Approximately 80 to 90 % of the initial arsenic in wood was evolved during pyrolysis. When using the TBAH solution, a reaction temperature of 450 "C or less is preferred to minimize losses of gas including arsenic compounds. The arsenic content in the gas after passing cotton filter increased with temperature up to 400 "C before leveling off at 2 to 3 % at pyrolysis times of 80 or 200 s. The content of arsenic for 1800 and 3600 s was larger at 350 "C. A wider variation reflects some instability in the amount of volatile products. However, the As content in the gas is generally low. At 450 "C it is higher for 80 than for 200 s (approx. 0.05 YO).The difference may be influenced by the gas yield, which is higher at 450 "C for 80 s. The percentage in pyrolysis oil increased with temperature up to 400 "C,then leveled off at 200 s. The shorter pyrolysis gave a higher percentage at 450 "C by around 5 %. An increase in temperature produced no significant increase in the arsenic content. Pyrolysis for 3600 s gave higher arsenic content in the pyrolysis oil at 350 "C. A shorter pyrolysis time resulted in a lower concentration in the oil. The relative portion of arsenic in wood charcoal was over 75 %, and independent of temperature up to 450 "C.It decreased markedly to 60 % at 500 "C. The total pyrolysis time caused no significant difference at temperatures between 350 and 450 "C. It is believed that arsenic is strongly bound to cellulose and lignin [ l ] . The decomposition of both chemical components occurs as temperature or the total pyrolysis time increase, resulting in more arsenic compounds being released. The combination of a temperature between 350 and 450 "C and total pyrolysis time of 200 s and lower is optimal for maintaining arsenic in the wood charcoal under these experimental conditions.
CONCENTRATION OF ARSENIC IN WOOD CHARCOAL.AND PYROLYSIS OIL The concentration of arsenic, which was calculated based on the initial weight of arsenic in wood charcoal, increased linearly up to 400 "C and then leveled off. A longer pyrolysis gave a higher concentration of arsenic at 350 "C. On the other hand, there was no difference at 450 "C between 80 and 200 s. The lower the arsenic content in the bio-oil, the lower the yield of bio-oil. These are thought to be the optimal conditions for the application of the bio-oil. This can be obtained between pyrolysis temperature of 450 and 500 "C(Fig.4). Once the volatile products escaped from wood, the Nz flow swept them from the hot zone to the exit in less than 1 s. In this study, 80 s was the minimum total pyrolysis time at 450 to 500 "C and 200 s at 500 "C. The combination of 80 or 200 s and 450 "C led to the desired result that is high retention of As in char along with a high oil yield. At longer pyrolysis times such as 1800 and 3600 s the products in the hot-zone undergo
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r-----'., wood charcoal
'"4 --
1800
0
3600
I
300
I
350 400 450 Target Temperature ("C) a Pyrolysistime: 200 sec
450
500
b F'yrolyis time: 80sec
Fig.2. Yield of pyrolyne oil and wood charcoal 88 a function of target temperatnre. Note: *,wood charcoal from CCA-treated wood, 0,wood charcod from o r i w wood, A;pyrolyais oil from CCA-treated wood, A;pyrolySis oil from original wood
Total pyrolysis time (sec) 20020018003600200802008020080 100 n
8 8
80 60
m
$
20
8 0
300 350 350 350 400 450 450 500400*50(1 Target temperature ( "c )
Fig. 3 Mass balance of As in pyrolysis products. (* : Original wood)
1401
I
I
I
I
I
I
I
Total .pyrolysis time (eec)
I -
I
300
350 400 450 Target Temperature ("C)
a Pyrolysis time: 200 sec
450
500
b Pyrolysis time: 80sec
Fig.4. Relation between concentration of arsenic in wood charcoal and target temperature. Note: .,wood charcoal from CCA-treated wood
secondary reactions, leading to a higher concentration of arsenic in the wood charcoal. Fig3 shows the relation between the concentration of arsenic in pyrolysis oil and target temperature. The concentration of arsenic in pyrolysis oil increased up to 350 "C, levels off and increased again at 450 and 500 "C. A longer pyrolysis time gave a much higher concentration at 350 "C. The reason for this is again that the products undergo secondary reactions that generate more volatile arsenic. Oil with a high concentration of arsenic is not suitable for any application. From this figure, the optimum conditions appear to be a temperature of less than 500 "C and a total pyrolysistime of 80 to 200 s. Based on data fiom literature, CCA-treated wood should be pyrolyzed at low temperature from the point of safety disposal [8]. However, the results of the present study indicate that a lower temperature and longer pyrolysis led to higher concentrations of arsenic in pyrolysis oil. Therefore, it is concluded that the optimum conditions for pyrolysis with respect to minimizing the concentration of arsenic in oil is 450 "C and 80 s. The application of fast pyrolysis for the treatment of CCA-treated wood seems to be advantageous because of the minimum secondary reactions and maximum retention of arsenic in charcoal.
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n
I
I
' 1
g1400-' rl O 1200 .9
db-
R
a
E
Total pyrolysis time (sec)
-
3600
OOO-
600
1 450
500
Target Temperature ("c) a Pyrolysis time: 200 sec
b Pyrolysis lime: 80sec
Fig. 5 . Relation between concentration of arsenic in pyrolysis oil
and target temperature. Note: A;pyrolysie oil from CCA-treated wood
CONCLUSIONS The pyrolysis system at laboratory scale described in this w o k proved to be suitable for producing bio-oil with a low (>0.04%) arsenic content fiom CCA-treated wood samples. The complete yield balance was determined using the experimental facility with a fluidized bed reactor In principle, treated wood behaves similar to the untreated with respect to the yields of pyrolysis products. However, more pyrolysis oil was produced from treated than fiom untreated wood at 400 "C potentially caused by the catalytic effect of CCA. The gas yield at low temperatme is increased markedly by prolonging the pyrolysis time. A lower content of arsenic in oil was obtained at higher temperature. Although TF3AH was used to trap arsenic compounds from the gases losses of almost 20% were measured. The concentration of arsenic in charcoal increased linearly with temperature before leveling off at 400 "C.The concentration of arsenic in pyrolysis oil increased markedly at low temperature and at long pyrolysis times. It is therefore concluded that the optimal temperature and duration of pyrolysis is 450 "C and 80 s, respectively. The fast pyrolysis technique has potential for the treatment of CCA-impregnated wood because of minimized secondary reactions and maximum arsenic retention are expected.
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ACKNOWLEDGEMENTS This research was supported by a Grant-in-Aid for Scientific Research (No. 10876038) from the Ministry of Education, Science, and Culture of Japan.
REFERENCES 1. Pizzi A,( 1982) The chemistry and kinetic behavior of Cu-Cr-Ash3 wood preservatives. IV fixation of CCA to wood. J. Polym. Sci., 20 739-764. 2. Broek K.V.D:,Helsen L., Vandecasteele C.& Bulck E.V.D. (1997) Determination
3.
4.
5. 6.
7. 8.
and characterization of copper, chromium and arsenic in chromated copper arsenate (CCA) treated wood and its pyrolysis residues by inductively coupled plasma mass spectrometry. Analyst, 122, 695-700. Dobbs A.J., Phil D., Grant C. (1978) The volatilization of arsenic on burning copper-chrome-arsenic (CCA) treated wood. Holzforschung 32, 32-3 5. Helsen L.& Bulck E.V.D. (1997) Release of metals during the pyrolysis of preservative impregnated wood. In: Developments of Thermochemical Biomass Conversion, (Eds. A.V. Bridgewater and D.G.B. Boocock) Vol. 1, pp. 220-228. Blackie Academic & Professional. Helsen L. & Bulck E.V.D. (1998) The microdistribution of copper, chromium and arsenic in CCA treated wood and its pyrolysis residue using energy dispersive X-ray analysis in scanning electron microscopy.Holzforschung 52 607-6 14. Wehle S., Meier D., Moltran J. & Faix 0. (1997) The impact ofwood preservatives on the flash pyrolysis of biomass. In: Developments of ThermochemicalBiomass Conversion, (Eds. A.V. Bridgewater and D.G.B. Boocock) Vol. 1, pp. 206-219. Blackie Academic & Professional. Wilkins E. & Marray F. (1980) Toxicity of emissions from combustion and pyrolysis of wood. Wood Sci. and Technol., 14, 281-288. Hirata T., Inoue M. & Fukui Y. (1 993) Pyrolysis and combustion toxicity of wood treated with CCA. Wood Sci. and Technol.,27, 3547.
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Fast Pyrolysis of Impregnated Waste Wood - The Fate of Hazardous Components D. Meier, T. Ollesch, 0. Faix Federal Research Centre for Forestry and Forest Products, Institute for Wood Chemistry and Chemical Technology of Wood, 0-21027, Hamburg, Germany
ABSTRACT: The behaviour of contaminated wood wastes during fast pyrolysis in fluidized bed reactors has been investigated. The waste samples comprised painted wood, demolition wood, cable drums, fences, window frames, railway sleepers, and wood impregnated with chlorinated organic compounds. The study was aimed at establishing mass balances for inorganic and organic contaminants. A continuously operating small pilot plant with a capacity of 5 k g h and a bubbling fluidized bed reactor was used. All experiments were performed at the same pyrolysis temperature of 475 "C. Wood samples were impregnated with known amounts of both inorganic preservatives containing chromium, copper, boron salts (CCB), complexes of Cu-HDO and A1-HDO, and 10 typical organic compounds such as lindane, pentachlorophenol, dichlofluanide, toluylfluanide, DDT, propioconazole, tebuconazole, permethrine, cypermethrine, and deltamethrine. Additionally, real wood samples from a commercial waste collection site were pyrolyzed. The heavy metals are collected with the char and the liquid bio-oil contains only traces of metals. The real samples show the same behaviour as those prepared in the laboratory. Fast pyrolysis of wood impregnated with A1 and Cu complexes gave similar results: aluminum and copper are predominantly found in the char fraction. The oil contains only traces of these metals. The stability of the organic compounds during pyrolysis was different. DDT was degraded completely, the most stable compounds were propioconazole and tebuconazole. Highly substituted PCDD/F were degraded to less substituted, but more toxic PCDD/F. Almost 99 % of PCDD/F can be found in the oil. INTRODUCTION In Germany there are increasing disposal problems in managing the amount of 8-10 million tons of waste wood which are accumulated each year. More restrictive environmental laws will not allow landfill in the near term. Therefore, the application of new technologies is under investigation. On the other hand, there already exists an increasing market for waste wood. Actual average prices for different assortments are listed in Table 1 (1). As can be seen, prices are slightly decreasing. But still, the price credits might have positive effects on the economics of thermal processing. 1405
Table 1 Average prices in Germany at gate of the wood collector (EURO/t).
H1 (untreated) H2 (treated) H3 1 (contaminated)
1
unshredded shredded not chopped chopped not chopped
April
October
January
May
1999 0 - 25 10-30
1999 0 - 25 10-30 20 - 50 10-35 100 - 250
2000 0-30
2000 0-30 15-45 10 - 45 0-30 60 - 140
10 - 50
I
0--35 100 -250
I
20-40
-
10 45
I
0-30 80 - 160
I
I
Table 2 shows the materials streams for German waste wood. Most of the input material comes from demolition and furniture. About one third is dumped on landfills and more than a quarter exported.
Table 2 Materials streams for German waste wood. OUTPUT (“XI)
INPUT (%)
As German regulations favor materials recycling over simple heat production, fast pyrolysis, which gives a liquid (bio oil) as main product, is getting more and more attention, because bio oil can also be used as a source for chemicals (2). On the other hand, fast pyrolysis has to compete with existing alternatives such as combustion and gasification. Therefore, the quality of the oil from pyrolysis of contaminated waste wood is of paramount importance, if this process is to be considered an attractive alternative approach for costumers such as disposal site operators or waste wood collection sites. Hence, this study aimed at pyrolyzing various waste wood samples with different contaminants, both organic and inorganic components. The fate of the contaminants is presented through mass balances. EXPERIMENTAL MATERIALS AND IMPREGNATION
Beech wood (Fagus sylvatica L.) particles of 1.5-3 mm and a moisture content of 10 % were generally used for the experiments. Impregnation was performed by firstly soaking biomass with the dissolved compounds and secondly applying vacuum for penetration of the compounds into the wood matrix. Where metal-organic complexes were used for impregnation, 10 g of each compound (see Table 3) were used to impregnate 3 kg of wood. 1406
Table 3 Metal-organic complexes used for wood impregnation.
copper HDO
aluminum-HDO
Three kilograms of biomass were impregnated with 3 g of the corresponding single organic contaminant. The list of contaminants tested appears in Table 4. Table 4 List of organic compounds used for wood impregnation. Chemical Name Hexachloro-cyclohexane (lindane)
Pentachlorophenol (PCP)
Dichlofluanide (DCFN)
Toluylfluanide (TFN)
1407
Chemical Formula
'*
C
I
5
(1,1,l-Trichloro-2,2-bis-(pchloropheny1)ethane (DDT)
u 6
Propioconazole
Tebuconazole
8
Permethrine
9
Cypermethrine
10
Deltamethrine
Biomass was also impregnated with an inorganic salt solution (CCB, 2 0 % concentration) consisting of 38 % K2Cr207*2H20,34% CuS04*5H20, 25 % HB03 and 3 % NaHS04.300 ml of the solution were used to impregnate 3 kg of wood. Furthermore, representative samples from a waste wood collection site (Otto DBrner, Hamburg) were subjected to fast pyrolysis and analysis. Two classes, H2 and H3, were selected. H2 contains typically particle boards and window kames, H3 consists of railway sleepers, fences, cable drums, and other materials impregnated with organic chemicals. ANALYTICAL METHODS
Organic contaminants Organic contaminants were analyzed by gas chromatography using a CHROMPACK CP 9000 model and a DB-1 column (30 m x 0.25 mm, 0.32 pm film thickness) with the following temperature programme: initial: 120 "C,heating rate: 5 OC, final 320 OC, hold for 10 minutes. For calibration solutions with two concentration levels (10 and 100 pg/ml), following DIN 38402, were used. Fluoroanthene was taken as internal standard. In order to reduce the influence of the bio oil on the results, a simple clean-up method was applied, using one-way 500 mg silica cartridges. They were conditioned with 2 ml of ethyl acetate. Then 200 pg of whole bio oil was placed on top of the silica 1408
and all the relevant toxic compounds eluted with 2 ml of ethyl acetate. To this solution the internal standard was added and the mixture subjected to GC.
Metallic contam inants Metals were analyzed using dispersive X-ray fluorescence spectroscopy (XRF) on a SPECTRACE 6000. 2 g of sample were used. For metals in char multi-level calibrations within a range of 10-20,000 ppm were carried out. For the liquid samples the range was set between 1 and 10,000 ppm.
Dioxines (PCDDLPCDF) Analysis of PCDD/PCDF was carried out at a certified laboratory (ERGO, Hamburg) using high resolution GCMS methods. Identification and quantification with a dissolution method was performed following VDI procedure 3499. Samples of bio oil and gas prior to combustion were sent to the laboratory. Gaseous contaminants were collected on a special polyurethane cartridge. PYROLYSIS PROCEDURE Fast pyrolysis of waste wood has been studied in a small pilot plant with a nominal capacity of 5 kgh. A schematic diagram is presented in Figure 3.
Fig. 1 Schematic diagram of pilot plant for fast pyrolysis of waste wood. Before starting the experiment the plant was checked for leaks. A pressure of 150 mbar nitrogen was applied and the losses measured over a time of 30 min. The plant was declared sealed when the losses were <20 mbar. After the leak test the plant was flushed with 1.5 m3 of nitrogen and the system heated to the desired temperature within approximately 2 hours. The on-line micro-GC, attached to the system for gas analysis, was calibrated with standards. Wood feeding was from the hopper via a vibration channel to a water-cooled screw feeder running at constant velocity. Throughput was controlled by the vibration frequency. Wood particles entered the reactor just above the distribution plate of the fluidized bed reactor. A weir inside the reactor was used to remove larger charcoal particles. Because of their lower density they sit on top of the sand bed and fall through the pipe into a collection drum.Vapors 1409
and smaller char particles were entrained with the fluidizing gas. Two cyclones separated the fine char particles and the residual vapors were condensed in a condensation train consisting of two tubular water-cooled heat exchangers (operated at 200 and 80" "C,respectively), an intensive cooler operated at -15 "C,and two electrostatic precipitators working at approximately -1 5 kV. The non condensable gases were circulated as fluidizing gas. Surplus gas was burnt in a flare.
RESULTS AND DISCUSSION
No significant differences in the yield of the main products liquid, gas, and char from untreated and treated wood were observed (see Fig. 2).
-
90
%
'0
f
-
50 40
0 4 400
I
450
500
550
600
Temperature ["C]
Fig. 2 Typical yields of fast pyrolysis experiments with the pilot plant.
Oil yields reached almost 70 % at 475 OC reactor temperature. With increasing temperature the gas yields increase and the char yields decrease correspondingly. CCB impregnated wood was pyrolyzed at 500 and 600 O C and the heavy metals copper and chromium were determined in the char and liquid by XRF (see Table 5). The data clearly show that typical heavy metals from wood preservatives are trapped together with the char fkaction. If the char is burnt for energy recovery, the heavy metals would remain in the ash. From here they might be recovered or dumped. The oil is practically uncontaminated. These results are consistent with those found already in a laboratory fluidized bed reactor (3).
1410
Table 5 XRF analysis of Cu and Cr in pyrolysis oil and char. 600 "C CCB with organic preservative
500 "C CCB alone
contaminant
char
oil
char
oil
copper ( P P 4
22560
12
8257
16
chromium (ppm)
3 1561
4
19768
3
XRF analysis was also performed on pyrolysis oil and char from fast pyrolysis of real wood waste obtained from a commercial collection and separation site. The XRF data show similar results compared to the samples impregnated in the laboratory. Various mono-fractions from the collection site were pyrolyzed such as cable drums, fences, railway sleepers, particle boards and window frames. The results are presented in Table 6 and are based on the whole oil whereas the data presented in (4)only refer to the cleaning oil fraction. Nothing could be detected in the oils collected from the condensation train. Table 6 XRF analysis (in ppm) of Cu and Cr from authentic waste wood samples cable drums
fences
railway sleepers char oil
particle boards char oil
window frames char oil
char
oil
char
oil
Cu
6921
0.42
1628
0.48
932
0,49
270
1.64
681
0.74
Cr
14103
1.12
4748
0.14
267
0.11
527
0.36
465
0.21
The metal-organic preservatives copper HDO and aluminum HDO are commonly used for limber used in landscape improvement. Therefore, samples impregnated with these preservatives were also pyrolyzed at 500 "C. Aluminum could not be measured in the oil because the chamber had to be evacuated whch would have contaminated the inner walls with pyrolysis liquids. Therefore, only the char was analyzed for A1 resulting in a recovery of 82.7 %. Hence, 17.3 % of the input aluminum is expected to remain in the oil. This is much more compared to Cu and Cr due to the lower melting and boiling point of Al. In the past, various chloro-organic compounds were used in preservative formulations in Germany. They are highly toxic not only towards bacteria, fungi and insects but also for humans. Therefore, it was of special interest to study their fate with fast pyrolysis. Each beech wood sample was impregnated with one compound from Table 4. In order to obtain measurable amounts after pyrolysis it was decided to use a high concentration of 1 g/kg which was thought to represent the worst case. In practice this concentration is unrealistic as no discrimination between the different organic preservatives can be made at the wood collection site.
141 1
Table 7 Recovery rate (%) of chloro-organiccompounds based on initial concentration Temperature "C ~
~~
450
500
550
580
Lindane
33.06
29.84
16.33
2.02
PCP
42.18
25.13
25.03
13.04
DCFN
10.65
7.45
3.13
2.40
Toluylfluanide
18.72
17.00
13.10
7.92
Propioconazole
40.47
26.92
23.80
16.66
0
0
0
0
Tebuconazole
40.16
28.11
26.19
15.93
Permethrine
32.51
22.76
13.51
6.83
Cypermethrine
4.26
0.68
0.47
0.32
Deltamethrine
3.00
2.10
1.61
0
DDT
Pyrolysis experiments were conducted at different temperatures: 450, 500, 550, and 580 "C. Quantitative analyses of each compound were performed by GC using the internal standard method. A special clean-up procedure with silica cartridges was used to enrich the compounds in the sample solution and to reduce matrix effects from the pyrolysis liquid. Generally speaking, the concentrations of the contaminants decrease with higher pyrolysis temperatures. PCP, HCH, propioconazole, tebuconazole, and permethrine are destroyed by 60-70 % already at the lowest pyrolysis temperature of 450 "C. At 580 "C the reduction reaches 80-95 "% based on the initial concentration (see Fig. 3). The pyrethroids deltamethnne and cypermethrine as well as toluylfluanide and dichlofluanide are even m h e r destroyed. At 450 "C they lost 80-98 % of their initial concentration and at 580 "C destruction rate is between 90 and 99 % (see Fig. 4). DDT is already completely destroyed at 450 "C. In general, the presence of chloro-organic preservatives in thermal processing makes the formation of polychlorinated dibenzodioxins and polychlorinated dibenzofiuans (PCDDR) highly probable. Their formation rate is highest at around 300 "C 5. In contrast to polyaromatic hydrocarbons (PAH), PCDD/F are formed especially at temperatures below 600 "C (deNovo synthesis). A prerequisite for reaction is the presence of char particulates, chlorine particulates and gaseous oxygen (6). As fast pyrolysis occurs in an inert atmosphere the formation of PCDDR might be reduced. To test this thesis, two experiments were camed out: one blank run with beech wood and another with beech wood spiked with
1412
45 40 35 J(t Tebuconazole
30
3
E
z?
25
$! 20
s
3 15 10 5
0 450
600
550
500 Temperature ("C)
Fig. 3 Decrease of recovery rate with increasing temperature of selected chloroorganic compounds
18
16 14
+I+ Toluylfluanide -ACypermethrine
+Deltamethrine
1
450
500
550
600
Temperature ("C)
Fig. 4 Decrease of recovery rate with increasing temperature of selected chloroorganic compounds CCB salt and technical grade PCP containing high concentrations of substituted heptaand octa-PCDDE (see Table 8). The highly substituted compounds are degraded in favor of less substituted, but more toxic compounds. Therefore, the toxicity of product bio-oil increased by a factor of 2.57 as demonstrated by the increase of the International Toxicity Equivalent (I-TE). However, it is not clear whether a de novo synthesis took place or just recombination reactions. From the total PCDD/F content the oil contains 98.65 %, the
1413
char has 1.34 and the gas has only 0.01 %. If the oil is used as fuel one has to carefully select the combustion conditions in order to destroy the PCDD/F completely. The results presented here show the general trend if chloro-organic compounds are present in the feedstock. The amount of chloro-organics in the feedstock used here were somewhat unrealistic but selected to better understand and measure PCDDR formation and concentration in the products of fast pyrolysis. In reality, there are no mono-fractions of wood contaminated only with chloro-organic compounds. Generally, pyrolysis of waste wood mixtures results in a much less contaminated oil as presented in this study. However, further investigations are necessary to prove if de novo synthesis takes place during fast pyrolysis of contaminated waste wood. CONCLUSIONS
The results of this study show that fast pyrolysis of contaminated waste wood is technically feasible without any restrictions. Heavy metal compounds introduced by wood preservatives are almost completely mechanically separated by cyclones together with the charcoal. The destruction rate of chlorinated organic compounds depends on the temperature, the higher it is the higher the extent of degradation. At 580°C the recovery was between 0 and 16.7 %. Polychlorinated dioxins or furans present in preservatives are partly destroyed. However, more toxic compounds with a lower degree of chlorine substitution are formed. Thus, high temperature combustion of these liquids is necessary to meet the German limit of 0.1 ng/m3. It should be noted that for the sake of analysis worst case samples were pyrolyzed in order to get measurable amounts of substances. In reality much less contamination is expected. No organic preservatives (except from railway sleepers) could be detected in the oil from fast pyrolysis of realistic waste wood mixtures. ACKNOWLEDGEMENTS
We acknowledge the ftnancial support of Deutsche Bundesstiftung Umwelt, Osnabriick, project no. 03 63 1
1414
Table 8 Concentration of PCDD/F in feedstock sample and pyrolysis products
Hexa- 1 Hexa-2 Hexa-3
total PCDD
PCDF Tetra Penta- 1 Penta-2 Hexa- 1 Hexa-2 Hexa-3 Hexa-4 Hepta- 1 Hepta-2 Octa total PCDF total PCDDRCDF I-TE n.d. = not detected REFERENCES Speckels, L., BFH, Institute for Wood Physics, personal communication (2000) Bridgwater, AV, Peacocke, GVC. (2000) Fast pyrolysis processes for biomass. Renewable and Sustainable Energy Reviews, 4, 1-73. Wehlte S., Meier D., & Faix 0. (1997) The impact of wood preservatives on the flash pyrolysis of biomass. In: Developments in Thermochemical Biomass Conversion, (Ed by A.V. Bridgwater & D.G.B. Boocock), pp.206-219. Blackie Academic & Professional, London. Bridgwater AV, Meier D, Radlein D. (1999) An overview of fast pyrolysis of biomass. Org. Chem., 30, 1479-1493.
1415
5
6
Kolenda J., Gass H., Wilken M., Jager J., Zeschmar-Lahl B.(1994) Determination and reduction of PCDDFCDF emissions fiom wood burning facilities, Chemosphere 29, 1927 -1938. Mohr K, Wilken M, Nonn Ch, Jager J. (1995) Behaviour of PCDD under pyrolysis conditions, Organohalogen Componds, 23, 293-297.
1416
Low-temperature pyrolysis as a possible technique for the disposal of CCA treated wood waste: metal behaviour L. Helsen and E. Van den Bulck Katholieke Universiteit Leuven, Department of Mechanical Engineering, Celestijnenlaan 300 A, B-3001 Heverlee, Belgium
ABSTRACT: Low-temperature pyrolysis is evaluated as a possible technique for the disposal of CCA treated wood waste. A theoretical and experimental study of the lowtemperature pyrolysis of CCA treated wood waste is performed in order to gain more insight in: (1) the metal (Cr, Cu, As) behaviour during the pyrolysis process and (2) the influence of CCA on the pyrolysis process. The experimental study focuses on the determination and characterisation of Cr, Cu and As in CCA treated wood and its pyrolysis residue and the study of the pyrolysis process. Based on the experimental observations some important conclusions are drawn with respect to the metal behaviour during the pyrolysis process. Furthermore, kinetic models are derived for the lowtemperature pyrolysis of CCA treated wood and the As release during the process. CCA is identified as a promoter of charcoal production and pyrolysis temperature and residence time are the most important parameters with respect to metal release during the pyrolysis of CCA treated wood waste. At the low temperatures considered, Cu and Cr do not pose any problems, As on the other hand is already released at temperatures between 300 "C and 400 "C. The mechanism responsible for As release is identified as the reduction of As(V) to As(III), which occurs sharply at 327 "C. Arsenic is probably released as As406,which is very difficult to capture and toxic. Due to the high leachability of As and Cu in the pyrolysis residue, this residue can not be landfilled without a pre-treatment. The presence of metal agglomerates in the pyrolysis residue, however, may allow a dry separation of the metals and the charcoal.
INTRODUCTION Waterborne salts have been used to preserve wood from insects, fungi and water damage for many years. One of the more common formulations contains copper, chromium and arsenic salts and is known as chromated copper arsenate or CCA. After impregnation of the wood with a CCA solution the metal compounds will be fixed to the cell walls of the wood matrix. Substantial amounts of CCA remain in the wood for many years and the disposal of scrap wood is a growing'problem in Europe and in the United States. Telephone poles, railway sleepers, timber from landscape and cooling towers, wooden silos, hop-poles, cable drums and wooden playground equipment generate wood waste for which environmental benign disposal technologies need to be developed. 1417
Pyrolysing the CCA treated wood at low temperature is a promising solution to the growing disposal problem since low temperatures and no oxidising agents are used. In turn, this leads to a smaller loss of metals than in combustion or even a total recuperation of the metals. However, the current literature clearly shows that the mechanism of metal volatilisation during the thermal decomposition is not yet completely understood. Fast pyrolysis of impregnated waste wood is studied by Meier [l, 21. In the present study, a low-temperature pyrolysis facility for CCA treated wood waste is optimised with respect to a minimal metal release upon a maximal mass reduction. During this study several experiments have been carried out with the aim to gain more insight in the behaviour of the metals (Cr, Cu, As) during the pyrolysis process. Among these experiments are: Determination of metal mass balances over the total system using inductively coupled plasma mass spectrometry (ICP-MS). Sequential extraction of the pyrolysis residue to study the leaching behaviour of the residual solid matrix. Speciation study of As in the pyrolysis residue using hydride generation coupled to ICP-MS (HG-ICP-MS). Study of the metal microdistribution in CCA treated wood and its pyrolysis residue by energy dispersive X-ray analysis coupled to scanning electron microscopy (SEM-EDXA). Labscale pyrolysis study: examination of the influence of temperature and residence time on the release of metals and resulting mass reduction. Thermogravimetric analysis (TGA) of CCA treated wood and the major arsenic compound in CCA treated wood (chromium arsenate (CrAsO,)).
In this paper, the most important experimental observations with respect to metal behaviour during pyrolysis of CCA treated wood are outlined and discussed. Based on the experimental observations and the new insights resulting therefrom, a reaction scheme for the pyrolysis of CCA treated wood and a kinetic model for the As release during pyrolysis have been developed. A critical evaluation of these kinetic models results in a hypothesis concerning the behaviour of metals, especially As, during pyrolysis of CCA treated wood. EXPERIMENTAL STUDIES DETERMINATION OF METALS THROUGH ANALYTICAL TECHNIQUES Fixation of CCA to wood The mechanism of interaction of the CCA preservative with the wood structure was initially elucidated by Dahlgren [3, 41 and Dahlgren and Hartford [ 5 ] . Subsequently, this scheme was modified by Pizzi [6, 7, 8, 91, who studied the kinetics of reactions between CCA or CCA component elements and cellulose or lignin model compounds. The kinetic data combined with results from analysis of reaction precipitates were used to determine the distribution of preservative elements between the major compounds in the wood structure. A summary of the products formed when CCA reacts with wood is presented in Figure 1.
1418
CuCr04 CrAs04
+ stable CuCr04/ lignin complexes
c
CrAs04/ lignin complexes inorganic CrAs04 precipitates on cellulose
Cr(OH)4Cr04 + inorganic precipitates on cellulose CU'+ / lignin complexes c u 2 + E
physically adsorbed on wood constituents CU*+/ cellulose complexes
HCrOi
+ HCrOi / lignin complexes
Fig. 1 Scheme of the interaction of CCA with wood. As can be seen, the CCA interacts mainly with the lignin and cellulose components of the wood structure, resulting in complexation, precipitation or adsorption. The fundamental reaction during the fixation process is the oxidation of wood components by Cr(V1) resulting in the reduction of Cr(V1) to Cr(II1). This can be described as an interaction of CCA with the functional groups (carbonyl, carboxyl, methoxy and phenolhydroxyl groups) of the wood components [lo], during which Cr is partly reduced, As and Cu are being fixed and some wood components are oxidised. The products formed when CCA reacts with wood are composed mainly of the CuCr04-lignin complexes, CrAs04-lignin complexes and CrAs04 precipitates on cellulose. These compounds make up 95 % of the system CCA-wood (excluding Cu2+ unreacted with chromium). Copper arsenates are not involved as final fixation products unless a preservative of high As content is used. The actual proportions of the various chemical species depend on the relative proportions of Cu, Cr and As in the preservative formulation.
Determination of the metal content To determine the total amount of Cr, Cu and As in CCA treated wood and its pyrolysis residue by inductively coupled plasma mass spectrometry (ICP-MS), the metals have to be brought in solution. Several leaching and dissolution methods were tested and compared to find the optimal procedure [ l l , 121. To determine the total amount of Cr, Cu and As in the dried CCA treated wood, a leaching procedure using H202 and H2S04 (BSI method) seems to be the most suitable method: reproducible, easy to handle and it guarantees that all metals are brought into solution. Since Cr is not completely dissolved when applying the BSI method to the pyrolysis residue, a more aggressive leaching procedure or a dissolution method (for example the dissolution in concentrated HN03) is necessary. This indicates that the Cr in the pyrolysis residue is even more strongly bound or less easily leached compared to Cr in the CCA treated wood.
1419
Sequential extraction A sequential extraction procedure, consisting of five steps, was applied to the pyrolysis residue [ 11, 121. The different extraction steps correspond to different fractions, sequentially liberated in more aggressive environments. This extraction procedure gives information about the mobility of the metals if the pyrolysis residue has to be disposed on a landfill. About 30 96 of both active compounds (As and Cu) retained in the pyrolysis residue is relatively easy leached in a neutral environment. The leachability of As and Cu will thus pose problems upon landfilling. The fixing agent (Cr) on the other hand, is only leached in a strongly oxidising or strongly reducing environment and will pose no problems upon disposal on a landfill.
Speciation study A hydride generator was coupled on-line with the ICP-MS system to determine the oxidation state of As in the first two extraction steps (neutral environments) [ll]. Identification of As(II1) in the reducing leach in the third step and the oxidising leach in the fourth step does not provide evidence for the valence state of As in the pyrolysis residue. Therefore, the leachates of these steps were not subjected to hydride generation. This speciation study shows that the major part of As in the pyrolysis residue is present in the trivalent state. The presence of As(II1) in the pyrolysis residue indicates that the As, present as As(V) compound (CrAs04) in the CCA treated wood, is partly reduced to As(II1) during the pyrolysis process. If the pyrolysis residue would be landfilled, mainly As(III), which is the more toxic form, will be liberated into the environment. Hence, the pyrolysis residue of CCA treated wood waste can not be landfilled without pretreatment.
Micro-analytical study Scanning electron microscopy coupled with energy dispersive X-ray analysis (SEMEDXA) was used to study the preservative distribution and the anatomical structure at the micro-level in CCA treated wood and its pyrolysis residue [13]. Comparison of the micro-distribution between CCA treated wood and its pyrolysis residue provides information about what happens to the wood metal bonding during the pyrolysis process. SEM-EDXA shows that the micro-distribution of metals is not necessarily the same before and after pyrolysis. The metal compounds may have migrated towards the ray cells during the pyrolysis process, the radial way through the ray cells being the main pathway of preservative penetration in softwoods. Furthermore, the metals appear as agglomerates rather than isolated elements (see Figure 2). These observations suggest the possibility of the use of a pneumatic centrifugal grinder-separator system, which aims at recovering clean charcoal, based on the density difference between the carbon and the metal particles [ 141.
1420
Fig. 2 Longitudinal radial sectional view of ray cells in the pyrolysis residue of CCA treated Pinus sylvestris sapwood ( 4 0 0 ~ )with metal agglomerates at the locations 1, 2 and 3 in the ray parenchyma cells.
EXPERIMENTAL PYROLYSIS STUDY Labscale pyrolysis experiments An experimental facility for the low-temperature (T < 500 "C) pyrolysis of CCA treated wood particles (between 2 and 135 mm long, between 2 and 17 mm wide and between 0.5 and 2 mm thick) was built with the aim of maximising the fraction of metals that is contained in the ash upon a maximal burnout of the wood. Two different configurations were tried: updraft [ 151 and downdraft [ 11, 12, 141 fixed bed. In the updraft system the mixture of pyrolysis gases and nitrogen is cooled in a highly compact coiled water-cooled heat exchanger before passing through a filter. Different types of filters were tried in order to capture the As and tar compounds as much as possible. The influence of pyrolysis temperature and residence time of the wood chips on the metal content of the resulting pyrolysis residue and the mass reduction was determined. Therefore, the reactor temperature was varied between 300 "C and 450 "C and the duration of the pyrolysis process between 20 and 60 minutes. More details of these pyrolysis experiments are given elsewhere [ 151. Since the liquid tar droplets and heavy gaseous compounds condensed on the spiral heat exchanger and formed there a dense sticky layer that could not be removed, nor analysed, the calculation of a complete metal mass balance over the system was impossible. To overcome these problems a downdraft system was built. The gas flow, leaving the downdraft pyrolysis reactor, is quenched in a water scrubber. The resulting aerosol is then forced through a glass fibre filter. This gas cleaning system has later been optimised to release as little metals as possible: the aerosol at the outlet of the water scrubber is forced through a plate-column, a liquid-gas separation vesseli an extra ice-water condenser and finally through a tube filled with cotton wool that acts as a filter. A metal mass balance was calculated over this system for the three metals (Cu, Cr and As). This downdraft pyrolysis system is described in more detail by Helsen et al. [ 11, 12, 141. 1421
The influence of reactor temperature and duration of the pyrolysis process on the weight percentage of As in the pyrolysis residue and on the mass reduction of the wood sample is illustrated in Figure 3. Since the same trends are observed for the two other metals (Cr and Cu), the retainment of these metals in the pyrolysis residue is not presented here.
t
90 80
300
350
400
reactor temperature (“C)
‘“I
450
500
I
Fig. 3 Percentage As retained in the pyrolysis residue and percentage mass reduction relative to dried wood as a function of the mean reactor temperature for different residence times (0for 20 minutes, x for 40 minutes, + for 60 minutes).
Figure 3 shows that the As content in the pyrolysis residue decreases as the temperature or duration of the process increases. The mass reduction is influenced in the opposite way. The influence of both parameters on the release of Cu and Cr is the same, but the
1422
amounts retained in the residue are higher. The arsenic compounds are thus more volatile than the copper and chromium compounds. The optimal combination of pyrolysis temperature and duration with respect to minimal metal release and maximum mass reduction, resulting from these experiments, is 350 "C and 20 minutes. At these conditions, the release of Cu and Cr is negligible, while the As release is about 25 %, using nitrogen at a flow rate of 5 Nm3/h. The corresponding reduction in mass is 56.7 % relative to dried wood. The percentages of Cr, Cu and As found in the CCA treated wood, the pyrolysis residue, the scrubber liquid, the glass wool filter and the gas stream obtained in a typical pyrolysis experiment using the downdraft fixed bed reactor are given in Table 1.
Table 1 Metal content (wt %) of the different products used as input or obtained as output in the pyrolysis process.
Cr (wt %)
c u (wt %)
As (wt %)
100
100
100
96.9 k 2.5 0.051 f 0.001 0.0092 f 0.0002 3.0 f 2.5
94.1 k 1.8 3.6 f 0.08 0.021 _+ 0.0005 2.3 f 1.9
85.4 f 2.0 3.3 f 0.06 4.0 f 0.08 7.3 k 2.1
product
dried wood pyrolysis residue scrubber liquid filters gas (losses) *
* calculated by subtraction The dried wood was heated according to the following reactor temperature profile: hold at 200 "C for 10 minutes, ramped to 350 "C in 20 minutes and hold at 370 "C for 10 minutes. The resulting mass reduction was 57 96 relative to the dried CCA treated wood. The metal content of the gaseous product (losses) was calculated by difference, while all other products were analysed by ICP-MS. From Table 1 it can be seen that low-temperature pyrolysis of CCA treated wood (with a precisely controlled temperature and residence time of the wood particles) may be a promising technique for this type of wood waste. The percentages Cu and Cr volatilised are low enough to conclude that they will pose no problems under the process conditions used. The amount of As volatilised, on the other hand, can not be neglected: it is partly captured in the water scrubber, partly by the filters, but a nonnegligible amount (about 7 %) still escapes with the gas flow. Arsenic can thus be identified as the problematic compound with respect to metal release.
Thermogravimeaic analysis A thermo-analytical study of untreated and CCA treated wood samples (more or less cylindrically shaped with diameter less than 2 mm) was performed in order to examine the influence of the presence of CCA on the pyrolytic behaviour of wood samples (TG study), as well as the release of metals (Cu,Cr and As) during the pyrolysis process in the kinetically controlled regime. Since arsenic is the most problematic compound during pyrolysis of CCA treated wood and almost all arsenic is present in the treated wood as CrAs04 (precipitated on cellulose or complexed with lignin) [9], a 1423
thermogravimetric analysis was also carried out on the compound chromium(II1) arsenate. TG analysis was performed using the apparatus and samples described in [ 16, 171. Comparison of the TG and DTG curves obtained for untreated and CCA treated wood (see Figure 4, corresponding to slow heating experiments) shows that CCA has a significant influence on the thermal behaviour of wood. 110
1
OO
1
100
200
° 300
400
500
1
temperature ("C)
Fig. 4 TG and DTG curves for untreated wood (dashed line) and CCA treated wood (solid line for CCAl and dash-dotted line for CCA2 with CCA2 retention > CCAl retention): slow heating (10 "Chin) experiments.
The temperature at the onset of pyrolysis as well as the temperature where the maximum rate of decomposition occurs are lowered, the rate of weight loss is much
1424
more peaked (the shoulder at the low-temperature side of the DTG curve has disappeared) and the final char yield is higher for CCA treated samples. A promoting action of the CCA compounds upon the pyrolysis reactions which favours the formation of char can be postulated. Based on the TG study of hydrated chromium(II1) arsenate, which is the major arsenic compound in CCA treated wood, and on the results of equilibrium calculations carried out by other researchers [ 18, 19, 201, it is concluded that the decomposition of hydrated chromium arsenate results in the formation of solid Cr203and gaseous H20, O2and As406,according to the scheme: 4 CrAs04.6H20 (s) 2 As205 ( s )
+ 24 H20 (g) + 2 Cr203(s) + 2 As205 (s)
+ 2 As203 (I) + 2 0 2 (8)
This scheme suggests that arsenic is released in trivalent state, as As406. From the study of the metal (Cr, Cu and As) release during pyrolysis, it can be concluded that the metal volatilisation is strongly dependent on both temperature and residence time of the wood sample at a given temperature. A critical point (10 minutes at 400 "C) is identified, below which the release of Cr and Cu is negligible and the release of As is lower than 10 %. Above this critical point (longer times at 400 "C), there is a dramatic increase in metal release for all three metals. Arsenic is again recognised as the most problematic metal compound, being released (or volatilised) at temperatures between 300 "C and 400 "C. These observations suggest that the release of As is not solely made up by processes that are governed by thermodynamical equilibrium principles, such as sublimation and evaporation. The time dependence suggests that processes that involve chemical kinetics are substantial to the As release mechanism as well.
THEORETICAL STUDIES: KINETIC MODELS LOW-TEMPERATURE PYROLYSIS OF CCA TREATED WOOD WASTE
A least square evaluation of the DTG curves has led to an acceptable mathematical description of the thermal decomposition of untreated and CCA treated wood [17]. The low-temperature pyrolysis of untreated wood is described by a scheme of three independent parallel reactions, while a combination of two independent parallel reactions and one subsequent reaction is needed to give a good presentation of the thermal decomposition of CCA treated wood [ 171. Comparison between the reaction schemes for untreated and CCA treated wood, each having its own kinetic constants, reveals the following observations. The first peak for CCA treated wood is characterised by lower activation energy, pre-exponential factor and peak temperature, resulting in a slightly higher peak with approximately the same shape but shifted towards lower temperatures. The second peak for CCA treated wood, on the other hand, is characterised by a higher activation energy and pre-exponential factor, resulting in a narrower peak located at lower temperatures. The third peak is better represented by a subsequent reaction, rather than a third independent reaction. The kinetic parameters obtained for this last peak have very low values, resulting in an ill-defined peak. 1425
Based on all previous experiments and the kinetic scheme derived for the pyrolysis of CCA treated wood, the following scenario with respect to the influence of the metals on the pyrolysis behaviour is suggested: (1) Originally, the metals Cu, Cr and As are preferentially bound to the cellulose and lignin compounds in the CCA treated wood, which means that no metals are bound to the hemicellulose compound [9]. (2) The thermal decomposition of hemicellulose (first DTG peak) is shifted to lower temperatures in the presence of CCA. The metals may have a catalytic effect on the thermal decomposition of hemicellulose, resulting in a lower activation energy. (3) The cellulose compound is originally loaded with CrAs04 precipitates and small amounts of other Cu and Cr precipitates or complexes [9], which may hinder the thermal decomposition at first, resulting in a high activation energy for the second DTG peak. However, once the decomposition of cellulose starts, it continues very fast (very high value for pre-exponential factor), resulting in a very narrow peak. As temperature increases, the CrAs04 precipitate decomposes into Crz03 and As205, which further dissociates in As203 and 0 2 at 327.2 "C. The other metal (copper and chromium) compounds, precipitated on or complexed with cellulose, are not volatile. The DTG peak temperature for the second independent reaction (which may be attributed to cellulose decomposition) is 327.9 "C, which is remarkably close to the dissociation temperature of AszO5. There may be a correlation between the decomposition of both compounds. An explanation could be that once the major part of the CrAs04 precipitate on cellulose has decomposed, the decomposition of cellulose continues very fast and may be catalysed by the metals that remained in the solid matrix after the decomposition of CrAs04 or accelerated by the oxidation by 0 2 that comes from the dissociation of AszOs, giving rise to the lower DTG peak temperature compared to untreated wood (371 "C). (4) The intermediate products formed by the thermal decomposition of hemicellulose and cellulose may decompose further at higher temperatures, together with lignin. The small values of pre-exponential factor and activation energy give rise to a very wide DTG peak and thus a very slow charring process. Lignin was originally complexed with CuCr04 and CrAs04 and small amounts of other Cu and Cr compounds [9], delaying the decomposition process. (5) At the end of the pyrolysis process, and maybe even earlier, the metals have been combined to agglomerates and are no more accessible for catalysis, which may be another reason for the slow charring process at higher temperatures.
ARSENIC RELEASE DURING THE LO W-TEMPERATURE PYROLYSIS OF CCA TREATED WOOD WASTE Arsenic may be assumed to be released according to its own kinetic scheme, independent of the release of other volatile compounds. A simple first order reaction is used to model the arsenic release. The derivation of this kinetic scheme is based on the arsenic releases that were experimentally determined for different combinations of temperature and residence time during the pyrolysis studies (labscale and TGA) [21, 221. According to the following equations, with Q the mass of As in the pyrolysis residue at time k, and m at time t (k, is defined as the time that the reactor temperature is 20 "C lower than the pyrolysis temperature, such that i n the period [k,,t] the pyrolysis process can be classified as nearly isothermal):
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imtliennal
=
K(T)(t-to)
K(T) = A e x ( g )
linear regression of -In(m/mo) as a function of (t-b) for a constant temperature T delivers the reaction rate constant K(T) at that particular temperature, and linear regression of In(K) as a function of 1/T for the different temperatures allows one to calculate the pre-exponential factor A and the activation energy E from intercept In(A) and slope (-Em).The resulting pre-exponential factor and activation energy are: A = 0.3875 mid' = 6.4586 E = 20.37 kJ/mol.
s-'
To reversely check the kinetic model, the integral rate equation (for non-isothermal conditions) describing the As release during pyrolysis of CCA treated wood is used in combination with the measured temperature profiles T(t) in order to calculate the corresponding As content of the pyrolysis residues. The calculated arsenic content of the pyrolysis residues is compared with the experimental values (labscale and TGA experiments) in the parity plot, presented in Figure 5 1201
201
P
20
40
60
(As content residue)calc(wt 8 )
so
1
Fig. 5 Parity plot: experimental As contents of the pyrolysis residues (0for labscale, * for TG experiments) are compared with the calculated values, using the first order single reaction kinetic scheme. The errorbars represent 95 % confidence intervals.
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Figure 5 shows that the agreement between calculated and experimental values is fairly good, which means that this simple first order single reaction kinetic scheme succeeds in describing the release of As during the low-temperature pyrolysis of CCA treated wood. It should be noted that the experimental As contents are higher than the calculated values for all TGA results, while for most of the labscale experiments the opposite is true. This may be explained by the fact that the labscale experiments do not consider one particle at one particular temperature, but a reactor filled with particles at a mean pyrolysis temperature. Consequently, the collection of these labscale results combined with the TGA results, obtained in the kinetically controlled regime, results in the description of an "averaged" As behaviour. Furthermore, the treated wood samples used in the TG experiments have a higher CCA content compared to the samples used in the labscale experiments. TG analysis performed on CCA treated wood samples with different CCA content [ 161 shows that the relative concentration of metals (Cr, Cu and As) in the pyrolysis residue increases with the CCA concentration of the original sample. Consequently, the relative amount of metals that is found in the pyrolysis residues resulting from the TG experiments should be higher, which is in agreement with the experimental observations. The kinetic scheme derived here does not account for the influence of the original CCA content. The kinetic constants (A, E) fall within the range of chemically meaningful values. Rate constants of the same order of magnitude were found for the reduction reaction of Cr(V1) to Cr(II1) by Pizzi [6, 91, who studied the fixation of CCA in wood. Because of this similarity the reaction responsible for the As release may be identified as a reduction reaction. The presence of As(II1) in the pyrolysis residue (see speciation study), while As is originally present as As(V) in CCA treated wood (as CrAsOd, confirms this hypothesis. As(V) is thus reduced to As(III), which is more mobile and more toxic, during the pyrolysis process and the resulting As(II1) is probably released as As4O6, which is a volatile compound. This scheme agrees with the thermal decomposition behaviour of CrAs04 (see TGA). The volatilisation reaction is not instantaneous, since As(II1) was found in the pyrolysis residue. Consequently, the release of As during the pyrolysis of CCA treated wood may be better described by two consecutive reactions (reduction followed by volatilisation) instead of one single reaction. The examination of this hypothesis may be a subject for further research. The good agreement between experimental and calculated values (see Figure 5), using the simple scheme derived here, already shows that the single reaction scheme with the corresponding rate constants is of high value as a first approximate to describe the As release during pyrolysis. It should be noted that the validity of this kinetic scheme is only tested for temperatures in the range from 350 "C to 450 "C, which means that good predictions are only guaranteed when lowtemperature pyrolysis processes at temperatures in this range are simulated. A more advanced model should account for the influence of the initial CCA content and incorporate the volatilisation reaction that follows the reduction reaction. Both extensions need more experimental work: (1) pyrolysis experiments using CCA treated samples with different CCA content and (2) speciation studies of the pyrolysis residues. Furthermore, extrapolations to temperatures lower than 300 "C and/or higher than 450 "C should be compared with experimental results.
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DISCUSSION Originally, to prepare the CCA solution, As and Cr are mixed in the form of As205and Cr03, respectively. They are thus both in the high oxidation state, being As(V) and Cr(V1). During the fixation process Cr(V1) is partly reduced to Cr(II1) in order to fix the As as CrAs04. The decomposition of CrAs04 results in Cr2O3 and As205, which further dissociates into As203 and 02. Both metals appear thus in the lower oxidation state, being As(II1) and Cr(III), after pyrolysis. The resulting chromium(II1) compound (CrzO3) is insoluble in water, acids, alkali and alcohols, which could be the explanation for the fact that Cr is more strongly bound in the pyrolysis residue compared to the CCA treated wood and for the lower mobility. Only at temperatures higher than 1000 "C, the chromium(V1) oxide (Cr03), which has a higher mobility and toxicity, is formed. All experimental observations (from this study and the studies of other researchers [23, 24, 251) confirm the hypothesis that As is released as the volatile arsenic trioxide As203 or As406, which is very difficult to capture and therefore rarely detected. Arsine (AsH3) is not present in the gas exhaust since it decomposes at 300 "C. The only way to release arsenic as As(V) compounds is in the particulate or aerosol phase, as condensed arsenates. These can be captured by scrubbing and filtering. The decomposition of wood components may give rise to methyl groups which can methylate arsine (at temperatures below 300 "C), resulting in volatile organic As(II1) compounds, which are less toxic than arsenic trioxide. According to the kinetic model that describes As release during pyrolysis, As is already released at temperatures lower than 300 "C, since the kinetic parameters A and E give rise to a very broad and low DTG peak. This is in contradiction with the reaction scheme derived for the thermal decomposition of chromium arsenate. According to the latter, As is only released for temperatures higher than 327 "C. The volatilisation of arsenic is preceded by a reduction of As(V) to As(II1). This reasoning suggests that the kinetic scheme for As release during pyrolysis, that was derived based on experimental results in the temperature range of 350 "C to 450 "C, may not give accurate predictions for temperatures outside this range. In the temperature range of 350 "C to 450 "C, the kinetic scheme for As release predicts a little rise in the rate of As volatilisation. This rise in volatilisation rate with temperature may also result from liquid-vapour equilibrium controlled by the temperature dependent vapour pressure of As203 (1). Vapour pressures as high as 100 mmHg at 332.5 "C and 200 mmHg at 370 "C, which give rise to non-negligible volatilisation, are reported for As203 [26]. Due to the continuous outflow of pyrolysis vapours, As203 (8) will be removed once it has been formed. Consequently, a rise in temperature results in a rise in volatilisation rate due to the increasing vapour pressure. According to this hypothesis As(V) is reduced to As(II1) before volatilisation occurs. For temperatures lower than 327 "C As205 is the most stable compound [2.1], but a sharp reduction reaction takes place at 327 "C. Since the melting point of As203 (1) is 313 "C [26], the As203 is formed as a liquid. For temperatures higher than 327 "C, the combination of phase equilibrium between As203 (1) and As203 (g) and the carrying off of As203 (g) by the outflowing pyrolysis vapours, .determines the amount of As203 (1) volatilised. To check the validity of this hypothesis, additional pyrolysis experiments at temperatures lower than 350 "C and different residence times are needed.
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CONCLUSIONS The most important conclusions with respect to metal behaviour during lowtemperature pyrolysis of CCA treated wood waste are summarised in Figure 6.
metal release:
CCA treated
327 "C
temperature
pyrolysis residue: no landfillingwithout pre-treatment(Cuand As) metal agglomerates
Fig. 6 Schematic presentation of the most important conclusions with respect to the metal behaviour during low-temperaturepyrolysis of CCA treated wood waste.
Low-temperature pyrolysis is suggested as a possible technique to dispose CCA treated wood waste. The presence of CCA influences the pyrolysis process significantly: CCA can be identified as a promoter of charcoal production. Pyrolysis temperature and residence time of the wood particles are the most important process parameters with respect to metal release during the pyrolysis of CCA treated wood waste. At the low temperatures considered, Cu and Cr do not pose any problems. Arsenic, on the other hand, is the problematic compound, being released already at temperatures between 300 "C and 400 "C.The mechanism responsible for the release of As is identified as the reduction of As(V) to As(II1). which occurs sharply at 327 "C. Limiting the As release implies thus controlling (avoiding) this reduction reaction, which means that the As has to remain in pentavalent form and thus, the temperature has to be kept low enough. Arsenic is probably released as arsenic trioxide As406, which is very difficult to capture and toxic. The pyrolysis residue (solid product) can not be landfilled as such due to the high leachability of Cu and As. The presence of As(III), which is more toxic than As(V), in the pyrolysis residue makes this leaching behaviour even more problematic. The recombination and agglomeration of metals during the pyrolysis process suggest the possibility of using a dry separation (pneumatic centrifugal separation based on density differences) of the metal-containing char residue, which makes the metal recycling process complete and therefore environmentally friendly.
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ACKNOWLEDGEMENTS The authors would like to thank the Department of Chemical Engineering of the K.U.Leuven for their valuable help in conducting the quantitative analysis of the samples. We are especially grateful to Prof. C. Vandecasteele from the Chemical Department for his comments through the course of the study. Also the assistance of Rudy de Vos in the SEM-EDX analysis at the Department of Metallurgy and Materials Engineering of the K.U.Leuven and of Steven Mullens and prof. Jules Mullens in the TG analysis at the Laboratory of Inorganic and Physical Chemistry of the "Limburgs Universitair Centrum" are acknowledged. Furthermore, we are grateful to Beaumartin S.A. and Mr. J.S. Hery in particular for the financial support, the helpful discussions and the wood samples.
REFERENCES 1. Meier D., Wehlte S . and Faix 0. (1997) Flash pyrolysis - a possibility to utilise contaminated wood. Plant Research and Development, 46,46-55. 2. Wehlte S.,, Meier D., Moltran J. and Faix 0. (1997) The impact of wood preservatives on the flash pyrolysis of biomass. In: Developments in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), pp. 206-219. Chapman & Hall, London. 3. Dahlgren S.E. (1974) Kinetics and mechanism of fixation of Cu-Cr-As wood preservatives IV. Conversion reactions during storage. Holtforschung, 28 (2). 5861. 4. Dahlgren S.E. (1975) Kinetics and mechanism of fixation of Cu-Cr-As wood preservatives VI. The length of the primary precipitation fixation period. Holtforschung, 29 (4), 130-133. 5 . Dahlgren S.E. and Hartford W.H. (1972) Kinetics and mechanism of fixation of Cu-Cr-As wood preservatives I. pH behaviour and general aspects on fixation. Holzforschung, 26 (2), 62-69. 6. Pizzi A. (1981) The chemistry and kinetic behaviour of Cu-Cr-As/B wood preservatives. I. Fixation of chromium on wood. J , Polym. Sci., 19,3093-3121. 7. Pizzi A. (1982) The chemistry and kinetic behaviour of Cu-Cr-As/B wood preservatives. 11. Fixation of the Cu/Cr system on wood. J. Polym. Sci., 20, 707724. 8. Pizzi A. (1982) The chemistry and kinetic behaviour of Cu-Cr-As/B wood preservatives. 111. Fixation of a Cr/As system on wood. J. Polym. Sci., 20,725-738. 9. Pizzi A. (1982) The chemistry and kinetic behaviour of Cu-Cr-As/B wood preservatives. IV. Fixation of CCA to wood. J. Polym. Sci., 20,739-764. 10. Chou C.K, Chandler J.A. and Preston R.D. (1973) Microdistribution of metal elements in wood impregnated with a copper-chrome-arsenic preservative as determined by analytical electron microscopy. Wood Sci. Techno/.,7 , 151-160. 11. Van den Broeck K., Helsen L., Vandecasteele C. and Van den Bulck E. (1997) Determination and Characterisation of Copper, Chromium and Arsenic in CCA treated Wood and its Pyrolysis Residues by Inductively Coupled Plasma Mass Spectrometry. The Analyst, 122,695-700. 12. Helsen L., Van den Bulck E., Van den Broeck K. and Vandecasteele C. (1997) Low-temperature pyrolysis of CCA treated wood waste: chemical determination
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13.
14. 15.
16. 17. 18. 19. 20. 21. 22. 23. 24. 25.
26.
and statistical analysis of metal in- and output; mass balances. Waste Manage., 17 (l), 79-86. Helsen L. and Van den Bulck E. (1998) The Microdistribution of Copper, Chromium and Arsenic in CCA Treated Wood and Its Pyrolysis Residue Using Energy Dispersive X-Ray Analysis in Scanning Electron Microscopy. Holzforschung, 52 (6), 607-614. Helsen L., Van den Bulck E. and Hery J.S. (1998) Total recycling of CCA treated wood waste by low-temperature pyrolysis. Waste Manage., 18,57 1-578. Helsen L. and Van den Bulck E. (1997) Release of metals during the pyrolysis of preservative impregnated wood. In: Developments in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), pp. 220-228. Chapman & Hall, London. Helsen L., Van den Bulck E., Mullens S. and Mullens J. (1999) Low-temperature pyrolysis of CCA treated wood: thermogravimetric analysis. J. Anal. Appl. Pyrolysis, 52,65-86. Helsen L. and Van den Bulck E. (2000) Kinetics of the low-temperature pyrolysis of chromated copper arsenate-treated wood. J. Anal. Appl. Pyrolysis, 53,5 1-79. Ebbinghaus B.B. (1993) Thermodynamics of gas phase chromium species: the chromium oxides, the chromium oxyhydroxides, and volatility calculations in waste incineration processes. Combust. Flame, 93, 119-137. Frandsen F.J., Helble J.J., Erickson T.A. and Kuhnel V. (1997) Prediction of trace element partitioning in utility boilers, presented at the EPA/DOE/EPRI Symposium. Wu C.Y. and Biswas P. (1993) An equilibrium analysis to determine the speciation of metals in an incinerator. Combust. Flame, 93, 31-40. Helsen L. (2000) Low-temperature pyrolysis of CCA treated wood waste. Ph.D. Thesis, Katholieke Universiteit Leuven, Belgium. Helsen L. and Van den Bulck E. (2000) Metal Behaviour During the LowTemperature Pyrolysis of Chromated Copper Arsenate Treated Wood Waste. accepted for publication in Environ. Sci. Technol.. McMahon C.K., Bush P.B. and Woolson E.A. (1986) How much arsenic is released when CCA treated wood is burned. For. Products J., 36,45-50. Hirata T., Inoue M. and Fukui Y. (1993) Pyrolysis and combustion toxicity of wood treated with CCA. Wood Sci. Technol.,27,35-47. Cornfield J.A., Vollam S. and Fardell P. (1993) Recycling and disposal of timber treated with waterborne copper based preservatives. 24th annual meeting, Orlando, USA, I R G N P 93-50008. Perry R.H. and Chilton C.H. (1973) Chemical Engineers' Handbook, 5' edn. McGraw-Hill, New York.
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Pyrolysis of Biomass as Pre-Treatment for Use as Reburn Fuel in Coal-Fired Boilers C. Storm, S. Unterberger, K. R. G. Hein Institut fur Verfahrenstechnik und Dampjkesselwesen (IVD), University of Stuttgurt, Pfuflenwaldring 23, D-70550 Stuttgurt, Germany
ABSTRQCT Biofuels attract increasing interest in power plant technology as sources of carbon dioxide neutral fuels. Besides using solid pulverised biomass as an additional fuel in coal-fired boilers a further possibility to run a combined coal and biomass process is to pre-pyrolyse or pre-gasify biomass and to inject the gas as reburn fuel into the coal-fired boiler. Within the pretreatment process the solid feedstock is separated into a h g h energy gas and a solid residue. The gas can be used as a reburn fuel in the coal-fired boiler with excellent NO, reduction properties. The separation of coal and biofuel ash within this process enables a specialised use of both residuals. At the Institute for Process Engineering and Power Plant Technology the pyrolysis and gasification of different biomass has been investigated with special emphasis on the gas, tar and char composition and on the NO, reduction efficiency of the produced gas. The gasification experiments have been carried out in an electrically heated lab-scale entrained flow reactor in understoichiometric atmosphere, for the reburn experiments a pulverised fuel combustion reactor was available. The gas, tar and composition as well as the NO, reduction efficiency have been analysed at temperatures of the entrained flow reactor between 400°C and 1300°C and of the fluidised bed reactor between 600°C and 900°C and air ratios of the pyrolysislgasification process between 0 and 0,4. Detailed results of the investigations using pyrolysis gas from coal [7, 81, sewage sludge [ 171, or biomass [ 181 for NO, reduction are published elsewhere.
INTRODUCTION Various forms of biofuels represent interesting feedstocks for combustion, gasification and pyrolysis processes in power plant technology as a source of C02 neutral fuels. Biomass has an estimated potential in Germany of about 660 PJ (1 Petajoule = 10'' Joule). This is equivalent to about 5 % of the annual primary energy consumption in Germany. [101 The aim of the investigations is the optimisation of a co-combustion process using biofuel pyrolysis/gasification gas as a reburn fuel with special view on 1433
parameters for NO, reduction, parameters for alkali, chlorine and heavy metal release. The advantage of ash separation within this combined combustion of coal and biomass offers a separate ash utilisation. While the ash utilisation of pure coal ash is technically approved and economically feasible already, different kmds of utilisation of biomass ash have to be proved. The utilisation of mixed ashes in building industry, which is the main market for ashes from coal combustion, is not foreseen in the legislation of the European Union. The investigated pre-treatment process is a suitable alternative for biomass co-combustion. FUNDAMENTALS
BIOMASS Biomass contains more linear and weaker connected macromolecules than the highly cross-linked coal with its aromatic components. Photosynthesis is the basic step in the formation of biomass. An endothermic reaction creates glucose from carbon dioxide and water. In a second step, the main biomass components cellulose, hemi-cellulose and lignin are created. The composition of cellulose (% wt.) is 44,4 % carbon, 6,2 % hydrogen and 49,4% oxygen. Due to the chemical composition it is evident, that biomass is an oxygen rich feedstock. Hemi-cellulose are like cellulose chains from glucose molecules, but their length is much shorter and they can have subchains. Unlike cellulose and hemi-cellulose li& consists of aromatic elements, which are arranged three-dimensional. The share of oxygen in lignin is much below the oxygen content of celluloses.
NITROGEN, CHLORINE, A L U I N E S , HEA W METALS Inorganic components differ between the different types of biofuels. For energetic use of biofuels the most interesting inorganic elements are chlorine and alkalis as they cause problems in conventional combustion systems. Nitrogen is another element, whch is important when burning biofuels, because it can either cause NO, emissions or it can help minimising them, when the nitrogen is bound in certain compounds i.e. like HCN or ammonia. Furthermore heavy metals and other trace elements are of interest as they can cause emission problems. Those elements usually occur only in anthropogenic made residues like sewage sludge, waste wood etc.
PYROLYSIS OF BIOFUELS Pyrolysis (or devolatilisation) is a thermal decomposition of organic matter in an inert atmosphere. In the gasification or partial gasification oxidation medium is added to the process. This oxidation medium usually is air or oxygen. Pyrolysis is also the first step in combustion and gasification. Besides solid residues carbon-, hydrogen-, and oxygen containing gases as well as condensable aromatic hydrocarbons, which are designated as tars, are formed during pyrolysis. [4] Homogeneous and heterogeneous reactions in the pyrolysis process cause a changing product composition, depending on the reaction temperature. The main reactions are shown in Figure I .
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Temperature [“C]
Figure 1 Equilibrium of different gasification reactions depending on the reaction temperature NO, REDUCTION WITH PYROLYSIS GAS Nitrogen oxides (NO,) are formed during combustion at high temperatures by the oxidation of air nitrogen and of nitrogen bound in the fuel [l, 5 , 6, 201. The sustainingly increasing demand of energy, which is mainly met by combustion of fossil fuels, makes the use of technologies for the minimising of NO, emissions essential. NO, Formation There are hown three different kinds of NO, formation in coal combustion: thermal NO, formation; prompt NO, formation; NO, formation from fuel nitrogen. All three h d s of formation result from different reaction zones in the flame and from two principle sources of nitrogen, the fuel nitrogen and nitrogen from air. NO, formation from fuel nitrogen plays the most important role for NO, emissions from coal combustion [3, 141. After the primary pyrolysis of the coal the nitrogen occurs in the solid phase bound into the char and in the gaseous phase. A small part is directly converted to NO, released as N2 or reduced to N2. From the fuel nitrogen in the gaseous phase mainly ammonia and HCN are formed. The char nitrogen is either oxidised heterogeneously or decomposed to N2 or further gaseous components llke HCN or radicals are formed. NO, Reduction NO, emissions can either be influenced by primary measures (less NO, formation) or by secondary measures (reduction of NO,) [2, 3, 7, 11, 191. The chemical reactions of
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NO, deformation during secondary measures are much more complex than those of the NO, formation. Not all ways of reactions are known by now. The reactions are differentiated into two types, the homogeneous and the heterogeneous, which characterise the reactions in the gas phase and the reactions between the gas phase and the parhcle surfaces. Homogeneous reduction: Nitrogen oxides can be reduced during the combustion process either by intermediate products or by added reducing agents. Reducing agents are divided into selective and non-selective agents. Non-selective reducing agents, like carbon monoxide, hydrogen, and hydrocarbons, reduce by catching the oxygen, which is still in the flue gas due to reasons of the combustion technique. This results in a high demand of reducing agent. In contrast to the non-selective agents are the selective reducing agents, which only reduce nitrogen oxides independent on the oxygen content in the flue gas. Both reaction mechanisms are supported exclusively by radicals. An increase of temperature leads also to an increase of radicals. There is also the possibility to influence the formation of radicals by decreasing the O2partial pressure. Heterogeneous reduction: The heterogeneous reduction of NO directly at the surface of char particles is extremely dependant on the temperature [15]. At low temperatures no reduction takes place at the particle surfaces in contrast to high temperatures, where the rate of reduction is about 95 %. This behaviour is due to the much faster oxidising reaction of nitrogen than the reduction of NO. The heterogeneous reduction mainly takes place in fluidised bed and grate ftring systems. In pulverised fuel combustion systems the influence of this reduction mechanism is much lower, because a high particle concentration and a high share of combustible solids is not given.
Reburning - In-Furnace NO, Reduction Methods of minimising NO, emissions already in the furnace by staged burner or air staging minimise the often very complex reduction systems after the furnace like DeNOx systems, injection of ammonia, which are often used for NO, reduction in flue gases. Air staging is already a technically feasible option of minimising NO, emissions and it is already integrated in many lignite combustion systems [9] and hard coal combustion systems [12]. Reburning, also called in-furnace reduction or hel-staging, is a three step combustion. In the first step the main fuel (usually coal) is burnt completely at an air ratio above one, within the second step the reburn fuel is injected into the furnace without adding any oxygen. In this reduction zone the air ratio is below one and creates conditions to reduce the NO,, which has been formed in the main combustion zone. To ensure a complete burnout air is injected into the third zone. The potential of NO, reduction by reburning depends on the air ratio in the reduction zone, the temperature in the reduction zone, the mixing conditions of the reburn fuel in the reduction zone, the residence time in the reduction zone, and the fuel and reburn he1 composition. Figure 2 shows the process of the thermal pre-treatment of biomass fuels in combination with the injection of biomass derived gas into a pulverised fuel combustion as reburn fuel.
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Figure 2 Process scheme: thermal pre-treatment of biomass and rebuming of the biomass derived gas in a coal-ftred boiler
The thermal pre-treatment process consists of a pyrolysis or gasification process, where the biomass is fed into. After the pre-treatment process the solid pyrolysis residues are separated from the energy-rich pyrolysis gas within a hot gas filtration unit. The gas is then used as reburn fuel.
TEST FACILITY The BTS (BrennstofRrennstuhng - fuel splitting and staging) test facility was built within the scope of the project ‘Combined minimizing of NO,-production and reduction of formed NO, during combustion of coal dust’ in co-operation with a German coal mining utility [16]. Figure 3 shows the scheme of the combined test facility. The test facility consists of following two main components:
0
entrained flow reactor with fuel feeding unit, carrier gas pre-heating, product gas analysis and hot gas filtration with precipitator, cyclone and ceramic candle filter. combustion reactor with fuel feeding unit, air pre-heating, pyrolysis gas and burnout air supply, gas analysis and flue gas filtration.
The two facilities can be operated either together for i.e. investigation of NO, reduction by the means of pyrolysis gas as reburn fuel or separate for investigations concerning pyrolysis, gasification and combustion of various solid fuels.
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pre-heater
entrained flow gas and tar analysls
cyclone
hot gas filtration
-
residues -injection air feeding unit primary air secondary air
Figure 3 Test facility for the research of pyrolysis and reburning The entrained flow pyrolysis reactor (20 k W d can be connected to the entrained flow combustion reactor (50 kW,). The pyrolysis residues are separated fiom the pyrolysis gas in the cyclone and in the hot gas filter. Afterwards this gas is led as reburn fuel into the reduction zone of the combustion reactor. All pipes are trace heated up to 350 OC to avoid tar condensing. The results of pyrolysis of different solid fbels discussed here were performed at the entrained flow reactor. The results of NO, reduction experimentshave been gained using the pyrolysis gas as reburn fuel in the pulverised fuel combustion reactor. E N T m E D FLOWREACTOR The dosing of the solid fuel into the entrained flow reactor takes place gravimetrically. The dosing rate of the fuel is usually between 0,5 and 2,5 kg/h. The reactor itself consists of an electrically heated ceramic pipe of the material ,pythagoras" (ceramic with 60 % A1203 heatable up to 1300 "C)with a total length of 2,7 m, an outer diameter of 100 mm, an inner diameter of 80 mm and a heated length of about 2 m. At the outlet of the reaction pipe the hot gas filtration unit with the connection for gas analysis is located. A heated pipe leads the gases fiom the hot gas filtration to the combustionreactor or to the flare, respectively.
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COMBUSTION REACTOR (PUL VERLSED FUEL) The pulverised fuel combustion reactor was used for reburn test runs using pyrolysis gas generated in the entrained flow reactor as reburn fuel. The main part of the reactor is the electrically heated furnace with a total electric output of 57 kWe,, devided into five heating zones. A ceramic pipe with a length of 2,5 m and an inner diameter of 200 mm is installed inside the furnace as reaction pipe. The feeding unit, the reactor itself and the flue gas filter are designed for a coal feeding rate between 1 and 5 kgih. The fuel and the combustion air, split in primary and secondary air, are injected through the movable block burner. The maximum wall temperature of the combustion reactor is 1500 "C.At a position of 0,9 m distance below the burner of the combustion reactor the pyrolysis gas is injected through a nozzle. To ensure a complete burnout air can be injected through a height adjustable burnout air probe. The reburn test runs have been performed under constant conditions in the combustion reactor. Besides the wall temperature (1300 "C)the fuel, the fuel feeding rate, and the probe positions stayed unchanged. The hard coal GBttelbom was burnt in the main combustion zone at an air ratio of 1,15. The fuel feeding rate amounts to 1 kgih.The position of the burnout air probe remained unchanged. It was adjusted to a calculated residence time of 1 q,'/h pyrolysis gas of 2 s in the reduction zone. Those settings (residence time, wall temperature) have been proved to be optimal for reburning in earlier projects [7,81. Only the burnout air was varied in order to get a constant O2concentration of 3 % in the flue gas.
FUEL ANALYSIS Table 1 shows the composition of the fuels used. Besides natural like biomass also anthropogenic treated biofuels like sawdust and sewage sludge have been used for experiments. The wheat straw and beech have been milled for the experiments in a cutting mill with sieve sizes of 1,5 mm, 1 mm and 0,5 mm, respectively. The sawdust or sanding dust comes from a chip board manufacturer and was delivered with the used particle size. The sawdust contains melamine from glue constituents. Melamine is nitrogen rich compared to the nature like biomass. This kind of biomass was chosen for NO, reduction experiments, because nitrogen plays a major role for the reburn efficiency of the pyrolysis gas. Several sewage sludges have been pyrolysed and the pyrolysis gases have been tested for their rebum and NO, reduction efficiency [171. The hard coal GBttelbom was used for combustion in the pulverised fuel combustion reactor as main fuel.
RESULTS PYROLYSIS GAS, TAR AND CHAR Main components of the pyrolysis gas are the gaseous components CO, C02, HZ,and aliphatic hydrocarbons as well as the condensable aromatic hydrocarbons, usually designated as tars.
1439
Table 1 Compositionof the rawfiels straw
fuel
beech
saw-
sew. sew. sludge sludge DD HG
hard coal GB
abbr.
ST
BU
dust SM
fixedC volatile ash
17,O
77,8 5,2
15,4 84,O 0,6
21,2 773 1,3
5,6 543 39,9
4,7 46,9 48,4
57,O 33,9 991
49,8 6,3
493 6,3 5,l 0,l
56,s 7,2 5,4 1,s
81,4 536
0,1
53,O 7,3 7,7 2,8 0,2
0,1
095
29,l 5,7
29,3 5,l
10,6 593
C H N
S Cl O(diff) mois%
49,7 6,l 0,s 0,2 1,l 42,4 8,3
0,l
*
43,7 5,9
*
39,O 5,3
195
130
60 pm 16473
*
36,6 0,8
0,6 10,2 2,s 36,7 2,3
*
*
13694
11287 30723
20,O 7,O 1,4 26,l 9,2 23,s
39,8 11,7 14,2 6,9
41,6 11,3 9,1 8,6
1,8
*
*
42,3 24,3 11,3
65
1 9 1
191
0,8
0,8
421 393
0,6
0,s
027
*
*
66
notdetermined Basic investigations on the pyrolysis of different biofuels under different process and fuel parameters have been carried out in previous investigations. Especially gaseous and condensable pyrolysis products have been the focus of those investigations [13, 171. For the sake of completeness main results of straw pyrolysis are presented here. Straw has been chosen as reference fuel. As shown in the previous investigations other biofuels behave similar regarding the main components. Even sewage sludge in spite of its high ash content produces almost the same gaseous and tar compopents like straw compared on a water and ash free basis (daf). Gaseous Components
Figure 4 shows the yield of the main components of straw pyrolysis. The composition of the pyrolysis gas changes with increasing pyrolysis temperature to the favour of
1440
light gases like H2 and CO. Carbon monoxide and hydrogen are generated in a higher share with increasing reaction temperatures. Carbon monoxide production starts already at 400 OC; hydrogen is significantly generated at temperatures higher than 800 "C. Aliphatic hydrocarbons are also produced at low temperatures and reach a maximum at about 800OC. The aliphatic hydrocarbons are decomposed to less complex compounds with rising temperature. Carbon dioxide is not strongly influenced by the reaction temperature. The high oxygen content in the raw fuel leads to an increased formation of CO and C02. I
'
I
.
1
'
I
.
I
*
I
.
I
V
I
.
1
.
I
straw, entrained flow pyrolysis main gaseous components
-
-
----CO,
-
0,l
0,o
1
400
.
1
500
.
1
600
700
800
900 1000 1100 1200 1300
pyrolysis temperature ["C] Figure 4 Main gaseous components of straw pyrolysis gas Tar Components The tar species are divided into two groups. The term heavy tars describes the s u m of all species with a boiling point above about 200°C. These are all aromatic hydrocarbons with more than two benzene rings (naphthalene and above). The light tars are mainly benzene, toluene and xylene. The boiling point of the compounds is below 200°C and the number of benzene rings does not exceed two. Figure 5 gives information about the tar production of straw. At low temperatures, more heavy components are produced. Light tars are formed by cracking out of heavy components with higher temperatures and reach a maximum between 800°C and 9OOOC comparable to the aliphatic hydrocarbons. With further increasing temperatures, both tars, heavy and light, are decomposed rapidly to less complex compounds like the light gases H2 and CO.
1441
220 200
1
’
I
.
*
I
I
.
I
.
-
\ 3180 il60
-3 v)
e
1
.
I
’
I
’
I
’
!
1 barn, Ah h=o, 21 -.total tars
\
140 120 -
I
straw, entrained Row pyrolysis condensable aromatic componds-
\
-
fraction (mainly BTX) - - light heavy fraction (> 1 ring) -
\ \
-
\
\
100-
80:
-
60;
-
20 -
40
0
l
‘
l
.
g
’
l
*
l
*
’
*
I
.
I
’
I
Figure 5 Production of tars during straw pyrolysis in dependence on the pyrolysis temperature Pyroljsis Residues
Char samples were taken from the pyrolysis experiments at the entrained flow reactor. Figure 6 shows the devolatilisation and the composition of the char from straw pyrolysis at different temperatures and air ratios and of the raw straw. The results of the thermogravimetric measurements are related to the ash content of the raw fuel. Straw shows a high devolatilisation already at low temperatures. The devolatilisation rate is almost not dependant on the air ratio in the entrained flow reactor. Deviations between the values are within the tolerance of the sampling and measurement system. Other biofuels (beech, sewage sludge) behave similarly to straw during the pyrolysis. The release of nitrogen during the pyrolysis is nearly independent from the pyrolysis temperature and from the fuel. Around 20 % of the nitrogen stays in the char independent of the fuel. Even the share of nitrogen release during pyrolysis of a nitrogen rich fuel like sewage sludge lays in the same order as the nitrogen release out of N poor fuels, like straw and beech. The independency of the temperature on the nitrogen release is only true for temperatures above 600 “C. Due. to the low degasification rate of straw pyrolysis at low temperatures there is only a minor nitrogen release below 600 “C.
1442
100
straw, entrained flow pyrolysis proximate analysis of pyrolysis residues in dependance on temperature I air ratio Z1 recipitator sample devolatilised A: h=O volatiles B: h=0,2 EZl fixedcaibon c: h=0,4 ash
90
-
80
d
.70
0
$ 60
rh
8 50
-2 E
40
{ 30 20 10 n
raw
I
B C 400
I
B C 600
A B C 800
A B I000
pyrolysis temperature ["C]
1200
Figure 6 Proximate analysis of straw (raw material and residual char), values related to ash content of raw fuel
NO, REDUCTION WITH PYROL YSIS GASES Using pyrolysis gas from biofuels as reburn fuel with good NO, reduction efficiency enables a way of contributing to the net COz reduction. It opens up a way of cocombustion without affecting the coal ash and causing operational problems in the boiler, and with concurrent minimising of emissions. The results are usually outlined in the diagrams as NO, emissions of the combustion reactor related to 6 % O2 in the flue gas over the fuel mass flow rate of the pyrolysis process. The higher the fuel mass flow rate of the pyrolysis process the more pyrolysis gas and thus reburn fuel is injected into the combustion. Directly related to the flow rate of the pyrolysis gas into the combustion is the air ratio in the reduction zone of the combustion reactor. The outline of the NO, emissions over the air ratio describes the reburn efficiency of the injected gas, whereas the outline over the fuel mass flow rate gives information on the biomass ratio on the overall process. Influence of Pyrolysis Temperature
Figure 7 shows the NO, emissions using pyrolysis gas from straw in dependance on the straw mass flow and the pyrolysis temperature. The pyrolysis gases were generated in the entrained flow reactor with the ceramic reaction pipe. With increasing straw mass flow into the pyrolysis process and thus with increasing pyrolysis gas flow into the reduction zone of the combustion reactor the NO, emissions of the combustion decrease until a minimum is reached at a certain straw mass flow rate. A distinct influence of the pyrolysis temperature on the reburn efficiency can be determined. Pyrolysis gas generated at 800 "C showed the best NO, reduction efficiency with a NO, minimum of 110 mg/m3 at a flow rate of 1,3 kg@. Gases generated at lower or higher pyrolysis temperature resulted in worse NO, emissions. More straw must be
1443
pyrolysed to reach NO, emissions below 200 mg/m3. When increasing the pyrolysis gas flow after reaching the NO, emission minimum the NO, emissions rise again due to an excessive supply of nitrogen and combustible compounds in the pyrolysis gas. Those compounds lead to higher NO formation in the burnout zone of the reactor. 1400 1200
I
I
I
I
pyrolysis gas as reburn fuel combustionfuel: GOttelborn (1 kg/h) Tm= 1300 "C residence time reduction zone: 2 s .
-
straw, entrained flow pyrolysis &=O, 22, pyrolysis temperature: -C5OO0C -0- 600°C -X- 800 "C - - L '1000"c -0-1200*c --V--I3OO0C
200
0
-
-
.
-
.
-
.
-
I
I
I
I
Figure 7 NO, emissions usingpyrolysis gaspom straw in dependance on straw mass flow and pyrolysis temperature The optimum settings for the pyrolysis process regarding the NO, reduction efficiency is 800 "C pyrolysis temperature and a straw mass flow of about 1,3 kg&/h. This is a mass ratio for the complete co-combustion process of 1,3 straw to coal mass flow. This is equivalent with a thermal power input of 21415 kJ/h straw or an overall share on the total thermal power input of about 40 % of straw. Outlining the air ratio in the reduction zone over the fuel mass flow rate of the pyrolysis process for different pyrolysis temperature (Figure 8) the air requirement of the pyrolysis gases is shown. There is a linear relation between the fuel mass flow and the air ratio. There is no big difference in air consumption between the gases generated at different pyrolysis temperatures. That means that the composition of the different gases is nearly the same regarding the oxygen demand. Oxygen compounds like CO and COz, which lower the oxygen demand of a gas, exist in all gases. The low air demand of the pyrolysis gas generated at 500 "C is due to the insufficient devolatilisation of straw in the pyrolysis process. The data of Figure 7 are outlined related to the air ratio in the reduction zone in Figure 9. As the air demand of all pyrolysis gases are nearly the same influence of the gases on the NO, emissions can be seen in Figure 9. Taking the information of the two figures the result is a difference between reburn efficiency and NO, reduction efficiency. Reburn efficiency means in that case the capability of producing a certain air ratio in the reduction zone. This is directly related with the air demand of the gas mixture. The reduction efficiency takes into consideration additionally the capability of reducing NO,. Therefore a sufficient amount of radicals have to be present. 1444
-
pyrolysis gas as reburn fuel combustion fuel: GLlttelborn (1 kglh) .
-
-
.-
'
-
entrained flow pyrolysis .-o 0,80- straw, -0 22, pyrolysis temperature:
!0,75 k-*5 0600°C 0 - & - 1000 "c -0-1200°C Ov7O -X800 "C --V--1300°C
. I -
.m
oc
0,651
0,o
-
I
I
I
0,s
1,o
1,s
I
2,o
2,s
fuel mass flow (straw) [kg,Jh] Figure 8 Relation between the air ratio in the reduction zone and fuel mass flow rate of the straw pyrolysis process 1400
I
-
I
-
-
I
I
I
I
'
I
'
I
I
'
straw, entrained flow pyrolysis =O 22, pyrolysis temperature: k 5 0 0 "C * 1000 "C
1200 -
-0-
-x-
600°C 800 "C
-
-0-1200°C . --V--1300 "C -
pyrolysis gas as reburn fuel combustion fuel: Gottelborn (1 kglh) reduction zone: 2 s
-
200 0
. -
'
l
.
"
l
'
l
"
'
l
'
l
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l
r
l
~
The air demand of pyrolysis gases produced at €ugh temperatures comes from light gases ldce CO and H2, which catch the oxygen very fast, but do not form radicals. The NO, reduction efficiency of a gas generated at 800 "C comes from the radicals, which are produced by cracking hydrocarbons (aliphatic or aromatic). The air demand 1445
of the gas is equal to the 'lighter' gas, but the oxygen consumption cames radicals with it. The content of hydrocarbons in the gas have their maximum in a pyrolysis gas generated at 800 "C. Because all investigated fuels have similar hydrocarbon maxima at around 800 "C in further comparisons of influences on the reduction the evaluation focuses on those settings. Influence of Air Ratio in the pVrolysis Process Having a technical feasible process in mind an addition of air to the pyrolysis process is practically. Experiments with air ratios of 0,lS and 0,3 in the entrained flow reactor have been performed with several fuels. Results gained with all fuels are similar therefore only straw results are presented. In Figure I0 a distinct influence of the air ratio in the pyrolysis process on the reburn efficiency is determinable. The NO, emission minimum of below 200 mg/m3 are achieved with a very high mass flow of straw. For the experiment with an air ratio of 0,3 the minimum could not be reached due to the limited feeding rate of the reactor. straw, entrained flow pyrdysis 22, pyrolysis temperature 800 "C
0'
0,o
I
1
03
I
1,o
I
13 2,o fuel mass flow (straw) [kg,Jh]
I 23
Figure 10 NO, emissions using pyrolysis gas from straw in dependance on straw mass flow and air ratio in the pyrolysis The air demand of the gas shows a dependance on the air ratio in the pyrolysis reactor (Figure 11) . Because the addition of air into the pyrolysis process already serves a part of the demand the air consumption in the combustion process is less.
1446
pyrolysis gas as reburn fuel combustion fuel: GOttelborn (1 kg/h) .
-
-
085
0,o
I
I
I
OS
1,o
13
I
23
2-0
fuel mass flow (straw) [kg,Jh] Figure 11 Relation between the air ratio in the reduction zone and fuel mass flow rate of the straw pyrolysis process (h test runs)
The reduction efficiency of the gases is not as much influenced by the air ratio in the pyrolysis process like when varying the temperature (Figure 12). The minimum NO, emission is shifted to lower air ratios in the reduction zone with increasing air ratio in the pyrolysis. Combining the information from Figure 10 to Figure 12 the result is, that during the pyrolysis process mainly gas components are oxidised, which do not form the necessary radicals. 1400
.
I
.
I
-
I
I
-
I
.
I
.
I
.
I
9
I
-
straw, entrained flow pyrolysis 22,pyrolysis temperature 800 "C
combustion fuel: G6ttelbom (1 kglh)
200 -
Y,15
1,lO
1,05
1,OO
0,95 0,W 0,85 0,80 0,75 0,70 0,65
air ratio in the reduction zone [-I Figure 12 NO, emissions using pyrolysis gas fiom straw in dependance on air ratio in the reduction zone and air ratio in the pyrolysis process 1447
Influence of Fuel Several fuels have been tested on their NO, reduction efficiency when pyrolysing them under different conditions. For further comparison the setting of 700°C pyrolysis temperature and inert atmosphere in the entrained flow reactor has been chosen. Detailed results of experiments with coal [7, 81, biomass [18] and sewage sludge [17] pyrolysis gas have been published elsewhere. The results are integrated into Figure 13 for comparison. The used synthetic gas mixture (26,5 %CO, 5,35 % COz, 7,75 %H2, 53,l % CH4, 7,3 % C2Hx)characterises the composition of a pyrolysis gas from coal generated at a reactor temperature of 700 "C without any N and tar compounds. The NO, minimum of 200 mg/m3 is reached at an air ratio of 0,75 in the reduction zone. Changes of the composition of the gas mixture (different CO, C02, H2, CH4 contents) did not have any effect on the NO, emissions and on the air ratio. Even 100 % of CH4 results in same NO, emissions like reburning with synthetic gases. The better NO, reduction behaviour of pyrolysis gas compared with the synthetic gas mixtures examined is due to the nitrogen compounds in the raw fuel and thus in the pyrolysis gas and due to the higher hydrocarbon content (especially aromatic hydrocarbons, tars). This can be seen from the results achieved with several reburn experiments with pyrolysis gas from coal and different biomass generated in the entrained flow reactor shown in Figure 13. The coal pyrolysis gas corresponds with the synthetic gas. But additional tar is contained in the real pyrolysis gas. A NO, emission minimum of 200 mg/m3 is already achieved at an air ratio of 0,9 in the reduction zone. Biomass with a rather low nitrogen content in the raw fuel (straw, beech) and biomass with a very high N content (sawdust, sewage sludge) have been chosen to show the occurring effects more clearly. The sawdust is a residue from wood-processing industry and it is contaminated with melamine (C&&), which is a constituent of the adhesives used. 14009'
.,-'??
I
'
I
pyrolysis fuel I .gas'as rebum . I ' I " , _ combustionfuel: GLIttelbom (1 kg/h) T =13OO'C re3enc.e time reduction zone: 2 s entrained flow pyrolysis &=O, 22/23,pyrolysis temperature 700 "C-47-sewage sludge .
'
-
-X-
200
beech
-
0 1,15
"
~ 1,lO
"
'
1,05
1.00
" 0,95
'
'
~
0,90 0.85
'
~
'
~
'
~
'
0,80 0,75 0.70 0,65
air ratio in the reduction zone [-] Figure 13 NOx emissions using pyrolysis gas from different fuels in dependence on air ratio in the reduction zone and pyrolysis fuel
1448
*
Straw and beech produce comparable gases because of their similar composition. The minimum NO, emission under same combustion conditions of about 200 mg/m3 can be found at air ratios between 0,84 and 0,87 in the reduction zone. Unlike the pyrolysis gases fiom beech the generated gases from sawdust and sewage sludge are very rich concerning nitrogen compounds. Due to the high N content in the pyrolysis gas and tars the distinct minimum of NO, emissions (about 160 mg/m3, 110 mg/m3, respectively) is shifted to higher air ratios (A = 0,92 and h = 0,95) in the reduction zone as shown in Figure 13. The increasing NO, emissions at low air ratios is caused by unreacted N-species of pyrolysis gas and tars, which are oxidised in the burnout zone of the combustion reactor. CONCLUSION
The results of the reburning test runs have shown that pyrolysis gases fiom biofuels are very well suitable for using them as a rebum fuel with excellent NO, reduction quality. Trials with pyrolysis gas of different biomass types resulted in NO, emissions of a coal furnace below 200 mg/m3. Especially with nitrogen rich pyrolysis gases low emissions were achieved. Nitrogen compounds bound in pyrolysis tars and gaseous compounds, like ammonia and hydrogen cyanid, have a big influence on the reduction capacity of gasificatiodpyrolysis gas, which has been proven by using a gas fiom a high melamine contaminated sawdust as feedstock for the entrained flow reactor. The reburn efficiency increases with a higher share of fuel nitrogen. But besides nitrogen compounds also radical forming hydrocarbons are important for NO, reduction in the reduction zone of the combustion reactor. Therefore pyrolysis gases generated at 800 "C show the best reduction efficiency due to their high content of aliphatic and aromatic hydrocarbons. Excellent NO, reduction can be achieved at pyrolysis temperatures around 800 "Cwith an overall straw to coal ratio of 1,3. This means that the share on the t h e m 1 input of 40 % is met by straw. Investigations on the behaviour of volatile and ash components have been carried out. The solid pyrolysis residues of several fbels have been analysed for their main ash composition and their devolatilisation. Increasing reaction temperatures result in a higher devolatilisation for all fuels. E.g. straw shows a devolatilisation of approx. 80 %dry above 600 "C pyrolysis temperature. But already at low pyrolysis temperatures of 400 O C a rather high devolatilisation of about 70 % can be reached. Other biofuel show a similar behaviour on ash and water free basis. Focusing the devolatilisation efficiency high temperatures in the pyrolysis process give best results, but also pyrolysis at low temperatures yields in sufficient devolatilisationrates. Gaseous hydrocarbons have a production maximum at about 800 OC reaction temperature for all feedstocks. Carbon monoxide and hydrogen are increasingly formed at high pyrolysis temperatures due to gasification reactions. Tar components are decreasing with increasing temperatures. A distinct release of about 70 % of the fuel nitrogen begins above 600 "C,but with rising temperatures the level of release is not increasing anymore. About 20% of the fuel nitrogen stays in the char and is separated from the gas within the filtration.
1449
ACKNOWLEDGEMENT
The research is h d e d in part by the Commission of the European Community,Joule I11 Research Programme (JOR3-CT95-0057),and the Bundesministerium fiir Wirtschaft (Federal German Ministry of Economic Affairs) via the A S Arbeitsgemeinschafi industrielle Forschung ‘Otto von Guericke’ e.V. (industrial research foundation). REFERENCES 1
2 3 4 5 6 7 8
9
10
11
12 13 14 15
16
Albrecht W.: NO,-Emissionen aus Kohlestaubflammen. VGB-Kraftwerkstechnik 72, Heft 7,1992 Beer J.: NO-Formation and Reaction in Fluidized Bed Combustion of Coal. J. Inst. Energy 54,198 1 Diirselen H.J.: Feuerungsseitige N0,Minderung bei Braunkohlenstaubfeuerungen, Einflu8 von brennstoff- und betriebsseitigen Parametem. Dissertation, Ruhr-Universitat Bochum, 1992 Evans R.J., Milne T.A.: Molecular Characterization of the Pyrolysis of Biomass. Energy & Fuels, Vol. 1, No. 2, 1987 Fenimore C.P.: Formation of Nitric Oxide in Premixed Hydrocarbon Flames. XI11 Comb. Symp., New York, 1971 Fenimore C.P.: Reactions of Fuel-Nitrogen in Rich Flame Gases. Combustion and Flame 26, 1976 Greul U.: Experimentelle Untersuchung feuerungstechnischer NO,-Minderungsverfahren bei der Kohlenstaubverbrennung. VDI Fortschrittberichte, Reihe 6, Nr. 388, VDI Verlag, Dusseldorf, 1998 Greul U., Magel C., Moersch O., Rudiger H., Storm C., Schnell U., Spliethoff H., Hein K.R.G.: Einsatz von kohlestihmigem Pyrolysegas als Reduktionsbrennstoff. Final Report, BMBF-Research Project No. 0326766A; IVD Berichte, Bd. I , Institut fiir Verfahrenstechnik und Damptlcesselwesen, Universitat Stuttgart, Stuttgart, 1997 Hein K.R.G.: NO,-Minderung bei Feuerungen fiir Rheinische Braunkohlen. Konferenz Einzelbericht, VGB Seminar Rauchgasreinigung und Reststoffentsorgung in Kraftwerken, Industrie- und Heizkraftwerken, Technische Vereinigung der GroBkraftwerksbetreiber VGB, No. 7, Pp. 1-24, Cottbus, 1990 Kicherer A., Gerhardt Th., Spliethoff H., Hein K.R.G.: Co-Combustion of Biomass / Sewage Sludge with Hard Coal in a Pulverized Fuel Semi-Industrial Test Rig. Final Report, EC-Research Programme, APAS-Contract COAL-CT92-0002, Institute for Process Engineering and Power Plant Technology, University of Stuttgart, Stuttgart, 1995 Pershing D.W., Wendt J.O.L.: Pulverized Coal Combustion: The Influence of Flame Temperature and Coal Composition on Thermal and Fuel NO,. XVI Symposium on Combustion, 1976 Petzel H.-K., Scholl G.,Tigges K.-D.: Modernste Verbrennungstechnologie zur Primarredizierung non NO,. VGB Kraftwerkstechnik, Jg. 73, No. 3, Pp. 231-237, 1993 Rildiger H. : Pyrolysegas von festen biogenen und fossilen Brennstoffen zur Erzeugung eines Zusatzbrennstoffes fur Feuerungsanlagen. Berichte aus der Energietechnik, Shaker Verlag, Aachen, 1997 Schu G.:Experimentelle Untersuchungen zur selektiven nichtkatalytischen Reduktion von Stickoxiden an einem Flammrohrkessel. Dissertation, Technical University of Munich, Munich, 1989 Schuler J.: Laboruntersuchungen zur NO-Bildung bei der Kohleverbrennung sowie NOReduktion an Pyrolysekoksen. Dissertation, University of Essen, Essen, 1988 Spliethoff H., Rudiger H., Greul U., Spliethoff H., Hein K.R.G.: Kombinierte Minderung
1450
17
18
19 20
der NO,-Bildung und Reduzierung von gebildetem NO, bei der Verbrennung von Steinkohle, Phase 3. Final Report, Research Project No. 0326535C, Bundesministerium fCr Forschung und Technoiogie BMFT, Bonn, 1993 Storm C., Harter M., Spliethoff H., Hein K.R.G.: Untersuchungen von NO,MinderungsmaBnahmenmittels Reduktionsgasen aus Kl&schl&unen in Kohlenstaubfeuerungen. Final Report, AiF Research Project No. 10640 N; IVD Berichte, Bd. 3, Institut fir Verfahrenstechnikund Dampfkesselwesen, Universitst Stuttgart, Stuttgart, 1998 Storm C., Harter M., Spliethoff H., Hein K.R.G.: Generation of a Gaseous Fuel by Pyrolysis of Biomass for Additional Use in Coal-Fired Boilers. Final Report, EC Research Project No. JOR3-CT95-0057; Institut fur Verfahrenstechnik und Dampfkesselwesen, UniversitP Stuttgart, Stuttgart, 1999 Suuberg E.M., Teng H., Calo J.M.: Studies on Kinetics and Mechanism of the Reaction ofNO with Carbon. XXIII Comb. Symp., 1990 Zeldovic Y.B.: The Oxidation of Nitrogen in Combustions and Explosions. Acta Physicochimica URSS 21, 1946
1451
Combustion of Bio-oil in a Gas Turbine R. Strenziok, U. Hansen, H. Kiinstner Institute of Energy and Environmental Technology University of Rostock, Germany
ABSTRACT: Based on a series of test runs in a combustion test facility and a gas turbine it could be shown that bio-oil from pyrolysis of wood can be employed for energy purposes. The fuel has to be clean without solids and of a homogenous consistency. The gas turbine combustion behaviour was examined in numerous tests during 1999 and 2000. The optimisation of the operating parameters was performed with adjustment of the fuel viscosity and fuel pressure in different load regions. The results set the priority for further investigations. The adaptation of the industrial gas turbine T2 16 to the wood based bio fuel was restricted by the design to part load. The emissions were measured for both bio-oil and diesel fuel operation. When compared to diesel fuel, characteristically,the emissions of CO and HC are higher for bio-oil at part load operation. The results indicate slightly lower NOx emissions for bio-oil.
INTRODUCTION A study of the future energy supply in Germany has shown that, five percent of the primary energy demand and ten percent of the electricity production in the year 2010 could be supplied from renewable energy. Also bio-mass as a renewable resource and an environmental friendly energy carrier is assigned an increasing importance for future energy supply. As a clean fuel bio-oil has a number of competitive advantages over diesel fuel, no sulphur dioxide and lower nitrogen oxide emissions. Furthermore, it has a favourable carbon balance contributing little to greenhouse gas emissions. Bio-oils have so far mainly been in the focus as an alternative fuel for internal combustion engines. The fuel obtained from fast pyrolysis, however, displays properties which do not lend themselves easily to this application. Intermittent combustion under high pressure in an reciprocating machine appears more demanding than continuous combustion under lower pressure in the combustion chamber of a gas turbine. This is the working assumption for the studies reported below. At the Institute of Energy and Environmental Technology the combustion behaviour of wood based bio-oil has been investigated in a combustion test facility of 300 kW-thermal capacity as well as in a gas turbine of 75 kW-electric capacity. The investigations were carried out within the framework of the EU- demonstration project
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"Scaling up and operation of a flash pyrolysis system for bio-oil production and application on the basis of rotating cone technology" (1 997-2000). The bio-oil used for the test programme was produced by BTG in Enschede (NL) with the Rotating Cone Technology from poplar as the bio-mass material. As alternative raw materials for pyrolysis wheat straw, Spanish thistle and other wood wastes have been investigated. The main objective of this paper is to demonstrate the power generation from bio-oil in a gas turbine, the optimisation of the combustion behaviour with respect to emissions and operation characteristics and the comparison of emission and operation data from bio-oil and conventional diesel fuel. For technical reasons the gas turbine had to be adapted to the operation in dual fuel mode with the possibility to switch between diesel and bio-oil. BIO-OIL CHARACTERISTICS
The bio-oil used in the tests was delivered from the project partner BTG in April 2000. For the chemical and physical analysis of bio-oil we used modified standard fuel oil methods. Table 1 shows the chemical and physical properties of the oil. The analysis was carried out in the Institute's chemical laboratory. Table I Properties of the bio-oil used in the gas turbine
The oil was produced from non-contaminated soft wood from Spain. Due to the low pH-value, acid proof materials had to be used in the fuel supply system. The pure biooil need to be preheated up to 50 "C to achieve a sufficient viscosity. Bio-oil from pyrolysis contain considerable amount of water between 20 to 30 per cent. However, combustion in the gas turbine was not hampered by the presence of water as the ignition of the bio-oil was initiated by the pilot flame burning diesel fuel. Bio-oil from pyrolysis is immiscible with petroleum based fuels and therefore a stable mixture of bio-oil and diesel fuel could not be obtained and tested in the gas turbine.
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EXPERIMENTS The experimental work was performed with a small commercial gas turbine type T 2 16 with a rated electric power output of 75 kW.It has a single shaft, single stage radial turbine and a single stage radial flow compressor with a pressure ratio of approx. 2,5. The turbine shaft speed ranges from 50 000 to 30 000 m i d and with a two stage reduction gear the turbine is coupled to a synchronous generator. For test purposes it is important that the combustion chamber is of a silo type arranged tangentially. A picture of the gas turbine is shown in Fig. 1.
Fig. 1 Gas turbine T 2 16 with exhaust measurement device
Fig.2 Gas turbine test bed for Dual Fuel operation
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Extensive alterations to the fuel system was necessary in order to bum bio-oil from pyrolysis. The flow diagram in Fig. 2 shows the layout of the gas turbine system after modifications. The temperatures and pressures refer to diesel fuel operation. The operation on bio-oil alone is not possible because the speed safety governor and the lubricating oil cooler is designed for diesel oil operation. Therefore some proportion of fuel must be diesel. The combustion chamber of the gas turbine was fitted with two inline fuel nozzles, a main nozzle and an ignition nozzle (Fig. 2). The ignition nozzle was operated with diesel oil. The supply of bio-oi1 to the main nozzle was made through a separately driven, external fuel pump (Dual Fuel Mode). It is designed for a maximum flow of 455 l/h at a pressure up to 80 bar and the effective flow controlled in a bypass. The flow through the main nozzle can be switched between diesel and bio-oil. The proportion of fuel mass flow of the main nozzle and ignition nozzle is given in Fig. 3.
-E
60
-
50
.
.a'=-
m
2 40
-
0 -
-
a*-
c
2
30 .
.. ~
l i l
c Y - '
- - --
- / -
e= m
g
20 -
3
10.
-
.
-
Diesel fuel main nozzle 6.5 GPH biooil main nozzle 6.5 GPH -Diesel fuel ignition nozzle 5.0 GPH -A4
0,
~~~
~
Fig.3 Mass flow rate of ignition nozzle and main nozzle
Fig. 4 External bio-oil pump The bio-oil from pyrolysis may contain solid residues and is highly viscous. Filtering and pre-heating are necessary before injection in the main nozzle. For the gas turbine tests the following properties were achieved
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-
particle size < 20 Micrometer injection viscosity about 12 cSt. The gas turbine was started with diesel oil. Then followed a change over of the fuel through the main nozzle from diesel oil to bio-oil. Due to the lower heating value of the bio-oil it was only possible to operate the gas turbine at part load in the dual-fuelmode. With the mass flows in Fig. 3 the following power levels are calculated and shown in Table 2. Table 2 Comparison of turbine power in diesel fuel and dual fuel operation
I Turbine power Ignition nozzle 5.0 GPH Main nozzle 6.5 GPH Total
I Diesel fuel Operation P = 354 kWth (p = 26 bar) P = 437 kWth (p = 26 bar) P = 791 kWth
I Dual fuel operation I P = 354 kWth (p = 26 bar) P = 226 kWth (p 30 bar) P = 580 kWth
-
The mass flow of bio-oil in relation to diesel fuel increases in the main nozzle with 6.5 GPH (gallons per hour) by up to 1.33. The power in the dual fuel mode is reduced to 73 % of the full power in the diesel fuel mode. The reduction is due to the specific limitations of the gas turbine used for the experiments and not inherent in bio-fuel operation. However, for later commercial use further alterations to the fuel system and the nozzle size would be required. The high viscosity makes pre-heating necessary and with high pre-heating temperatures the danger of coking in the main nozzle arises. In order to prevent overheating in the nozzle it is equipped with a cooling adapter and a thermocouple.
Fig. 5 Deposits in the gas turbine combustion chamber after bio-oil operation
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After short combustion tests in the gas turbine deposits were found in the combustion chamber as shown in Fig. 5 . The deposits are of a lacquer like nature and can only be removed mechanically. At present, the deposits in the combustion chamber do not constitute a problem for the operation of the gas turbine. However, there are also slight deposits on the turbine blades. If the deposits are to be removed, the turbine needs to be taken apart. Due to the high speed of the turbine larger deposits cannot be tolerated. Means to prevent fouling of the turbine is a priority requirement. The extent to which fouling occurs may depend on the quality of the bio-oil and can only be determined from duration tests.
EXHAUST EMISSIONS The equipment for measuring and analysing emissions is shown in Fig. 1 and 2. The oxygen content "as measured" is 17.5 %. The excess air ratio is in the diesel fuel operation 3.3 and in the dual fuel operation about 6.
Fig. 6 Bio-oil emissions in the gas turbine in relation to diesel operation (1 00 YO) When compared to diesel fuel the emissions of CO and HC are higher for bio-oil. The results indicate slightly lower NOx emissions for bio-oil. All measurements were taken at part load operation due to the limitation of the gas turbine. The initial tests in dualfuel-mode were performed with a relatively "cold" combustion at part load which may be the reason for the incomplete combustion of the hydro carbons. Also the combustion air fuel ratio was not adjusted for bio-oil operation at part load. The combustion of bio-oil is clearly recognised by the exhaust gas odour and a light exhaust gas cloudiness. The tests performed at the Institute for Energy and Environmental Technology are among the first involving exact measurements of emissions and are comparable to emissions data found in the technical literature (Andrews and Fuleki, 1997). Due to limitations of the particular gas turbine and restricting the test series to one type of biooil the results cannot claim to be generally valid for the comparison of bio-oil from pyrolysis to a standard fuel like diesel oil but should give some indication of the qualitative outcome. The studies have demonstrated the possibility to bum bio-oil in a gas turbine for the production of electricity and heat.
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The principal modifications involve the fuel handling and the fuel supply system. Basic results and insights of bio-oil combustion were obtained in the combustion test facility. REFERENCES 1.
2.
3.
4.
5. 6. 7. 8.
Bridgwater A. & Hogan E. (1996) Bio-oil Production and Utilisation. Proceedings of the 2"d EU- Canada Workshop on Thermal Biomass Processing. Cpl press , ISBN 187269151 X Andrews R. G. & Fuleki D. (1997) Results of industrial gas turbine tests using a biomass derived fuel. Proceedings of the Third Biomass Conference of the Americas. Oasmaa A. & Lepp2maki P. (1997) Physical characterisation of biomass-based pyrolysis liquids. VTTPublications No. 306, Espoo Bridgwater A. & Czernik S. (1999) Fast pyrolysis of biomass: A Handbook. Cpl press, ISBN 1872691 072 Hansen U. (1999) Einsatz von Pyrolyseol als regenerativer Energietrager zur dezentralen Versorgung mit Strom und Warme. Freiberg, Germany. Strenziok R. (1 999) Methods and problems of bio-oil analysis. PyNe-Workshop Montpellier, France. Morris K. & Piskorz J. (1999) Bio ThermTM: A Systemfor Continuous-Quality, Fast Pyrolysis BioOil. 4'h Biomass Conference of the Americas. Meier D. & Faix 0. (1999) State of the art of applied fast pyrolysis of lignocellulosic materials- a review. Bioresource Technology 68
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Stirling Engine with Flax@ Burner Fuelled With Fast Pyrolysis Liquid Andreas Bandi and Frank Baumgart Centerfor Solar Energy and Hydrogen Research Baden- Wuerttemberg, (ZSW),Hessbruehlstr. 21 C, 70565 Stuttgart, Germany
ABSTRQCT Fast pyrolysis liquid was combusted with a modified propane gas FLOX’ burner (WS) mounted to a Stirling CHP unit (SOLO 25 kW,). The propane burner was modified by an air pressure atomiser designed for pyrolysis liquids. Nearly 110 hours of operation with different fuel loads have been achleved, engine performance and emissions were recorded. The Stirling tests proved that pyrolysis liquids can be burned efficiently and with low emissions in a FLOX’ mode operating burner attached to a Stirling CHP unit.
INTRODUCTION In the fiame of the EU-project “Small-Scale Combined Heat And Power Production (Chp) From Bio-Crude Oil (Bco) Fuelled To A Stirling Engine” (JOR3-CT98-0310, co-ordinator CRES, Greece) ZSW is investigating in close collaboration with industrial partners, WS, SOLO, Germany, the use of BCO in a Stirling CHP unit with a FLOX@burner. ”FLOX is acronym for ”FLame-less Oxidation”. The special feature of FLOX’ burners, developed by WS, Germany, is that at temperatures as high as 850-900°C thermal NO, is drastically lowered by internally mixing combustion air and exhaust gas, avoiding temperature peaks of the flame. This is even true at high air preheating temperatures. FLOX@has an additional advantage for burning of pyrolysis liquids. As the residence time of the fuel in the burning chamber, is higher than in a normal burning, a more efficient carbon burnout becomes possible. At present FLOX’ burners fuelled by propane or natural gas are used for high temperature process heat production. A future, promising application of these burners is the use in Stirling CHP units (SOLO, Germany). Within this project ZSW designed and realised an injection-atomisation system for BCO. A WS propane burner [ 1,2] was modified with a ZSW atomiser and attached to a SOLO Stirling engine 161 (4-9 kW,, 10-25 kW,) [3]. This paper presents experimental results on the burning of BCO in a FLOX@burner and in a Stirling CHP unit with a FLOXB burner.
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The combustion of bio crude oils has been tested so far in stationary Diesel engines, boilers and gas turbines [4]. There are several atomisation systems available for combustion of mineral oils. Based on physical properties, BCO can be classified between light fuel oil and low sulphur heavy oil. However BCO has a number of unique properties, such as high acidity (PH c 2-3), low heating value (Lower Heating Value, LHV = 15-19 MJkg), high viscosity (30-2000 cSt at ambient temperature) [ 5 ] and high instability, that makes it difficult to use as a combustion fuel. The wide range of organic components with different boiling temperatures and the char content have a negative effect on the emissions (CO, HCs and soot) by burning BCO. In addition, the high water content (generally > 20%) and the high oxygenated molecule concentration impair the ignition and increase the colung tendency.
ATOMISATION EXPERIMENTS
Viscosity, solid particle content, particle size and heating value of the liquids are determining factors in the selection of the proper atomisation system. The atomisation technologies available include pressure atomisers, air or steam pressure atomisers and rotary atomisers. Pressure atomisation is mostly used for the combustion of oils with low viscosity (2.5-5 cSt), e.g. light oils. In this system a pump raises the pressure up to 10 and more bar and forces the oil through a narrow nozzle hole producing a fine spray of oil with droplet sizes ranging between 30 and 100 pm. Due to the generally narrow sized nozzle holes, the pressure atomisers tend to get blocked. These atomisers are less suitable for liquids with high viscosity, relative high solid particle content and low fuel charge. The rotary atomisation is suitable only for large liquid flows. For the atomisation of BCO a pressurised air atomiser with external mixing chamber fiom (Schlick-Diisen GmbH, Germany) was selected. Using this type of atomiser the liquid mass flow can be varied independently from the air flow, providing a constant atomisation quality at different BCO mass flow. The selected nozzle hole was 0.8 mm and the spray cone angle 30". In order to provide the proper atomisation temperature for BCO, the nozzle has been equipped with a temperature control (coolingheating) system. The atomiser was designed to operate with different BCO mass flow up to 8-10 lh,by a liquid pressure of about 0.2-1.0 bar. For an air flow of 15 m3/h (2 bar), the droplet size was estimated from the propellant air/BCO mass ratio to about 10-50 pm (Sauter mean diameter). In order to study the atomisation properties of BCO, an experimental setup has been constructed, including the atomiser, storage tanks for BCO and methanol (for cleaning), and control systems for the BCO temperature in the nozzle, liquidair pressure and atomising air pressure. A schematic presentation of the experimental setup is given in Fig. 1. A picture of the equipment is shown in Fig. 2.
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2 ~.......................................,...................*.............~
1. 2. 3. 4. 5. 6. 7. 8. 9.
compressed air valves .pressure adjustment selector valve methanol tank BCOtank atomiser connection plate temperature measurement
---
I..........
-
cooling/heating system BCOlMeOH line atomising air
Fig. I: Scheme of the experimental installation
Fig. 2: Experimental set-up for BCO atomisationtests
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For the atomisation and burning tests BCO supplied by Ormrod Diesels was used. Some of the characteristic properties of this BCO are presented in Table 1. Table I : Characteristic properties of the BCO used in experiments
Mass flow experiments at different liquid pressures were performed at constant temperature. For 30°C and 5OoC the results are shown in Fig. 3. A typical spray cone at 2 bar and a BCO mass flow of 8 l/h is shown in Fig. 4. The mass flow and atomising experiments have been carried out without any difficulties related to BCO properties.
BCO pressure [bar]
Fig. 3: BCO mass flow dependence on pressure (nozzle hole diameter: 0.8 mm) 1462
Fig. 4: Atomisation cone with BCO (nozzle hole: 0.8 mm; BCO flow: 8 yh,air pressure: 2 bar
COMBUSTION EXPERIMENTS WITH A FLOX@BURNER A traditional propane burner of WS was modified for BCO combustion. The burning chamber with the modified burner was pre heated up to 900°C with an auxiliary burner placed opposite to the BCO atomiser (Fig. 5). The nozzle was cooled down to 30°C. Once 900°C was reached in the burning chamber, the BCO injection was started. After the ignition of BCO, which occurred immediately with supply of the first amount of BCO, the auxiliary burner was turned off. The supply of BCO was started with about 2 l/h and than increased to about 9-10 Yh (liquid pressure 0.9 bar). The air supply was adjusted to about 3% O2in the flue gases. The burning of BCO in FLOX@mode continued stationary without any difficulties for several hours. Table 1 summarisesthe main experimentalresults. The experiments proved that BCO can be burned without any noticeable residues or soot formation and with practically no CO emission in stationary F L O P mode. Improvement in NO, emission can be expected by varying excess air and exhaust gas recirculation ratio. However, the nitrogen content of BCO is itseIf a source of NO, whch can not be reduced below a certain limit. The low CO concentration provides evidence for a high quality atomisation which allows a total carbon burn out. Fig. 6 shows the combustion chamber fuelled with BCO ( F L O P mode). The temperatures in the burning chamber (on different places) and in the nozzle were recorded. (see Fig. 5). In order to avoid the nozzle plugging, the BCO temperature in the nozzle was controlled closed to the nozzle tip and was kept at 25-3OoC, therefore a temperature caused blocking (coke formation) can be excluded.
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flue gas
burning chamber
air supply coolinglheating /water inlet
,,,/
\
ID 40 mm 8 holes a 6 mm ID
0
in the nozzle
coolinglheating water outlet
Thermocouples for temperature measurement
Fig 5: Modified WS propane FLO?
burner for BCO combustion
Table. I : Experimentalresults to BCO F L O P burning in a modified WS propane burner
COMBUSTION EXPERIMENTS WITH A STIRLING CHP UNIT EQUIPPED WITH A FLOX@BURNER The BCO-Stirling experiments were conducted with different pyrolysis liquid @ads and air pressures. In order to prevent the blocking of the nozzle during heating-up the engine, the atomising air supply was started simultaneouslywith the start the heating. After reaching the operation temperature, BCO supply was started and the auxiliary burner was turned off. The burning of BCO in FL,OX@mode was stable without any diniculties related to the engine and the burning of BCO. The CHP test rig is shown in Fig. 7.
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Fig. 6: Combustion chamber &elled with BCO (FLOX@mode, WS) In order to avoid the blocking of the nozzle due to the solid particles and agglomerates in the BCO, a mechanical cleaning system was integrated in the atomisation system. The cleaner was activated only when due to break of BCO supply a temperature decrease at the heater was observed. During the experiments engine performance and emission data were recorded. A summary of the records is presented in Table 2. The Stirling tests proved that fast pyrolysis oils can be efficiently burned in FLOX@ mode with low emissions. Several trial runs were carried out in order to improve the engine performance and emission characteristics. BCO flow rate, aidfuel ratio have been varied and emissions have been measured. Ash deposition on the burning chamber walls and on the heat exchanger was not observed. Electrical and thermal efficiencies of the CHP unit were not satisfactory (overall efficiencies 50-60 %). It is considered that the main reasons for low efficiencies are the introduction into the engine of not preheated atomiser air, and the burning chamber geometry. As indicated by the temperature measurements, at high fuel loads the burning takes place behind the heat exchanger, causing heat loss for the engine.
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7
Table 2: Emission and engine performance data for Stirling CHP unit operation
Atom. air press., bar Soot numbex
BCO flow, vh
Emissions
m m3 4.40
2.5
6.46
2.0
m m3
m m3
20-40
6.1
I
I
137
7*97 5*0 I2O l4 German Emission Standards for diesel engine < 5 MWth (02 concentration 5%): NO& 500 mg/m3;CO: 650 mg/m3;HC: 100 mg/m3;Soot number: 2 O2concentration in flue gas: 6- 10%
I
6*2
Fig. 7: Stirling CHP unit test system
CONCLUDING REMARCS (1) Experiments carried out with modified propane burner proved that BCO can be combusted in FLOX@mode without noticeable residues with low emissions.
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(2) Due to its special feature, the FLOX@ mode is highly suitable for BCO combustion. (3) Combustion experiments of BCO in a Stirling CHP unit with F L O P burner were conducted for nearly 110 hours with different loads. (4) After preheating the burning chamber up to 800-850 OC with an auxiliary burner, the BCO ignition and combustion were carried out without any difficulties. ( 5 ) Blocking of the atomiser nozzle due to the agglomerates and solid particles contained in the pyrolysis liquid created difficulties in the fuel supply. (6) The measured emission figures with Stirling CHP unit were below the German Standards for diesel engines in similar power range
REFERENCES 1. Wiinning, J., "GASWAMEZnternationaF', Vol. 47, Vulkan-Verlag, Essen, 1998,
pg. 322 2. Wunning, J., Regemat Burner, WS report, 1999 3. Baumiiller, A., G. Lundholm, L. Lundstram and W. Schiel, Development History of the V 160 and SOLO 161 Engines, International Stirling Engine Conference, South Africa, 1999 4. Shaddix, R. C., G. R. Hardesty and S. Gust, Combustion Properties of Biomass Flash Pyrolysis Oils: Final Report, Sandia National Laboratories, Albuquerque, 1999
5.
D. Meier, A. Oasmaa, G.V.C.Peacocke, in "Developments in thermochemical biomass conversion", ed. A. V. Bridgwater and D. G. B. Boocock, Blackie Academic & Professional, London, 1998, pg. 394
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Pyrolysis oil combustion tests in an industrial boiler Anja Oasmaa*, Matti KytO**, Kai Sipila* * VTTEnergy, P.O.Box 1601, 02044 m,Finland ** Oilon Oy, P.O.Box 5, 15801 Lahti, Finland
ABSTRACT Combustion properties of various pyrolysis oils were studied at Oilon (Lahti, Finland) in a test boiler on 4 MW output level. Burner settings and boiler dimensions as well as oil properties were tested. A special emphasis was on oil specifications for boiler application. It was concluded that pyrolysis oils can be combusted in slightly modified conventional oil furnaces and boilers fairly well. Some existing oil burners are also capable to combust pyrolysis oils, but they might need slight modification. However, oil specifications must be defined, including at least water and solids. Special care should be taken in handling and pumping of oils. Support fuel is needed at the burner to start the combustion and recommended for supportingpyrolysis oil in case of poor-quality oil. Emissions depend on oil quality. They are generally equal to those of light fuel oil or light bunker oil, but the content of particles in exhaust gases is higher. The results are applicable with small modifications in the boiler size class down to 1 MW*. Research into replacing light fuel oil in boilers below 1 MW&continues.
INTRODUCTION The industrial organisations supporting VTT in the biomass pyrolysis R&D are Vapo Oy, Fortum Oil & Gas, and Oilon Oy. Vapo, the biggest domestic biofuel producer in Finland, is interested in the production of fast pyrolysis liquid from domestic fuels. Fortum is an international energy company, which was formed from the Neste and N O Groups. Fortum's Oil & Gas Division explores and produces oil and gas, refmes, sells and markets petroleum products, and transmits, distributes and sells natural gas. In pyrolysis area, Fortum is focusing on the use of pyrolysis liquid in medium-size boilers [l]. Oilon is the biggest burner manufacturer in Finland and is interested in boiler applications in burner size class of 350 kW-45 MW. Additionally, Wartsila NSD Oy has carried out a study on using pyrolysis liquid as diesel power plant fuel together with VTT in the 1990s [2]. At VTT, the following issues are addressed in the present study: pyrolysis of solid biomass [3, 41, hot vapour filtration, pyrolysis oil quality [5] and fuel oil specifications [6],storage and handling properties of pyrolysis oil, boiler applications [7], and techno-economic assessment of pyrolysis systems [8]. A considerable amount of work 1468
has been carried out on testing and further developing standard fuel oil analyses and test methods for pyrolysis oils [9, 10, 11). In the EU Project JOR3-CT95-0025, co-ordinated by VTT Energy, i.a. the quality of pyrolysis oil of pilot plants was controlled, using Swedish pine as raw material [12]. The oils produced were used in combustion tests at Oilon Oy. Two main targets were set for the combustion tests at Oilon: Optimisation of the boiler unit and combustion conditions in 4 - 5 MW size class. With small modifications, the research results will be applicable up to the burner output of 45 MW. Comparison of combustion properties of various qualities of pyrolysis oil. Other aims included: Identification of burner types, in which pyrolysis oil can be burned, and definition of critical issues in successful combustion Homogenisation of phase-separated inferior-grade pyrolysis oil, and determination of its combustion characteristics Development of a feed system for pyrolysis oil Study of material corrosion.
MATERIALS AND METHODS TEST ARRANGEMENT The combustion tests were carried out in the combustion test laboratory of the R&D Center of Oilon Oy employing a test furnace of 8 MW nominal capacity manufactured by TKH (Finland). The test furnace was cylindric, the inner diameter of the furnace being 2.4 m and length 5.2 m. A number of changes were made in the furnace during the combustion tests with pyrolysis oils. The flue gases escape upwards from the rear of the furnace. The fiunace is watercooled. During the run,the temperature of incoming circulation water ranged 60 85 "C and that of return water 80 - 105 "C (1.27 bar). Lots of the heat released in combustion escaped along with flue gases, and hence, the flue gases were rather hot (300 - 500 "C). There were several viewports in the furnace for visual monitoring of the flame. The main components and apparatuses of the combustion test equipment were Feed tanks of pyrolysis oil: two 1.5 m3steel tanks with propeller mixers Feedpump Preheater of pyrolysis oil: electrical mass preheater of Oilon, oil flowing in pipes inside it Dual fuel oil lance and nozzle suitable for dual fuel combustion Modified Oilon Lenox GRT-5L dual fuel burner Forced-draught fan controlled by an inverter Oilon test furnace: nominal power output 8 MW, test run output about 4 MW Measuring assembly for flue gas emissions. 0
The burner of type Oilon Lenox GRT-5L is a mixing burner for oil and gas combustion in boilers. The burner was equipped with the dual fuel lance, which enables the use of two different liquid fuels either separately or simultaneously. Heavy fuel oil (POW 180) was used as support and start up fuel. A KM-nozzle type 1469
2388802 (Y-jet principle) of Oilon with atomisationmedium was used in the dual fuel oil lance. Compressed air of about 6 bar was mainly used for atomisation in the tests. The combustiontests were divided into three main tasks: 1. Testing and optimisation of combustion conditions with one pyrolysis oil - The aim was to test the combustion of pyrolysis oil and to adjust the combustion equipment and slightly modify the used burner 2. Combustion tests with typical pyrolysis oils - The aim was to run comparable combustion tests with different 'Yypical" pyrolysis oils 3. Combustiontests with pyrolysis oil high in water - The aim was to apply, if needed, support combustion for difficult pyrolysis oils (high water content, inhomogeneity) and run comparable combustion tests. MEASUREMENTS AND SAMPLES The following systematic emission measurements were carried out during the tests: Gasmet gas analysis for emission gases: COYCOz, NO, NOz, N20, Cb,C2&, H20as momentary measurements (IR-principle) Continuous backup measurement for emission gases: COYNO and NOz Continuous oxygen measurement with two separate measuring devices Particles content measurement according to standard SFS 3866, isokinetic sampling for the filter Soot measurement according to standard EN 267 with a Bacharac standard pump, and numerical value with an optic tone value gauge.
NOx concentration in the exhaust gases is calculated on an anhydrous oil basis from dry exhaust gases according to SFS 5624. In this article, "particles" mean dust, particulates and solid particles contained in flue gases and measured as weight increase of the filter. Particles differ from soot due to the measuring method. Soot means the darkness level of the spot taken with the Bacharac pump and soot paper,and this sample is compared between absolute black and white. The cleanliness of the boiler was determined by visual observations inside the boiler after the test runs. Comparisons to the heavy fuel oil runs were done. The cleanliness of the nozzles was determined by checking each hole of the nozzle by a steel wire after the test run. The blockage of the nozzle holes was also observed visually through the small window on the back wall. Samples were taken as follows: Pyrolysis oil samples after the preheater for each measuring point Possible coke or impurity sampIes from the nozzle after the test cycle Possible coke samples from the boiler after the test cycle. The samples were analysed according to fuel standard modified for pyrolysis oils
[9,11, 101, PYROLYSIS OILS The pyrolysis oils chosen for the Oilon combustion tests were produced from various hardwoods and softwood using either a bubbling fluidised bed or circulating fluidised bed. The capacity of production ranged up to 1 t/h. Since the oils are not yet commercial and producers still try to improve them, the oils are marked as in Table 1. 1470
Table I Pyrolysis oils in Oilon combustion tests. Oilon Class Feedstock Process No. 1-6.6 Testing Hardwood 1 CFB* 9 Combustion tests Hardwood 2 CFB* 10 + 11 Combustion tests Hardwood 3 CFB* 12 + 13 Combustion tests Pine 1 CFB* 15 Combustiontests Pine 2 BFB** 7-8 High-water content Pine 3 BFB** * Circulating fluidised bed **Bubbling fluidised bed
Production date 11/1996 2/1999 2/1999 3/1999 311999 6/1998
Hardwood pyrolysis oil 1 (Oilon 1-6.6) was used as "test oil", its amount being largest. The oil was inhomogeneous due to a long storage time (2.5 years) outdoors. The oil was homogenised with 20 vol% (abt. 10 wt?h) methanol addition. Simultaneously, the viscosity of oil reduced from about 200 cSt to 35 cSt (@ 50 "C). Oil was stirred with a propeller mixer slowly continuously after methanol addition to keep it homogeneous. Hardwood oils 2 and 3 and Pine oih 1 and 2 were classified as "typical oils'' for the combustion tests. The most important fuel characteristics are heating value, viscosity, and water and solids content. These characteristicsof all oils used in the measurements are presented in Table 2. Nitrogen content was below 0.1 wt?? for each oil. Figure 1 shows the viscosity curves as a function of teqerature for all pyrolysis oils used for combustion.
Table 2 Fuel oil analyses of pyrolysis oils used for combustion. OILON No. 6.2 6.3 6.4 6.5 6.6 9 10+11 12+13 15 7+8
(Testing) ' (Testing) ' (Testing)' (Testing) ' (Testing)' (Typical) (Typical) (Typical) (Typical) ' (High-waater content)
Feedstock
LW
Viscosity 50 'C
Water
Solids
cst
Wt%
Wt%
Hardwood 1 Hardwood 1 Hardwood 1 Hardwood 1 Hartj,vood 1 Hard# ood 2 Hardwood3 Fine 1 Fine2 Rne3
Wks 17.4 17.3 17.4 17.2 17.5 17.6 16.4 15.3 15.3 11.2
42 38 40 37 42 41 33 27 6
20.4 21.o 19.3 21.1 19.9 20.7 24.8 22.6 24.5 35.7
0.40 0.41 0.38 0.40 0.40 1.20 1.86 0.17 0.25 0.03
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6
Fig. 1 Viscosity curves of pyrolysis oils used in the Oilon combustion tests. Testing oils: light oil POK15 used by big real estates, marine purposes and industry, and heavy he1 oil PORL 180, most generally used in size class of 1 10 Mw. * Contains about 10 wt% methanol.
-
RESULTS ADJUSTMENT OF COMBUSTION SET-UP Several adjustment combinations to the feeding and boiler set-up were tested. During previous tests at Birka Energi [13] it was concluded that a special type of front head was needed inside the boiler for preventing the heat loss. This was used also at Oilon. The use of an extra cylinder inside the fiunace yielded a more "dense" flame, higher temperature and hence faster volatilisation. The burner head was changed from a diverging type to a converging one which yielded a m o w e r and more intense flame. A swirler of 60" yielded a stable flame although a large amount of primary air was used.
ADJUSTMENT OF COMBUSTION C0NDIT.IONS Combustion conditions were optimised and basic regulations of the combpstion system were adjusted by burning "test oils". Simultaneously, an operation model for the basic measurement was finished and the operation of measuring assembly was tested. In total 11 m3 of hardwood 1 pyrolysis oil was burned in the test runs. 10 wt% of MeOH was added to the test oils to increase the homogeneity of oil, to improve the combustibility and ignition and to reduce particles emissions. In part for the same reason the highest NO, contents were measured for these oils. The hghest NO, level, 121 mg/MJ was measured for the run,in which the pyrolysis oil had been heated to 80 "C prior to atomisation. However, the difference in NO, emissions was 1472
not great when compared to the test run,in which the temperature of oil was 54 "C (1 17 mg/MJ). On the contrary, the amount of incombustibles increased clearly when the temperature of the oil was raised. The mean combustion results (oils at about 50 "C prior to atomisation) for the oils are given in Table 3. Table 3 Mean combustion results for test oils. 0 2
vol% 3.45
NO, mg/MJ 109
co mg/MJ 1.2
HC mg/MJ 0.3
Soot Bac. 2.1
Particles mg/MJ 19
In general, the test oils burned relatively well and the emission values were fairly good. No si&icant problems appeared in the combustion assembly, in particular after determining the basic regulations. The flame was usually unbroken and stable, and no sigmficant splash occurred. The boiler did not foul due to pyrolysis oil after the initial stage. COMBUSTION TESTS WITH TYPICAL PYROLYSIS OIL GRALlES
The lowest NO, contents were achieved for "typical pyrolysis oils" (Table 2). The reason for this was apparently the lack of MeOH and perhaps also the higher water content. It was typical of these oils that the particle emissions were on average 4.5fold, but respectively, the NO, emissions were about 20% lower, compared to the values measured for the test oils. The soot number (average 2.4) was not much higher than that measured for the test oils (average 2.1). The combustibility of the oils was relatively good, with the exception of Pine 2 pyrolysis oil. The problems with this oil appeared primarily in the nozzle, which clogged quickly despite regulation measures and other efforts. Hardened, floury, dry and crumbly particles were find in the holes of the clogged nozzle. The measurement run could be carried out only by adding methanol. This phenomenon was not found for the other oils of the group. Methanol might have dissolved crust and/or prevented its formation for a time. The oil was otherwise of relatively "high grade". The average combustion results for typical pyrolysis oils are presented in Table 4. The emission values for heavy and light fuel oil and natural gas using same! boiler with same! adjustments are also presented as Testing. The emission values using a low-NOx system are in brackets. In Finland the current emission standards for 50 - 150 MW boilers are 120 mg/MJ NO,. In smaller boilers NO, is not specified and the limits for particles are: 90 mg/MJ in 1-3 MW, 60 mg/MJ in 3-5 MW and 40 mg/MJ in 5 - 50 MW boilers. Table 4 Average combustion results for typical pyrolysis oils. Emission results for heavy fuel oil (HFO), light fuel oil (LFO),and natural gas (GAS) are included. 0 2
vol% Pyro HFO LFO GAS
3.5 3.5 3.5 3.5
NO, co mg/MJ mg/MJ 88 4.6 193 (75) 3(10) l(5) 70(38) 55 (27) l(6) 1473
Soot Bac. 2.4 2(3) OS(1) 0 (0)
Particles mg/MJ
86 23 (28) 2(2) 0 (0)
The values presented in Table 4 are average emission values obtained for all typical pyrolysis oils in combustion tests. The combustion tests with typical pyrolysis oils usually proce'eded relatively well. The expected difficulties and possible problems were related to the stability of flame, combustion characteristics of oil, simultaneous achievement of low emission levels, fouling of furnace, and cleanness of nozzle devices. Changes also occurred in these dwing the tests. For example, the difficult pine 2 oil would not have burned during the measuring cycle (30 min) without special monitoring and "manipulation". On the other hand, the combustion of the best oils could have been continued a long time without any changes or problems in combustion or fouling. Figures 2 4 show photos of a typical stable flame of clean pyrolysis oil, a flame of a difficult pyrolysis oil, and a flame of pure support fuel.
-
Fig.2 Typical stable pyrolysis oil flame without problems.
Fig. 3 Difficult unstable and broken pyrolysis oil flame.
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Fig. 4 Flame of support hel.
A particle sample collected from the flue gases of the No. 12+13 pine pyrolysis oil was studied by light and electron microscopy. Single particles contained a lot of inorganic substances, mainly calcium, other substances being iron, silicon, magnesium, aluminium, potassium, chrome and vanadium. The amount of organic substances (e.g. tars) or carbon is likely small. COMBUSTION TESTS WITH "WATER OILS"
The combustion tests with pyrolysis oils containing an abundance of water (Table 2) proceeded in general relatively well. The expected dificulties and possible problems were related to the stability of flame, combustion characteristics of oil, simultaneous achievement of low emission levels, fouling of fimace, and cleanness of nozzle devices. Changes also occurred in these during the tests. The use of support flame immediately improved the stability and combustion in general, as i.a. the particles emission immediately reduced by about 15%, but respectively the nitrogen oxides increased by about 50%. Support he1 power of about 1 MW (25%) is sufficient for securing good combustion. On the other hand, combustion proceeds reasonably well without any support fuel, ifthe other parameters are correct. Combustion of water-rich oil could have been continued much longer than the measurement required. No dramatic changes or problems were monitored in combustion or fouling during the measuring cycle. The average results for water-rich pyrolysis oils in combustion with pure pyrolysis oil and with support fuel are presented in Table 5. Table 5 Average combustion results for water-containing pyrolysis oils.
Supportfuel
O2 vol%
No Yes
3.50 3.50
NO, mg/MJ 75 114
CO
HC
mss/MJ mgMJ 0.6 0.3
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0.1 0.1
Soot Bac.
Particles mgMJ
6.0 6.4
98 83
MATERIAL CORROSION Acidity of pyrolysis oil (PH2 - 3) and the combined effect of specially acids and high water content when raising the temperature set special requirements for materials. Acid-proof steel, certain special steels [11, brass [14], and plastics stand well pyrolysis oil [9]. In these tests, usual carbon steel was wittingly used, as the aim was to test combustion of pyrolysis oil in a common fossil oil burner system. The pipeline and the heat exchanger of pyrolysis oil feed were of St 37.4 DIN 2391 precision steel pipe (size 10 x l), so-called Ermeto-pipe. The pipeline clearly corroded during the tests, which was also seen in an increase in the iron content of oil. It is self-evident that the durability of materials, both steel and seals, has to be ensured in the design of a special pyrolysis oil burner and fuel line to it.
or.HANDLING AND SPECZFICAATIONS The homogeneity of pyrolysis oil is one of the main issues to be notified. A fairly slow (max. 1,200 rpm) propeller type mixing after homogenisation of the oil ensures a steady feed and combustion of the oils. Even an old phase-separated pyrolysis oil can be kept homogenous by slow mixing after adequate methanol addition. An Eccenter screw pump (‘mono’-pump) was used successively for oil feeding. A typical coarse filtering (45 pm) was used through the whole test period without any problems. The best preheating temperature for pyrolysis oils varied depending on the original viscosity of the oil. A good atomisation was obtained for this type of burner and boiler system in viscosity range of 15 20 cSt. The most important fuel oil properties for pyrolysis oils besides the homogeneity were water, solids in the oil, viscosity, and lower heating value. High water content decreased the heating value, and the viscosity and in combustion decreased NO, and increased solid emission in the exhaust gas. A high content of solids in the oil yielded high particle emissions in exhaust gas. The optimum viscosity range yielded good atomisation as mentioned above. The lower heating value has to be defined for adjusting steady power output.
-
DISCUSSION A summary of typical emission values for “test oils“, “typical oils“ and “water oils“ at the fuel power of about 4 MW under as identical conditions as possible is presented in Table 6. There were no principal problems in the feed or combustion of the pyrolysis oils except for pine oil 2 (Oilon 15), which caused blockages in the nozzle. Despite these problems one basic measurement was successhlly carried out for this oil by adding MeOH to the oil.
Table 6 Typical emission values for different oil grades at nominal fuel power of 4 MW. Combustion conditions: total output 3.4 - 3.8 MW, compressed air used as atomisation medium of pyrolysis oil, oil temperature 48 58 O C , regulation of primary air 50%, regulation of secondary swirl S’4.7, primar~lair swirler 60”. Particles had been measured during a longer-term test cycle.
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Test oils Typical oils OILON oil 6.6 10+11 12+13 10 0 MeOH. wt% 0 Raw material Hardwood 1 Hardwood 3 Pine 1 0 2 , vol% 3.4 3.5 3.3 - 3.4 3.3 - 3.6 10-25 1-2 2-5 c o , ppm 193-205 159- 164 108 NO,, ppm 0 0 0 NOZ,ppm Soot, Bac. 1.3 2.0 2.8 PARTICLES, g/MJ 15 100 92
-
15 10 Pine 2 3.5 2-6 110 5 3.5 53
Water oils 7+8 0 Pine 3 3.4 - 3.7 2-4 88 - 105 1 5.9 98
The reason for the nozzle blockage is unclear. There may have been some operational problems during the production of pine oil 2, which could have contributed to thls kind of peculiarity. However, the GC-MS analyses showed that pine oil 2 contained significantly more methoxy-substitutedphenols than the other oils tested. These monosubstituted fragments are particularly reactive (due to fiee position on aromatic ring) and prone to recombinatiodcondensation reactions. In pine oil 2, flaky sticky precipitation increased in the course of time, which may consist of polymerisation products of these fragments and be the reason for the nozzle blockage. This precipitation was left on the bottom of the container and the rest of the oil was used in the tests. This lund of precipitation did not appear in other pine oils. It is quite clear that the chemical composition of pyrolysis oil affects significantly the atomisation behaviour of the oil. However, more data is needed on the reactivity of various commercial pyrolysis oils before final conclusions can be drawn. The test oils were relatively easy to burn and emission values were fairly good. The combustability was significantly improved by methanol addition. After adjusting the burner, no greater problems appeared and the boiler remained clean. Neither did the nozzle clog easily. The combustion of pyrolysis oils includes two main phases: 1) vaporisation and combustion of light fraction, vaporisation of water and gasification of heavier fractions, and 2) combustion of the heaviest lignine fractions. As the initial stage requires a long residence time at a sufficiently high temperature, the front chamber improves the total combustion process. It is essential in the dimensioning of the front chamber that the flame shall not touch its surfaces in any stage, as coke is rapidly deposited on the surface and the particle content increases. If the shape of flame is well suitable in the front chamber, the particle emission is sharply reduced. The heat release rate per volume (kW/m3)of the funace should not be as high as in mineral oil combustion, as the pyrolysis oils burn slower and the flame of respective power is larger. Hence, the power output of existing mineral oil boilers reduces when pyrolysis oil is used. It might be possible to apply co-combustion in large boilers, e.g. by running, for instance, the highest burner level or some of the burners with mineral oil. The possible tar-like particles, soot and PAH compounds of pyrolysis oil flames burning on the lower level are mixed in oil flames and burn prior to convection surfaces. The atomisation viscosity should be adjusted suitably so that the size and size distribution of oil droplets are correct and their penetration into the air pattern of the burner is as desired. A burner dimensioned for oil combustion can be used or it should at least be slighly modified. However, if the best result is wanted a special burner should be designed. A too high temperature (> 60 "C) increases the risk of nozzle blockage. Particle emissions, soot and usually also nitrogen oxide emissions increase 1477
compared to a situation, where the oil temperature is lower and the viscosity is suitable. As regards emission values, a general conclusion is that the nitrogen oxide level was rather high for the test oils, while the particles emissions were small. The NO, level increased due to the reduction of water content and the heat of initial flame due to rapid combustion of added methanol. Also the burner had to be adjusted so that the flame was suitable, stable and the level of particle emissions reasonable. On the other hand, this contributed to oil vaporisation and particles gasification and hence to small particles emission. The typical pyrolysis oils were nearly invariably relatively easy to burn. The emission values were fairly good, but not as good as those for the test oils. A general conclusion fkom the emission values is that the nitrogen oxide level of typical pyrolysis oils was lower than that of the test oils, but the particles emissions respectively clearly higher. The NO, level was reduced by the higher water content and the disuse of methanol, which on the other hand resulted in higher particles emissions. The "water oils" were surprisingly easy to burn in dual fuel mode and the emission values were relatively good. The soot number rose to a higher level than that of typical oils. One reason for the high soot number may be the unproper (too low) viscosity level of these oils. As soon as the suitable burner adjustments were found, no great problems in combustion occurred and the boiler remained clean. The nozzle also remained fairly clean in all test runs. Mean NO, and particles emissions for different pyrolysis oil groups are shown in Figure 5.
Fig. 5 Mean NO, and particles emissions for different pyrolysis oil groups. It can be concluded fiom the emission values, that the nitrogen oxide level of water oils was fairly low without support oil, but grew when increasing support combustion. Support fuel increased NOx emissions by two ways: the average nitrogen content of the fuel mixture increased and the flame temperature increased yielding to an increase in thermal NOx. The particle emissions were relative high. The NO, level was reduced by the high water content and the disuse of methanol. Due to these factors, the effect on particles emissions was contrary. On the other hand, the particle content of the water-rich oils was very low, and hence, the final results was not poor at all. It should be borne in mind, that the composition of soot and particles of pyrolysis oils typically differs from that of conventional oil combustion. As there is no separate 1478
soot classification for pyrolysis oils, the Bacharac soot method of oil combustion was used for describing the tone value. The particles consisted mainly of very fine brownish material on the filter. CONCLUSIONS Different grades of pyrolysis oil were burned at 4 MW nominal power output in a test furnace of Oilon Oy 's R&D Centre. Operation, dimensions, adjustment parameters, and characteristics of pyrolysis oils were tested and compared, and emissions in different cases were measured. A general conclusion is that the results were relatively clear and the main issues were retrieved well. There were clear differences in combustibility and in particular in emissions for different oil grades. The most lmportant parametres of pyrolysis oil are viscosity, water and particles content, metha1101 addition, oil raw material, and oil age. Good and poor oils or at least difficult oils were distinguished in combustion. The burner and boiler modifications improve the combustion result but cannot help much if the oil quality is poor. When optimising .combustion conditions for the test oils, i.a., the following effects of oil grade were found: feedstock and/or pyrolysis process yields various reactivities of oil components, whch may cause blockages in the feed line, oil agelunhomogenity yields uneven combustion, methanol addition homogenises pyrolysis oil and improves combustion, solids content affects, mainly the amount of incombustibles, and an increase in water content reduces NO, emissions and increases particle and soot emissions. As regards adjustments, the following factors improved combustion and flame: clean nozzle, strong swirl, intense symmetrical flame, pressure air atomisation (compared to steam), increase of air coefficient and combustion power (having enough residence time though), suitable atomisation viscosity (abt 15 20 cSt). At the optimum adjustments of this combustion system, the mean combustion results and emission values of typical pyrolysis oils were as follows: 0 2 3.5 vol%, NO, 88 mg/MJ, CO 4.6 mg/MJ, hydrocarbons 0.1 mg/MJ, soot 2.4 Bac., and particles 86 mg/MJ. In addition to the characteristics of pyrolysis oils and burner settings, different fiunace constructions were compared in the test runs.By insulating the forepart of the furnace the mean temperature levels of flame were increased and hence combustion &roved. The main results of the combustion tests are: Pyrolysis oils can be burned relatively well in conventional f h a c e s and boilers. Boilers and oil burners may require small modifications or additions. The flame is larger and combustion takes a longer time than with mineral oils. Handling and pumping of pyrolysis oil should be performed according to exact recommendations. Support fuel is required at the start of Combustion and possibly in the combustion of pyrolysis oil of poor quality to maintain good and stable combustion. The nozzles should be kept clean and in good conditions. Extra cooling air for the nozzle could be useful during combustion. Emissions fiom combustion are in general between those from light fuel oil and the lightest heavy oil, but the particle content is higher. There are no SO, and net COzemissions.
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Quality specifications should be defmed for pyrolysis oil, including especially water and solids contents. Viscosity range is significant for good atomisation. The quality of pyrolysis oil has a strong impact on emissions. High solids content in pyrolysis oil yields high particulate emissions. Hence, solids removal fiom pyrolysis vapours or oil is highly recommended. High (above 30 wt-) water content also yields high particulate emissions. These emissions can be decreased to a certain extent by using a support fuel and optimizing the atomization viscosity. Methanol addition (max 10 wt%) homogenizes the inferior-grade oil and decreases particulate emissions. The costs for methanol addition and oil combustion in a commercial boiler are most probably lower than those of incinerating poor-quality oil in a special incineration plant for hazardous wastes. Further research is required on combustion properties of various commercial pyrolysis oils in order to identify the reasons for emission behaviour, nozzle blockages and related phenomena. ACKNOWLEDGEMENTS This work was funded by the VTT Tekes PROGAS Programme, VTT Energy, Birka Energi (Sweden) and Oilon Oy. Thanks are due to Eero Pekkola, Managing Director, and Seppo Hotti, Product Group Manager, for the leadership of the project, and to Mikko Mikkola, Timo Nironen, Hannu Sipilainen and Kauko Tuovinen for the technical assistance at Oilon. From Birka Energy, thanks are due to Bjbrn Hallgren, LarsErik Hagerstedt, and Ham Vuolutera, and from Fortum, to Steven Gust. At VTT Energy thanks are due to Jarmo Kleemola, Ilkka Isoksela, Eija Tapola, Johanna Levander, Eeva Kuoppala, Jaana Korhonen, Kaisa Lanttola, Raili Silvasti, and Paula Kayhko for their technical assisstance, and to Antero Moilanen.
REFERENCES 1.
2.
3.
4. 5.
6.
Gust, S. (1997). Combustion of pyrolysis liquids. In: Kaltschmitt, M. & Bridgwater, T. (eds.). Biomass gasification & pyrolysis. State of the art and h t w e prospects. Newbury: CPL Scientific Ltd. Pp. 498 - 503. Jay, D., Rantanen, O., Sipila, K. & Nylund, N.-0. (1995). Wood pyrolysis oil for diesel engines. Proc. 1995 Fall Technical Conference, Milwaukee, Wisconsin, 24 27 Sept. 1995. New York ASME. Sipila, K., Oasmaa, A,, Solantausta, Y., Arpiainen, V. & Nyronen, T. (1999). Perspectives for pyrolysis oil production and market in Scandinavia. In: Sipila, K. & Korhonen, M. (eds.). Power production fiom biomass III. Espoo: VIT. Pp. 277 292. (V’IT Symposium 192). Sipilli, K., Oasmaa, A., Solantausta, Y., Arpiainen, V. & Nyronen, T. (1999). Pyrolysis oil - a new product for heat and power. To be published. Oasmaa, A. & Czernik, S. (1999). Fuel oil quality of biomass pyrolysis oils state oftheartfortheendusers.Energy&Fuels,vol.13,no.4,pp. 914-921. Diebold, J., Milne, T., Czemik, S.,Oasmaa,A., Bridgwater, A., Cuevas, A., Gust, S., Hufhan, D. & Piskorz, J. (1997). Proposed specifications for various grades of pyrolysis oils. In: Bridgwater, A. & Boocock, D. (eds.). Developments in thermochemical biomass conversion. Vol. 1. London: Blackie Academic & Professional. Pp. 433 447.
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Sipila, K., Oasmaa, A., Arpiainen, V., Westerholm, M., Solantausta, Y., Angher, A., Gros, S., Nyronen, T. & Gust, S. (1996). Pyrolysis oils for power plants and boilers. In: Chartier, P. et al. (eds.). Biomass for energy and the environment. Vol. 1. IOddlington: Elsevier. Pp. 302 - 307. Solantausta, Y., Podesser, E., Beckman, D., ostman, A., Overend, R. P. (2000). IEA Bioenergy Task 22: Techno-economic assessments for bioenergy applications 1998-1999. Final report. Espoo: VTT Energy. 241 p. (VTT Research Notes 2024). Oasmaa, A., Leppimdci, E., Koponen, P., Levander, J. & Tapola, E. (1997). Physical characterisation of biomass-based pyrolysis liquids Application of standard fuel oil analyses. Espoo: VTT. 46 p. + app. 30 p. VTT Publications 306). Oasmaa, A. & Meier, D. (1999). Analysis, characterization, and test methods of fast pyrolysis liquids. In: Biomass - a growth opportunity in green energy and value-added products. Vol. 2. Kidlington: Elsevier Science. Pp. 1229 - 1234. Bridgwater, A. , Czermk, S., Diebold, J., Meier, D., Oasmaa, A., Peacocke, G., Piskorz, J. & Radlein, D. (1999). Fast pyrolysis of biomass: A handbook, Newbury: .CPL Scientific Ltd. 188 p. Bio fuel oil for power plants and boilers. Confidential Final Report. Contract JOR3-CT95-0025. Espoo: VTT Energy, 1999. Hallgren, B. (1996). Test report of Metlab Miljo AB. Skellefiehamn:Metlab Miljo AB. 17 p. (Reg. no. ALL-1668, 1996 02 08-09.) Fuleki, D. (1999). Bio-fuel system materials testing. PyNe Newsletter, March 1999, Issue 7, pp. 5 - 6.
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Transport, Handling and Storage of Biomass Derived Fast Pyrolysis Liquid G.V.C. Peacocke' and A.V. Bridgwate2 I Conversion And Resource Evaluation Ltd. 9 Myrtle House, 5 Cassowary Road, Birmingham, B20 INE, UK Bio-Energy Research Group, Chemical Engineering and Applied Chemistry, Aston University, Birmingham, B4 7ET, UK
ABSTRACT Biomass fast pyrolysis liquid is being developed for he1 and chemical applications. As these developments proceed, the liquid product is increasingly being transported by air, water, rail and road to satisfy user demands for products. This paper addresses the legislative requirements and regulations for the safe transport of this liquid. As biomass derived fast pyrolysis liquid is not on the UN approved carriage lists; its own classification has been determined from the UN manual as: UN 1993 Flammable Liquid [Fast Pyrolysis Liquid], n.o.s.,3, l"(a), 2"(a), 1
Ths classification should be used on all packages containing biomass fast pyrolysis liquid. Protocols for the labelling of packages and containers of all sizes are given with the aim of compliance with transport regulations in the EU, Canada and the USA. In conjunction with the requirements for packaging and labelling, guidance on the details to be enclosed on the transportation documents are given, with appropriate MSDS for the liquids. Guidance on the handling of fast pyrolysis liquid and its storage are given with procedures for treatment of spills. INTRODUCTION
Biomass derived fast pyrolysis liquid is now being actively produced for research, testing and evaluation purposes, for use as a chemical feedstock andor as an alternative fuel for use in boilers, engines and turbines. As fast pyrolysis technologies develop, and utilisation of the liquid increases, there will be a greater demand for the transportation of the liquid by all possible routes. To ensure that the liquid is transported in a safe and environmentally secure manner, due care and attention must be taken to ensure that the appropriate national and international regulations are met. To this end, it is llkely that fast pyrolysis liquid will be classed by the regulatory authorities as a "dangerous" or a "hazardous" substance for transportation puiposes. 1482
The UK Department of Transport Dangerous Goods Division and the UK Health and Safety Executive were consulted to discuss the classification of this liquid for transport. (1). Based on discussions with these organisations, their opinion was that fast pyrolysis liquid would be classed as a "dangerous good" or hazardous material, due to the chemical complexity and composition. The pyrolysis community may feel that the classification of fast pyrolysis liquid as "dangerous" or "hazardous" material does not reflect the true nature of the liquid. However, due to the hlghly variable chemical and physical properties, a wide spectrum of liquids must be considered. This paper addresses legislative requirements, transport, storage and handling of the liquid and the mitigation of spills. Thls paper and its recommendations applies only to biomass derived fast pyrolysis liquid, although it may well be equally applicable to other liquids such as carbonisation tars, gasification tars and other liquids produced by thermal processing of biomass or other solid materials or wastes.
INTERNATIONAL REGULATIONS ON THE TRANSPORTATION OF DANGEROUS GOODS The scope of this legislative review pertains to the transportation of fast pyrolysis liquids in the EU, USA and Canada for all modes of transport. This paper cannot cover all the national regulations in force, however, most national transportation regulation are based on, or use the UN Regulations, as described below.
UN REGULA TIONS At the United Nations level, all work related to the transport of dangerous goods is coordinated by the Economic and Social Council [ECOSOC] Committee of Experts on the Transport of Dangerous Goods, whch produces the "Recommendations on the Transport of Dangerous Goods", also called the "Orange Book" (2). These Recommendations and Regulations are addressed to all Governments for the development of their national requirements for the domestic transport of dangerous goods, and also to international organisations such as: 0
0 0
The International Maritime Organisation [IMO]; The International Civil Aviation Organisation [ICAO] and; Regional commissions such as the Economic Commission for Europe [ECE];
These are included for regulations and internationalhegional agreements or conventions governing the international transport of dangerous goods by sea, air, road, rail and d a n d waterways. The UN Recommendations addresses the following areas: 1. 2. 3. 4.
Dangerous goods most commonly carried, their identification and classification, Consignment procedures: labelling, marlung, and transport documents, Standards for packaging, test procedures, and certification, Standards for multi-modal tank-containers, test procedures and certification.
These recornmendations contain all basic provisions for the safe carriage of dangerous goods, but they may have to be supplemented by additional requirements 1483
that are applicable at national level or for international transport depending on the mode or modes of transport envisaged. These recommendations are presented in the new form of Model regulations so that they can be more easily transposed into national or international legislation (2). INTERNATIONAL E G'C
OF DANGEROUS GOODS BYAIR
There is no EU Agreement concerning the international carriage of dangerous goods by air. The ICAO issues technical instructions for the transport of dangerous goods by air (3). Regulations are also produced by the International Air Transport Association [IATA] (4) and these regulations are typically followed by cargo carriers for international transport. The IATA Dangerous Goods Regulations manual is based on the ICAO Technical Instructions. For example, the Civil Aviation Authority [CAA] in the UK has produced a guide for the transport of dangerous goods based on the IATA regulations (5). Dangerous goods can be transported safely by air transport provided certain principles are strictly followed. Air transport incorporates additional operational requirements that provide a harmonised system for airlines to accept and transport dangerous goods safely and efficiently. Users of the IATA Dangerous Goods Regulations are assured that they are meeting all legal requirements for shipping dangerous goods by air internationally and its use is strongly recommended. The IATA Regulations include a detailed list of individual articles and substances specifying the United Nations classification of each article or substance and their acceptability for air transport as well as the conditions for their transport. INTERNATIONAL CARRIAGE OF DANGEROUS GOODS BY WATER
As noted above the IMO is responsible for the provision of guidance on the safe transport of dangerous goods on water. The International Maritime Dangerous Goods [IMDG] Code (6) is accepted as an international guide to the transport of dangerous goods by sea and is recommended to governments for adoption or for use as the basis for national regulations. It is intended for use not only by the mariner but also by all those involved in industries and services connected with shipping and contains advice on terminology, packaging, labelling, stowage, segregation, handling and emergency response. INTERNATIONAL CARRL-IGEOFDANGEROUS GOODS BYRAIL
The transport of dangerous goods by rail is serviced by the Intergovernmental Organisation for International Carriage by Rail [OTIF]; it is responsible for ensuring harmonisation between ADR [Regulations concerning the i n t e ~ t i o n a lcarriage of dangerous goods by road - see next section], RID [Regulations concerning the international carriage of dangerous goods by rail - see next section] and ADN [Regulations concerning the international carriage of dangerous goods by water - see next section]. EU TRANSPORT REGULATIONS
EC subsidiary bodies deal with the transport of dangerous goods. These bodies are 1484
subsidiary bodies of the Inland Transport Committee, and therefore they are concerned only with inland transport, i.e. road, rail and inland waterway. These bodies are part of The Working Party on the Transport of Dangerous Goods [known as WP. 151, which is responsible for:
0
0
The European Agreement concerning the International Carriage of Dangerous Goods by Road [ADR] an& The European Provisions concerning the International Carriage of Dangerous Goods by Inland Waterways [ADN] and; The Joint Meeting of the Working Party on the Transport of Dangerous Goods and the RID Safety Committee also called the RID/ADR/ADN Joint Meeting.
The RID/ADR/ADN Joint Meeting is serviced jointly by the ECE secretariat and the secretariat of the Intergovernmental Organisation for International Carriage by Rail [OTIF]; responsible for ensuring harmonisation between ADR, RID [Regulations concerning the international carriage of dangerous goods by rail) and ADN. European Agreement concerning the International Carriage of Dangerous Goods by Road [ADR] and Rail
ADR is based on the UN Recommendations as regards the listing and classification of dangerous goods, their marking and labelling and packaging standards (7), but it also contains much more detailed provisions as regards: 1. 2. 3. 4.
The types of packaging which may be used, The consignment procedures, Transport equipment [vehicle to be used, vehicle construction and equipment], Transport operation [training of drivers, supervision, emergency procedures, loading and unloading, placarding of vehicles].
The ADR is intended primarily to increase the safety of hternational transport by road, but it is also an important trade facilitation instrument. Except for dangerous goods which are totally prohibited for carriage, and except when carriage is regulated or prohibited for reasons other than safety, the international carriage of dangerous goods by road is authorised by ADR on the territory of Contracting Parties provided that the conditions laid down in annexes A and B are complied with. There are at present 34 Contracting Parties to ADR and RID, including all of EU, USA and Canada. The guidance within RID provides similar guidance to ADR (8). European Agreement concerning the International Carriage of Dangerous Goods by Inland Waterways[ADN] (9)
The status of the European Provisions for the International Carriage of Dangerous Goods by Inland Waterways is different from that of ADR as ADN is only a recommendation directed to Governments for their national regulations and to river commissions for regulating the international carriage of dangerous goods on specific inland waterways under their responsibility. One well-known example of such regulations is the "Regulations for the Carriage of Dangerous Substances on the Rhine [ADNR]" developed by the Central Commission for the Navigation of the Rhine 1485
[CCNR]. A draft European agreement concerning the international carriage of dangerous goods by inland waterways ["ADN" agreement] has recently been completed under the joint auspices of the UNECE and the CCNR (9).
USA AND CANADA REGULATIONS USA Regulations on the Transport of Hazardous Materials The United States Department of Transportation [USDOT] regulates the transportation of Hazardous Materials to, from and through the United States. Title 49, Code of Federal Regulations, Subtitle B "Other Regulations Relating to Transportation" Parts 100 - 199 set forth the standards for Hazardous Materials transportation, commonly known as "49 CFR" and these are readily available (10). There are significant differences between 49 CFR and other international regulations such as ICAO, IATA and IMO, despite efforts to reduce these differences and in addition, the US Regulations change regularly. All shipments of Hazardous Materials to, from and through the United States must comply with 49CFR and other regulations. These regulations are administered by the Research and Special Programs Administration of the Department of Transportation [RSPA] and enforced on a federal level by the Federal Aviation Administration [FAA] for air transport; the Federal Highway Administration [FHWA] for ground transportation and the United States Coast Guard [USCG] for water transportation. Enforcement on a local level is usually the State Police or Highway Patrol. Exporters, importers, shipping lines, airlines, forwarders and couriers have to comply with 49 CFR, and failure can lead to delays in transport or in the case of non-compliance to potentially massive fines. 49 CFR Part 171.11 allows the use of ICAO Technical Instructions to be used for air shipments as long as certain additional requirements set forth in 49 CFR are also complied with. 49 CFR Part 171.12 allows the use of the International Maritime Organization's International Maritime Dangerous Goods Code for import and export shipments by vessel as long as certain additional requirements set forth in 49 CFR are also complied with.
Canadian Regulations on the Transport of Hazardous Materials Transport Canada is responsible for the transportation of dangerous goods in Canada legislated by the Transportation of Dangerous Goods [TDG] Regulations (1 1). These regulations are based on the UN Regulations and are very similar in content. A detailed manifest of all shipments is required to allow easy tracking of shipments until they arrive at their destination or exit the country. An emergency response plan must also be submitted to Transport Canada and a Summary of Emergency Response Plan; the reference number of the Emergency Response Plan filed with Transport Canada and a 24 hour telephone number must be included on the transportation documents with the consignment.
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PROPERTIES OF FAST PYROLYSIS LIQUID The nature of fast pyrolysis liquid means that there is no "generic" analysis to cover the wide spectrum of liquids producible from biomass. In the UN or EU regulations, there is no classification for fast pyrolysis liquid or its derivatives, fractions or by-products. As noted in the introduction, a submission to the EC to have fast pyrolysis liquid listed in ADR, RID and ADN could take 2 to 3 years, therefore a self assessment of the substance was made, using the methods described in the UN Manual (2). It is also likely that the outcome of a submission would be that fast pyrolysis liquid would be classed as a "dangerous substance". The assessment of a substance as a dangerous good considers the chemical and physical properties of fast pyrolysis liquid and these are summarked below.
PHYSICAL PROPERTIES OF FAST PYROL YSIS LIQUID Specific physical properties, if not known, can be determined to UN test methods (1 2). As fast pyrolysis liquid is not listed as a substance in its own right in ADR or the UN Regulations, it can be categorisd with a not otherwise specified [n.o.s.] classification. The relevant physical properties of the fast pyrolysis liquid used are given in Table 1, based upon average values in the literature. Table 1. Applicable hysical properties offast pyrolysis liquid (13) Physical moper& Moisture content pH -2 Specific gravity Viscosity [cp @ 40°C] Kinematic viscosity [cSt] Flash point ["C] Pour point ["C]
Fast wrolvsis -20% -1.2 -50
20-1000@ 25°C 50-61 -23
From these properties, fast pyrc-jsis liquid is a Class 3 substance - Flammable Liquid. The exact specification is then related to its chemical composition to determine its level of hazard, as described below. CHEMICAL COMPOSITION OF BIOMASS FAST PYROLYSIS LIQUID
There are numerous references that include chemical analyses of fast pyrolysis liquids from a variety of sources, including slow pyrolysis tar, fast pyrolysis liquid and fractions thereof. As fast pyrolysis liquid may be raw, treated, filtered, catalytically derived, upgraded and dependent on the process parameters, a "worst case" must be taken for the liquid composition, i.e. high variability. As noted, fast pyrolysis liquid falls within Class 3, based on its physical properties, however, this is further complicated by the presence of other chemicals that are in Class 6.1 - Toxic substances Iphenols, etc.] which are in concentrations above 0.1 wt%. A typical analysis of several fast pyrolysis liquids from different laboratories is given in Table 2. The Aston and IWC liquids were produced in a fluid bed and the BTG liquid was producedin the rotating cone reactor, all using wood feedstocks. 1487
Table 2 GC-MS Analysis of Fast Pyrolysis Liquids (1 4) Chemical
Aston
BTG
IWC
(SH)-Furan-Zone 2,4- and 2,5-Dimethyl phenol 2.5-DimethoxytetrahydroahydrofUran (cis) 2-Furaldehyde 2-Furfiuyl alcohol 2-Hydroxy- 1-methyl-1-cyclopentene-3-one 4-Ally- and 4-Propyl syringol 4-Ethyl guaiacol 4-Methyl guaiacol 4-Methyl syringol 4-Vinyl guaiacol
0.63 0.04 0.12
0.34 0.07
0.61 0.37 0.47
5-Hydroxymethyl-2-firaldehyde Acetic acid Acetoguaiacone Acetol Eugenol Guaiacol Homovanillin Hydroxyacetaldehyde Isoeugenol (cis) Isoeugenol (trans) Levoglucosan m-Cresol o-Cresol p-Cresol Phenol Syringol Vanillin Water
-
0.22
0.35 0.01 0.13
0.12 0.73
0.08 0.30
0.05 0.37 2.65 0.16 5.78 0.18 0.44 0.16 10.40 0.2 1 0.55 4.47 0.26 0.04 0.02 0.05
0.03 0.28 2.56 0.17 3.57 0.13 0.20 0.13 10.89 0.13 0.20 4.46 0.08 0.08 0.08 0.13
0.21 21.4
0.29
-
18.6
0.15 0.19 0.07 0.13 0.27 0.05 0.00 4.23 0.07 3.54 0.05 0.14 0.08 7.07 0.06 0.27 3.20
0.03 0.29 0.09 10.0
UN CODE FOR FAST PYROLYSIS LIQUID
To this end, the following generic classification is proposed based on the guidance in the UN Manual and the ADR for basic labelling purposes. The nearest UN n.0.s. classification that can be used [withthe addition of "Fast Pyrolysis Liquid"] is: UN 1993 Flammable liquid, [Fast Pyrolysis Liquid], n.o.s., 3, lo(a), 2"(a), 1
Only two chemical groupings need to be indicated for the components comprising the most significant risk. This classification is an interim classification until an application is made to the European Union for the inclusion of fast pyrolysis liquid on the dangerous goods list. This classification should be used on all labels and for all sizes of shipments, in particular on the transportation documents and the MSDS. For tank containers, and bulk shipments, additional placarding is required and this is discussed below.
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PACKAGING OF FAST PYROLYSIS LIQUID The most critical aspect of transport of dangerous goods is packaging. Appendix I1 gives the fuller details of packaging codes and weight restrictions for specific UN approved packaging types. Fast pyrolysis liquid can be shipped from small samples of the order of a few mg to tonne quantities, in a variety of receptacles [single package or combination packaging] and for different purposes. RECEPTACLE REQUIREMENTS TO COMPLY WITH UN REGULATIONS
Receptacles that are acceptable for international transport must meet the specifications of packaging described in the UN regulations (2). Packaging relates to samples of liquids of all sizes. However, depending on the classification of the dangerous goods, there are limitations to the quantities, which may be shipped per package. Packages may also be single, e.g. drums, or combination packages, e.g. plastic bottles inside a cardboard box. In summary, the physical requirements for packaging are in Table 3. There are also particular limits for Class 3 substances, depending on the package, materials of construction and type. The wide ranges of combinations are not discussed here, but a summary of the limits for innex packages is given in Table 4. Table 3. Minimum Package requirements for Fast pyrolysis liquid Packing group Receptacle required minimum test pressure Degree of filling of receptacle [at 15"CI: Hazard symbols
1 or "X" 250kPa g 90%
see below
Table 4. Limitations on inner packages Max. uermissible capacitv rl1 T w e of inner packaging 5 Glass, porcelain or stoneware packaging Plastic packaging 30 Metal packaging 40 1 Other types of small packaging, e.g. tubes PACKAGING REQUIREMENTS
Fast pyrolysis liquid can be transported in varying quantities, from grams to tonnes. There is no stated maximum for the shipment of Class 3 liquids; however, the labelling and packaging requirements will vary, depending on the size of package. Guidance is given on the package specification for a range of shipment sizes. Very small quantities [up to 1 litre]
For small samples, it is recommended that polypropylene [e.g. NalgeneTM]bottles be used with a cap insert inside the neck. This type of plastic packaging is extremely resilient to compression and damage. Glass bottles should be avoided where possible. All very small quantities should be shipped as a combination package an inner package in a UN approved box containng a suitable fabric adsorbent.
-
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Small quantities11-10 litres] For larger quantities, individual 1 litre containers should be used either with a cap insert, or small plastic drums with non-removable heads. A supplier of such packaging can provide a suitable receptacle and it is recommended that for quantities of less than 10 litres, an outer package is used, e.g. a cardboard box filled with adsorbent or a wooden box. If a 10 litre plastic dnun were used, it would be preferable to place it for shipment in an outer package, e.g. a steel dnun or wooden box filled with adsorbent. Drums can also be shipped as a single package.
Moderate quantities [I 0-450 litres [maximum 400 kgii For moderate quantities, the UN limitations mean that maximum volumes are only 30 1 for plastic packages [drums] or 40 1 for metal drums [see Table 41. Metal drums should be stainless steel, or a PTFE lined mild steel drum if used as an inner package. For single packages, e.g. dnuns,the maximum weight is 400 kg; therefore, standard drums can be used, provided they are stainless steel or polypropylene barrels [or lined steel drums]and are filled to less than 90% of maximum capacity. In addition to the packaging types noted above, the other important packaging, which is occasionally used for fast pyrolysis liquid, is the Intermediate Bulk Container [IBC]. An IBC is a rigid, or flexible portable packaging, other than those specified in Appendix A S of the UN guide (2). According to the requirements of ADR, IBCs are not to be used for Packing Group 1 liquids.
Large quantities [more than 400 kg] Large samples need to be transported in larger containers or tanks. Containers are specifically defined in the ADR regulations (2). The additional requirement for tadcontainers is the use of a placard on road containers, displayed the appropriate UN code of 1993 on the bottom and 33X on the top [see next section]. Containers and tanks will be the preferred method with time for larger quantities for land transport. Again, acid resistant containers and tanks are required.
LABELLING OF PACKAGES Marking Each package shall be clearly marked with the substance identification numbef of the goods to be entered in the transport document, preceded by the letters "UN". For fast pyrolysis liquid, this classification is 1993.
Danger labels Packages containing substances or articles of this class shall bear a label conforming to model No. 3 [Class 3 - Flammable liquids] and shall in addition bear a label conforming to model No. 6.1 [Class 6 - Toxic substances] as shown in Figure 1. Packages containing receptacles, the closures of which are not visible from the outside, and packages containing vented receptacles with or without outer packaging shall in addition bear on two opposite sides a label conforming to model 11- see Figure 1. 1490
There is a requirement for this label, due to the presence of phenols in the liquid. Labels No. 3 and No. 6.1 shall be diamond shaped and measure at least 100 x 100 mm. They have a line of the same colour as the symbol appearing on the label 5 mm inside the edge and running parallel to it. If the size of the package so requires, the dimensions of the label may be reduced, if they remain clearly visible [see Table 51. Labels to be affixed to vehicles, to tanks of more than 3 m3or to large containers shall measure not less than 250 x 250 mm.
Figure 1. Labels for fast pyrolysis liquid receptacles Any label required to be carried on a package shall be securely fixed to the package with its entire surface in contact with it and the label shall be clearly and indelibly printed. The colour and nature of the marking shall be such that the symbol [if any] and wording stand out from the background to be readily noticeable and the wording shall be of such a size and spacing as to be easily read. The package shall be so labelled that the particulars can be read horizontally when the package is set down normally. The dimensions of the labels required for the package are given in Table 5 . For tank containers, as defined above, a placard is required. The top code for the placard is: 33X, the bottom code is 1993. Large tanks are fitted with replaceable codes. The placard dimensions are typically a minimum of 30 cm high by 40 cm wide, numerals to be a minimum of 10 cm hgh, black on an orange background.
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Table 5 . Label Requirementsfor Packages CaDacitv of Package Not exceeding 3 litres if possible Exceeding 3 litres but not exceeding 50 litres Exceeding 50 litres but not 500 litres Exceeding 500 litres
Dimensions of label at least 52 x 74 mm at least 74 x 105 m at least 105 x 148 mm at least 148 x 210 mm
M m D AND EMPTY PACUGING Mixed packaging
Fast pyrolysis liquid should not be shipped with other dangerous goods. Up to 5 litres [inner packaging] may be shipped with goods not subject to the provisions of ADR, provided they do not react dangerously with each other. Fast pyrolysis liquid shall not be packed together with substances and articles of Classes 1 and 5.2 [explosives substances and articles and organic peroxides respectively, other than hardeners and compound systems] and material of class 7 [radioactive material]. Only 0.5 litres of fast pyrolysis liquid per inner packaging and 1 litre per package, which are classed under (a - highly hazardous), may be shipped with up to 5 litres of Class 3 substance not exceeding 5 litres if they classified under (b- moderately hazardous) or (c- low hazard). This applies if mixed packaging is also permitted for substances or articles of these classes, andor with goods that are not subject to the provisions of ADR, provided they do not react dangerously with each other. Dangerous reactions are classed as: 1. 2. 3. 4.
Combustion andor giving off considerable heat, Emission of flammable andor toxic gases, Formation of corrosive liquids, Formation of unstable compounds.
If wooden or fibreboard boxes are used, a package shall not weigh more than 100 kg relating to the inclusion of adsorbent packaging and the provision of a leak proof outer package and that all individual packages are clearly singly labelled (2)].
Empty packaging In the case of empty tank vehicles, empty demountable tanks and empty tank containers, uncleaned, this description shall be completed by adding the words, "Last load", together with the name and item number of the goods last loaded, e.g.: Last load 1993 Flammable liquid [Fast Pyrolysis Liquid], l"(a) HANDLING AND STORAGE OF FAST PYROLYSIS LIQUID HANDLING OF FAST PYROLYSIS LIQUID
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Generally, handling and storage activities are a common feature of the majority of workplaces. They are also one of the principal causes of death and injury due to this interface between people and the wide range of materials handled. This section sets out the measures necessary on the part of employers, employees, manufacturers, designers, importers and suppliers of fast pyrolysis liquid used at work to ensure safety and the minimisation of risks to health in connection with the use, handling, storage and transport of fast pyrolysis liquid. The guidance in this section is based on the UK legislation (15) and Guidance from the Environment Agency, UK on the storage of hazardous materials. Handling and storage covers a broad range of areas, including: 1. 2.
3. 4. 5.
6. 7. 8.
The use of fixed and mobile handling equipment; Manual handling operations; Design of the workplace; The provision of a suitable working environment; Specific requirements for the handling and storage of identified hazardous substances; Controls on the use of hazardous materials; Specific requirements for labelling of hazardous substances; and The selection, provision and use of personal protective equipment.
The law on handling and storage is diverse. For example in the UK these range from the more general requirements such as the Health and Safety at Work Act [HSWA] to the specific requirements of the Control of Substances Hazardous to Health [COSHH] Regulations 1994 and the Highly Flammable Liquids and Liquefied Petroleum Gases Regulations 1972. Each country has its own legislation and national guidance for the handling of goods and these should be used where appropriate. The handling and storage of materials has, in many cases, great potential for pollution incidents, particularly in the case of hazardous materials that may be discharged by natural seepage to water and land resulting in groundwater pollution in particular. Within the EU, this issue is tasked within the new Integrated Pollution Prevention and Control [IPPC] regulations that came into force on 1st August 2000. Guidance from the Management of Health and Safety at Work Regulations [UK], 1992, proposes that the following activities should be instigated in the case of the handling and storage of hazardous materials: 1. 2.
Risk assessments;
Implementation of management systems for the effective planning, organising, controlling, monitoring and review of any preventive and protective measures arising from a risk assessment; 3. Appointment of competent persons; 4. Establishment of emergency procedures to be followed in the event of serious or imminent danger; 5. Provision of comprehensible and relevant information; 6. Consideration of human capability; 7. Provision of health and safety training; and; 8. Consultation with safety representatives.
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STORAGE OF HAZARDOUS SUBSTANCES Prior to storing and handling fast pyrolysis liquid, it is essential to consult sources of hazard data, typically the MSDS or other available sources (16, 17). The chemical compatibility of hazardous materials must be given particular consideration. Potentially reactive material must be stored separately [mixing may occur due to spillage, leakage or accident e.g. during a fie]. The following precautions are necessary to ensure the safe handling and storage of dangerous goods andor chemical substances with fast pyrolysis liquid: Meticulous standards of housekeeping should be maintained at a l l times; Smoking and the consumption of food or drink should be prohibited in any area in which substances are used or stored; 3. Staff must be reminded regularly of the need for good personal hygiene, in particular washing of hands after handling chemical substances; 4. The minimum quantities only should be stored in the working area; extra bulk storage may be required separately and well away fiom the work area; 5. Containers and transfer containers should be clearly and accurately marked; 6. Chemical substances should always be handled with care and carriers used for Winchester and other large containers; 7. Fume cupboards should operate with a minimum face velocity of approximately 0.4 d s e c when measured with the sash opening set at 300 mm maximutn, and performance should be checked fiequently in accordance with the COSHH Regulations; 8. Staff should always wear personal protective clothing and equipment e.g. eye protection, face protection, aprons, gloves, protective boots, whenever handling fast pyrolysis liquid, in particular when handling large quantities; 9. Responsibility for safe working should be identified at senior management level, and written procedures published and used in the training of staff. 1. 2.
Bulk chemical storage [drums, barrels, tanks and similar containers]
In the design and use of bulk storage facilities, the following aspects need attention: 1. 2.
The range and quantities of substances to be stored; Dependent upon (1) above, the degree of segregationby distance of a. The store fiom any other building; and b. Certain chemical substances within the store from other chemical substances stored.
Purpose-built chemical stores should be of the detached single-storey brick built type or constructed in other suitable materials, such as concrete panels, with a slopmg roof of weatherproof construction. The structure should have a notional period of fire resistance of at least one hour. Other features include [based on advice from the Environment Agency, UK]: 1.
Permanent ventilation by bigh and low level airbricks set in all elevations, except in those fonning a boundary wall; low-level airbricks should be sited above doorsill level; 1494
2.
Access doors constructed from material with at least one hour notional period of fire resistance; doorways should be large enough to provide access for fork lift trucks, with ramps on each side of the door sill [also to contain any internal spillage]; separate pedestrian access, which also serves as a secondary means of escape, should be provided; 3. An impervious chemical-resistant finish to walls, floors and other surfaces; 4. Artificial lighting by sealed bulkhead or fluorescent fittings, to provide an overall luminance level of 300 lux; 5 . Provision of adequate space, with physical separation and containment for incompatible substances, each area to be marked with the permitted contents, the hazards and the necessary precautions, and incorporating an area for the storage of empty containers; 6 . Fire separation of individual areas sufficient to prevent fire spreading; 7. Provision of the following equipment in a protected area outside the store: 0 Fire appliances (drypowder and/or foam extinguishers); 0 Fixed hose reel appliance; 0 Emergency shower and eyewash station with water heating facility to prevent freezing; 0 Personal protective equipment i.e. safety helmet with visor, impervious gloves, disposable chemical-resistant overall, with storage facilities for same; and 0 Respirator and breathing apparatus in a marked enclosure; 8. A total prohibition on the use of naked flames and smoking, appropriate warning signs should be displayed; 9. A prohibition on the use of the store for storage of other items or for any other purpose; and 10. Provision of racking or pallets to enable goods to be stored clear of the floor.
External drum storage Drums, barrels, carboys and other similar containers for fast pyrolysis liquid should be stored in the external air on an impervious and durable surface, which is in excess of 4 m to any risk area, bund or open boundary. The area should be protected by a bund wall, dished or ramped to contain spillages, with the walls and floor impervious to the materials stored. The bunded area should contain no drains or valves. Vehicular access to such areas should be protected by a ramp or a channel ensuring that the ramp itself does not cause regular spillages. Generally, no container should be stored within 2 m of any window, escape route or door. Much will depend on the nature of the substances stored and the design of the storage area. Where a storage area is constructed with fire resistant walls, these distances can be reduced. Ensure that overflow pipes on all tanks discharge within the bunded ar5a. Any tank situated on a roof may drain to the surface water system via the guttering therefore roof storage should be avoided. Flammable liquids should be stored in a purpose-built external flammable materials store and not in a warehouse. Much will depend upon the quantities to be stored. Small quantities should be stored in a lockable metal cupboard, suitably marked. Dnuns should be stored within a bunded area to contain any spillages., Vehicular access to such areas should be protected by a ramp or a channel ensuring that the ramp itself does not cause regular spillages. Automatic cut-offs 1495
should be provided on all delivery pipes to prevent spillage due to overfilling.
TREATMENT OF SPILLS Spills can range in size from grams to thousands of litres. The prevention of spills and the problems associated with spillages can be minimised by good handling, storage practices, and other preventative measures. Spillages of hazardous substances can arise because of poor storage systems, the use of unsuitable or defective containers, during refilling of tanks and other containers, or because of human error. Procedures for dealing with both small and large spillages should be outlined in the Material Safety Data Sheet [MSDS]. In most cases, small spillages can be dealt with immediately by absorption in sand, sawdust or proprietary absorbent granules for subsequent disposal to a waste container. Large spillages will need a considerable degree of attention to ensure, in particular, that substances do not enter a drainage system or natural watercourse in concentrated form. This may entail containment of the spillage using drain cones, sand bags, polythene sheeting and a range of other materials, so that it can eventually be pumped into a disposal container. Spillage of hazardous materials should never be washed to surface water drains. An absorbent material used to clean up a spillage of a pyrolysis liquid may have to be disposed of as special waste. The potential for water and ground pollution because of spillages very much depends upon the handling systems operated on site. The following recommendations with regard to materials handling are made: All loading and discharge points should be designated, marked and isolated from the surface water drainage system. 2. Routes of transfer for all materials should be identified and the complete route should be protected against spillages to the surface water system. 3. Underground pipework should be avoided where possible, as faults are more difficult to detect and can lead to groundwater contamination. 4. Manual handling should be avoided where possible to reduce the risk of human error and accidents. 5 . Yard areas used for materials handling or materials handling processes must be isolated from the surface water drainage system by bunding. Roofing over such an area is an advantage to prevent the accumulation of rainwater, fire regulations permitting. 6. Appropriate containers should always be used for different materials. They should be sturdy, in good condition, clearly labelled and not liable to leak. 7. The necessity for materials handling and transfer should be minimised to reduce risks. 1.
To handle spills, the following actions are proposed
Small Ouantities TuD to 1 litre] 1. Wear rubber gloves and suitable eye and face protection. 2. Cover contaminated area with sawdust, or other suitable inert adsorbent, e.g. vermiculite or montmorillonite, 3. Collect contaminated absorbent and place in closed container, 1496
4.
Transport to approved landfill or incinerator for disposal.
Large Ouantities [above 1 litre] 1. Evacuate area. 2. Wear rubber boots, rubber gloves, suitable eye/face protection and NIOSWMSHA approved respirator. 3. Cover contaminated area with sawdust or vermiculite. Take up sawdust or vermiculite and place in closed container. Transport to approved landfill or incinerator. For large spills, a spill kit is advised to contain the spill and prevent its incursion to local watercourses. The recovery of fast pyrolysis liquid as an adsorbed waste may mean that it must be treated as "special waste" for disposal. Where possible, contaminated adsorbent should be recovered and stored in sealed containers for subsequent disposal.
CONCLUSIONS Fast pyrolysis liquid may be classed as a Class 3 Flammable liquid for the purposes of transport by any mode. The use of the UN manual for the assessment of fast pyrolysis liquid has allowed a generic classification to be made to ensure that liquid that is shipped in UN approved packaging, or packaging conforming to UN requirements is meeting the relevant internationalregulations. The recommended coding is:
UN 1993 Flammable Liquid [Fast Pyrolysis Liquid], nos., 3, l"(a), 2"(a), 1 The use of the UN code will allow samples to be shipped in a manner complying with the regulations currently in force in the EU, Canada and the USA. The use of an approved supplier of UN packaging will ensure that samples are shipped in the correct packaging and facilitate transport, which is more rapid, with the required information for the transportation document and the provision of MSDS.
RECOMMENDATIONS The following recommendations are made: 1.
2.
3. 4.
Data on fast pyrolysis liquid should be provided to the UK DETR - Dangerous Goods Branch to permit an application to the EU for the listing of fast pyrolysis liquid. This automatically provides access to the analogous approvals bodies in 34 countries. Data is required on the toxicity of fast pyrolysis liquid to ensure that MSDS can provide sufficient mformation for handlers and transporters of the liquids, The relative quantities of adsorbents required to deal with spills should be assessed, Work should be carried out to understand how to remediate spill areas and assess environmental consequences.
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ACKNOWLEDGEMENTS
The authors wish to thank the EC and IEA Bio-energy Funded Pyrolysis Network (PyNe) for providing the funding to enable this work to be carried out. The full report from which this paper is taken will be published by the PyNe Network. GLOSSARY 49 CFR
ADN ADNR ADR CAA CCNR COSHH (UK) DETR (UK) ECE ECOSOC FAA (USA) FHWA (USA) HSWA (UK) IATA IBC ICAO IMDG IMO MSDS MSHA NIOSH OTIF PTFE RID RID/ADR/ADN
RSPA (USA) TDG (Canada) USCG USDOT
United States Regulations Relating to standards for Hazardous Materials transportation The European Provisions concerning the International Carriage of Dangerous Goods by Inland Waterways Regs. for the Carriage of Dangerous Substances on the Rhine EU Agreement on the Carriage of Dangerous Goods by Road Civil Aviation Authority Central Commission for the Navigation of the Rhine Control of Substances Hazardous to Health Regulations 1994 Department of the Environment, Transport and Regions, UK Economic Commission for Europe Economic and Social Council Committee of Experts Federal Aviation Administration Federal Highway Administration Health and Safety at Work Act International Air Transport Association Intermediate Bulk Container International Civil Aviation Organisation International Maritime Dangerous Goods International Maritime Organisation Material Safety Data Sheets Mine Safety and Health Administration, USA National Institute of Occupational and Safety Hazards, USA Organisation for International Carriage by Rail Polytetrafluoroethene EU Agreement on the Carriage of Dangerous Goods by Rail The Joint Meeting of the Working Party on the Transport of Dangerous Goods and the RID Safety Committee Research and Special Programs Administration of the Department of Transportation, USA Transportation of Dangerous Goods Regulations, Canada United States Coast Guard United States Department of Transportation
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REFERENCES 1. Approved Supply List [4th Edition] - Information Approved for the Classification and Labelling of Substances and Preparations Dangerous for Supply, HM Stationery Office, 1998. 2. UN Recommendations on the Transport of Dangerous Goods Model Regulations, 1lth Edition, United Nations, New York, January 2000. 3. Technical Instructions for the Safe Transport of Dangerous Goods by Air, 19992000 Edition. Doc 9284-AN/905, ICAO. 4. Dangerous Goods Regulations CD-ROM (ref395 15-41) IATA Dangerous Goods Regulations manual, IATA, Montreal, Quebec, Canada. 5 . T e c h c a l Instructions for the Safe Transport of Dangerous Goods by Air [Doc 9284-ANI905 and supplement] from CAA, Printing and Publications Service, Cheltenham, UK. 6. International Maritime Dangerous Goods Code (IMDG Code) - 1994 [including Amendment 29-98], IMO-213E, 1999, IMO. 7. European Agreement concerning the international carriage of dangerous goods by road [ADR] and protocol of signature, United Nations, New York and Geneva, ECE/TRANS/130 (Vols 1 and 2), 1998, ISBN 92-1-139062-1. 8. Regulations concerning the international carriage of dangerous goods by rail [RID) 1999 edition [Annex 1 to Appendix B to the Convention concerning international carriage by rail [COTIF)), HM Stationery Office 1998. 9. European Provisions concerning the International Carriage of Dangerous Goods by Inland Waterway [ADN], ECE/TRANSMTP.15/148, EU, 2000. 10. See Internet site http://haanat.gov.com/ 11. The Export and Import of Hazardous Wastes Regulations and the Transportation of Dangerous Goods [TDG] Regulations, 1992, Transport Canada, Ottawa, 1992. 12. Recommendations on the Transport of Dangerous Goods - Manual of Tests and Criteria - Third revised edition, United Nations, New York, 2, January 2000. 13. Bridgwater A.V. (1996) Production of hgh-grade fuels and chemicals from catalytic pyrolysis of biomass. Catalysis Today, 29, 1-4, pp. 285-295. 14. Meier D., Oasmaa A. and Peacocke G.V.C., (1996) "Properties of Fast Pyrolysis Liquids: Status of Test Methods", In: Developments in Thermochemical Biomass Conversion, (Ed. by A.V. Bridgwater and D.G.B. Boocock), volume 1, pp 391408, Blackie Academic and Professional. 15. Health and Safety At Work Act [UK], 1974 and Management of Health and Safety at Work Regulations [UK], 1992 16. Diebold J.P. (1999) "A Review of the toxicity of biomass pyrolysis liquids formed at low temperatures", In: Fast pyrolysis of biomass: a handbook, (Ed. by A.V. Bridgwater, S. Czermk, J. Diebold, D. Meier, A. Oasmaa, C. Peacocke, J. Piskorz and D. Radlein), pp. 135-163. CPL Press Ltd, UK. 17. Chemicals (Hazard Information and Packaging for Supply) (CHIP 2) Regulations 1994, UK.
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Levoglucosenone - a Product of Catalytic Fast Pyrolysis of Cellulose G. Dobele, G. Rossinskaja, G. Telysheva Latvian State Institute of Wood Chemistry, 27 Dzerbenes St., L V-1006, Riga, Latvia D. Meier, S. Radtke, 0. Faix Federal Research Centre for Forestry and Forest Products, institute for Wood Chemistry and Chemical Technology of Wood,0-21 027, Hamburg, Germany
ABSTRACT: Five celluloses, differing by polymerization degree and crystallinity index, impregnated with phosphoric acid and thermally pretreated have been subjected to fast pyrolysis in order to investigate the formation of levoglucosenone, a dehydrated 1,6-anhydrosaccharide. It was established that the yield of levoglucosenone depends on both properties of cellulosic raw material and conditions of thermal pretreatment. Thermal pretreatment allows to change the degree of polymerization of cellulose. When the pretreatment temperature exceeds 100 "C, monophosphates of cellulose are formed, the development of intermolecular cross-linking is promoted and depolymerization of cellulose is hindered. The highest yield of levoglucosenone (29 % b.0. cellulose) was obtained by pyrolysis of microcrystalline Munktell cellulose impregnated with 3.5 % of phosphoric acid and pretreated at 100 "C. Realization of high yields of levoglucosenone from celluloses with a higher degree of polymerization requires larger amounts of phosphoric acid.
INTRODUCTION It is known that the qualitative and quantitative composition of the thermal degradation products from polysaccharides can be altered by use of different catalysts [ l ] . Inorganic acids are not well selective catalysts as they affect both the degradation'and condensation reactions in the pyrolysis process. A relationship between dehydration, degradation and condensation reactions is determined by the individual properties of the acid, the characteristics of the cellulose structure and by the pyrolysis conditions
PI.
In the presence of acid additives thermal degradation of cellulose is intensified at lower temperatures, due to the occurrence of the dehydration reactions [3]. Condensation reactions result in the decrease of volatile products at 450°C combined
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with an increase of their water portion [4]. Slow heating rates and atmospheric pressure promote the dehydration and subsequent condensation reactions [ 5 ] . Depending on the properties of the acid used as an additive, the cellulose dehydration process proceeds according to two mechanisms, i.e. via formation of intermediate carbon ions or esters. In the case of phosphoric acid dehydration is accomplished predominantly via a repeated formation and splitting of esters [ 5 ] . The cel.lulose phosphate esters are sufficiently thermostable [6], therefore cellulose condensation reactions in the low-temperature range (up to 300 "C) do not develop considerably. The phosphoric acid dehydrating action as well as its hydrolytic activity leads to formation of dehydrated bicyclic 1,6-anhydrosaccharide - levoglucosenone (more than 20% b.0. cellulose) during cellulose pyrolysis at 350 "C [7, 81. The unique chemical structure of levoglucosenone offers a variety of synthesis of biologically active products on this basis [9]. Latest reports demonstrated that levoglucosenone can be obtained from slow and fast pyrolysis of various cellulose containing raw materials impregnated with phosphoric acid [ 10-121. Slow pyrolysis of impregnated cellulose (3-7 % of phosphoric acid) promotes preferentially levoglucosenone formation. Fast pyrolysis experiments at analytical scale revealed that apart from levoglucosenone also levoglucosan is formed which is the main depolymerization product of untreated cellulose. The yields of levoglucosan and levoglucosenone in the composition of pyrolysis volatile products amount to 75 to 85%, regardless of the cellulose type, amount of phosphoric acid (3.57%) used for impregnation and pretreatment temperature ( 100- 160°C) [ 121. The goal of the present study was to investigate the effect of structural peculiarities of different types of cellulose on levoglucosenone formation under catalytic fast pyrolysis conditions. Moreover, amount of phosphoric acid and pretreatment temperature should be determined to provide maximum possible yields of levoglucosenone. EXPERIMENTAL MA TERIALS
Five cellulose samples were used: microcrystalline celluloses (Munktell, Avicell and Thermocell), cellulose from sulphate pulping process (Taircell), and recycled Kraftpulp (IFAB UKP, Sweden). The method for preparation of microcrystalline Thermocell cellulose was developed at IWCh, Latvia [ 131. Crystallinity index was calculated based on data of X-ray diffractometry [ 141, The degree of polymerization was determined by the viscosity measurements from cellulose cadoxen solutions (the relative error of the method was 5 %) [ 151. IMPREGNATION AND PRETREATMENT 87 % phosphoric acid (puriss., Fluka) was used. An aqueous solution of phosphoric acid (3.5, 5 and 7 wt.% based on dry cellulose) was mixed with cellulose (cellulose/acid solution = 15). Water was evaporated under vacuum, and the impregnated material was dried at 40 "C under the reduced pressure. The dried samples were heated at 100, 160 and 200 "C for 1 hr.
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DETERMINATION OF PHOSPHORUS The amount of phosphorus present in the sample was measured through the combustion of impregnated cellulose and the subsequent determination of phosphorus as a triple phosphorus-molybdate-vanadate complex by photometry on a Specord UV/VIS at a wave length of 400 nm [ 161. For the determination of phosphorus chemically bound to cellulose, impregnated samples were washed with water with a modulus of 1:200, then dried under reduced pressure at 40 "C. The washed sample (20-100 mg) was suspended in a potassium chloride solution (0.75 g of KCl in 25 ml of water), and was titrated potentiometrically with 0.025 N NaOH. The degree of substitution for phosphoric acid esters (mono-, diand tri-esters) was determined by titration curve differentiation [ 171.
ANALYTICAL PYROLYSIS ( 0 - G C )
A CDS Pyroprobe 100 combined with a gas chromatograph (CP 9000) was used to pyrolyze and analyze a 70 mg sample. Pyrolysis temperature was set to 500 "C. Heating rate was 600 "C" and pyrolysis time was 10 s. A DB 1701 column (60 mx0.25 mm, 0.25 mm film thickness) was used for the separation of monomeric volatile compounds. Further details of the procedure can be found elsewhere [ 181. External calibration was used for quantification of levoglucosenone.
RESULTS AND DISCUSSION CELLULOSE CHANGES DURING IMPREGNATION PRETREATMENT Effect of interaction of acid and cellulose at the stages of impregnation and thermal pretreatment depends on the cellulose properties. The celluloses under study have different ratios of ordered and amorphous regions. They differ also by their degree of polymerization (Table 1) and hydrophilic properties. The presence of phosphoric acid affects the system of the cellulose hydrogen bonds, the crystallinity index and leads to the formation of esters [ 1 1,12, 191.
Table I Characteristics of cellulose supramolecular structure and degree of polymerization (DP). Cellulose
Munktell Avicell Thermocell Taircell Kraft-pulp
Crystallinity index untreated 7 % H,PO,, 160 "C 0.86 0.85 0.81 0.78 0.72 0.7 1 0.78 0.76 0.72 0.59
DP
200 220 300 1040 2450
To evaluate the cellulose reactivity with respect to the esterification reactions, the amount of chemically bound phosphorus in impregnated materials was determined and related to the amount of acid and pretreatment temperature. Two celluloses with a
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different crystallinity index and DP were selected for comparison: microcrystalline Munktell cellulose and Taircell cellulose from sulphate pulping process. The data presented in Table 2 demonstrate that pretreatment of impregnated celluloses at 100 "C does not lead to chemically bound phosphorous and formation of esters, i.e. phosphoric acid was completely washed off from cellulose. However, phosphoric acid esters as monophosphates could be detected in the impregnated cellulose samples after heating them at 160 and 200 "C.
Table 2 Changes in the amount of chemically bound phosphorus and cellulose DP depending on the quantity of acid and pretreatment temperature. WO,,
P
(%)#
(%)#
Pretreatment temperature, ("C)
Amount of bound P
(%)#
DP
(%)
b.0. P Taircell 3.5 3.5 3.5 5 5 5 7 7
1.2 1.2 1.2 1.7 1.7 1.7 2.5 2.5
100 160 200 100 160 200 100 160
0.35 1.16
29.2 92.5
0.61 1.50
35.8 95.9
1.02
42.9
Munktell 3.5 1.1 100 3.5 1.1 160 0.18 16.1 3.5 1.1 200 0.84 76.8 5 1.6 100 5 1.6 160 0.48 30.0 5 1.6 200 1.36 81.2 7 2.3 100 7 2.3 160 0.78 33.9 7 2.3 200 1.82 81.5 *amount of cellulose insoluble in cadoxene, b.0. cellulose, % # based on cellulose
310 265 (9*) n.d. n.d. n.d. n.d. 380 230 (41*)
190 380 (lo*) n.d. n.d. n.d. n.d. 195 320 (43*) n.d.
By increasing the pretreatment temperature and the amount of the acid the content of chemically bound phosphorus increased for both celluloses. At 200 "C the less ordered cellulose Taircell binds practically all phosphorous (92.5 to 97.2%). For the more ordered microcrystalline Munktell cellulose the amount of chemically bound phosphorus was lower compared to Taircell and reached after thermal treatment' at 200 "C 8 1 'YOof the amount introduced. Besides ester formation another effect of phosphoric acid is hydrolytic splitting of glycosidic bonds. This effect is more pronounced for celluloses with a higher DP and depends also on the amount of acid used for impregnation and the pretreatment temperature.
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After thermal pretreatment at 100 "C of Taircell cellulose, which has the highest DP amongst other celluloses under study, its DP decreased from 1040 to 3 10 and 380 after introduction of 3.5 and 7 % phosphoric acid, respectively (Tables 1 and 2). An increase of DP after introduction of larger amounts of acid (7 % vs. 3.5 %) could be explained by intermolecular linking reactions, which compete with the hydrolysis of glycosidic bonds. At similar impregnation and pretreatment conditions the DP of Munktell cellulose, having the lowest DP (200), was practically left unchanged. Increase of pretreatment temperature promotes intermolecular cross-linking reactions, which are additionally stimulated by increasing amount of phosphoric acid. Pretreatment at 160 "C resulted in a decrease of cellulose solubility in cadoxene. At 3.5 % of phosphoric acid concentration, the amount of the cadoxene-insoluble fraction for both celluloses was approximately the same, i.e. 9-10 %. Increasing the acid addition from 3.5 to 7 % gave four times higher values. (Table 2). At the same time the DP of the soluble fraction of Taircell cellulose, impregnated with 3.5 and 7% of phosphoric acid and pretreated at 160"C, decreased to 265 and 230, respectively, indicating the ongoing hydrolysis of glycoside bonds. Unlike Taircell cellulose, the DP of the soluble fraction of microcrystalline Munktell cellulose increased, to 380 and 320, respectively, i.e. in this case the intermolecular cross-linking reactions are prevailing.
PYROLYSIS OF IMPREGNATED CELLULOSE, YIELD OF 1,6ANHYDROSACCHARIDES Cellulose changes taking place during impregnation and thermal pretreatment stages find their reflection in the alteration of the composition of volatile pyrolysis products (Table 1, 2). As a result of dehydration and depolymerization reactions the yield of levoglucosan (Lg) decreased, while levoglucosenone (Lg-none) became the main pyrolysis product of cellulose impregnated with phosphoric acid (Table 3, Fig. 1). Thermal pretreatment of non-impregnated cellulose also decreased the Lg yield during pyrolysis (Table 3). The changes in the yields of 1,6-anhydrosaccarides, under the conditions of analytical pyrolysis which model the fast pyrolysis conditions, were different for each cellulose (Fig.1.). The celluloses studied can be subdivided into two groups with respect to the Lg-none yield: Avicell and Kraft pulp in one group, and Munktell, Theirnocell and Taircell in another.
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Well
Taircell 33,
20
40
60
la, 120
80
140
20
160
40
M
80
103
120
140
t, "C
t "C
20.
OJ 20
40
60
a,
IM
120
140
160
t, "C
Fig 1.
Dependence of levoglucosan (dashed line) and levoglucosenone (solid line) yields on the amount of phosphoric acid and pretreatment temperature
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160
Table 3 Yield of levoglucosan obtained by analytical pyrolysis of non-impregnated cellulose. Cellulose Munktell Avicell Thermocell Taircell Kraft
Levoglucosan yield, b.0. cellulose (%) blank 79.3 36.6 13.6 19.6 8.2
after pretreatment @ 160 "C 46.2 28.2 12.4 17.0 7.3
Despite the fact that celluloses of the first group (Avicell and Kraft pulp) differ considerably by their structure, DP and chemical composition (Table l ) , the yields of 1,6-anhydrosaccharides during pyrolysis alter in a similar way depending on the impregnation and pretreatment conditions (Fig. 1). The highest levoglucosenone yield was 16% by introducing 5% phosphoric acid without thermal treatment. Pretreatment negatively affects the formation of levoglucosenone in this case. The yield of levoglucosan slightly depends on impregnation and pretreatment conditions and changes in the range of 3 to 9 %. For the second group of celluloses (microcrystalline celluloses Munktell, Thermocell and cellulose Taircell) the highest yields of levoglucosenone were 29, 28 and 25 %, respectively. They were obtained of samples impregnated with phosphoric acid and pretreated at 100°C (Fig.1). However, the amount of the acid to obtain the highest yield of Lg-none from celluloses of this group differs for the row of above mentioned celluloses and is equal to 3.5, 5 and 7 %, respectively. The least amount of phosphoric acid is required for cellulose Munktell which is characterized by a homogeneously ordered supramolecular structure and low polymerization degree. The results obtained show that the formation of 1,6-anhydrosaccharides was affected by both the properties of the cellulose and its structural changes obtained by impregnation and thermal pretreatment. The highest yield of levoglucosenone was achieved from pyrolysis of impregnated Munktell cellulose having the highest index of crystallinity and levelling-off DP. Pyrolysis of Munktell also gave higher levoglucosan yields compared to the other celluloses. However, for the tested series of celluloses no coi relation was observed concerning the yields of levoglucosan and levoglucosenone obtained from pyrolysis of initial and impregnated celluloses. Formation of cellulose phosphates during pretreatment at temperatures higher than 100 "C inhibited levoglucosenone formation. Obviously, the main positive effect of pretreatment at 100 "C was the change of the hydrogen bonds system under the action of phosphoric acid and the decrease in DP. Celluloses with a low initial DP needed less acid to give a maximum yield of Lg-none (Table 1, Fig.1). So, in the order of celluloses Taircell (DP 1040), Thermocell (DP 300) and Munktell (DP 200) produced the highest yield of Lg-none at 7, 5, and 3 % of phosphoric acid, respectively. The yield of levoglucosan decreased with an increase of phosphoric acid. For lower acid amounts (3.5 %) the yield of LG obtained from Munktell and Taircell varied within 1 1 and 7 %. Taircell impregnated with 3.5% acid and without pretreatment gave levoglucosenone yields of up to 17 %. The amount of levoglucosan is the same as for initial cellulose, namely 19 %, which is a proof for a catalytic effect of phosphoric acid on cellulose depolymerisation. Thermocell and Kraft pulp showed an increase in the summary yield of 1,6-anhydrosaccharides after impregnation with of 3.5 % acid 1506
compared to initial samples. This indicates that phosphoric acid catalyzed thermal depolymerization of celluloses having a more disordered supramolecular structure. Probably, under fast pyrolysis conditions dehydration and depolymerization processes with levoglucosenone formation proceed first of all in more disordered regions of the cellulosic structure where phosphoric acid is concentrated during impregnation. Thermal depolymerization takes place simultaneously in more ordered regions of cellulose where the dehydrating effect of phosphoric acid is less pronounced resulting in the formation of the non-dehydrated 1,6-anhnydrosaccharide, levoglucosan. CONCLUSIONS From the results obtained, the studied celluloses could be arranged in the following order with the respect to the yield of levoglucosenone obtained by catalytic flash pyrolysis: Kraft = Avicell < Taircell< Thermocell < Munktell The highest yield of levoglucosenone (29 %) has been obtained by pyrolysis of microcrystalline Munktell cellulose, which is characterized by the highest index of crystallinity, levelling-off DP and the lowest hydrophility. In general, a high yield of levoglucosenone by catalytic pyrolysis of cellulosic raw materials can be achieved when: - the amount of phosphoric acid needed for the process of cellulose hydrolytic depolymerization correlates with DP of cellulose, i.e. the higher the DP, the higher the amount of acid needed; - thermal pretreatment should be realized in the temperature range of 40100 "C, because higher temperatures promote formation of monophosphates, which slow down the depolymerization processes.
REFERENCES Einsele U., Meier P., Herlinger H.( 1979) Zusammenhange zwischen Dehydratisierungsreaktion und Flammschutzeffekt bei Cellulosefasern. Cell. Chem. And Technol., 13,57-75. 2. Hendrix I.E., Rostic I.E., Olsen E.S. (1970) Pyrolysis and combustion of cellulose. 1. Effect of triphenylphosphate in the presence of nitrogenous bases. J . Appl. Polymer Sci., 14, 1701-1723. 3. Lewin M., Basch A . (1978) Structure, pyrolysis and flammability of cellulose. In: Flame Retardant Polymeric Materials, 2, pp. 1-41. New York- London. 4. Dobele G., Dizhbite T., Rossinskaja G., Telysheva G. (1995) Thermocatalytic destruction of cellulose. In: Cellulose and Cellulose Derivatives: Physico-chemical Aspects and Industrial Application (Ed. by. J.F.Kennedy, G.O.Phillips, & P.A.Williams) pp. 125-130, Woodhead, Cambridge, UK. 5. Jermolenko I.N., Ljubliner I.P., Gulko N.V. (1982) Element-Containing Carbon Fibrous Materials. Nauka I Technika, Minsk. 6. Jacevskaja M., Komarov V. (1979) The properties of activated carbon obtained from metalsubstituted phosphorylated wood. Kim@ Drev., 2, 91-95. 7. Rossinskaja G., Dobele G., Domburg G. (1986) Thennocatalytic transformation of cellulose and lignin in the presence of phosphoric acid. 3. Characteristics of the higher - boiling oil fraction of carbohydrates pyrolysis. Khimija Drev., 6 , 72-76. 8. Shafizadeh F., Furneaux R.H., Stevenson T.T. (1979) Some reactions of 1.
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levoglucosenone. Carbohydr. Res., 71, 169-191. Witczak Z.I. Levoglucosenone; past, present and further application (1 994) In: Levoglucosenone and Levolucosans, Chemistry and Application, pp, 3- 16, ATL Press, Mount Prospect, IL 10. Dobele G., Rossinskaja G., Telysheva G., Meier D., Faix 0. (1 999) Cellulose dehydration and depolymerisation reactions during pyrolysis in the presence of phosphoric acid. J. Anal. Appl. Pyrolysis, 49,307-3 17. 11. Dobele G., Rossinskaja G., Dizhbite T., Telysheva G., Meier D., Faix 0. (1999) Cellulose as a raw material for levoglucosenone production by catalytic pyrolysis. In: Recent Advances in Environmentally Compatible Polymers (Ed. by I. Kennedy, G.Phillips, P.Williams), Woodhead publ., Cambridge, UK (in press). 12. Dobele G., Meier D., Faix O., Radtke S., Rossinskaja G., Telysheva G. (2000) Volatile products of catalytic fast pyrolysis of cellulose. J. Anal. Appl. Pyrolysis (in press). 13. Maskavs M., Kalninsh M., Reihmane S., Laka M., Chernyavskaya S. (1999) Effect of water sorption on some mechanical parameters of composite systems based on low-density polyethylene and microcrystalline cellulose. Mechanics of Composite Materials, 35, 1, 79-90. 14. Ioelovich M., Tupureine A., Veveris G. (1989) Investigation of the crystalline structure of cellulose in plant materials. Khzmija Drev., 5 3 - 9 . 15. Bolotnikova L., Danilov S., Samsonova T. (1966) The method for the measuring of cellulose viscosity and degree of polymerization. Zhurn. Prikl. Khim., 39, 176180. 16. Uniland F., Jansen A., Tyring D., Wunsh P. (1975) Complex Compounds in Analytical Chemistry, Mir, Moskaw. 17. Koch H., Bommer H.D., Koppers I. (1982) Analytische Untersuchungen von Phosphatvernetzten Starken. Starch, 34, 1, 16-21. 18. Meier D. and Faix 0. (1992) Pyrolysis - gas chromatography - mass spectrometry. In: Method in Lignin Chemistry (Ed. by S.Y.Lin & C.W.Dence), pp. 177-199, Springer, Berlin. 19. Dobele G., Rossinskaja G., Rone B. (1996) Thermodestruction of cellulose and levoglucosenone obtaining. In: The Chemistry and Processing of Wood and Plant Fibrous Materials (Ed. by J.Kennedy, G.Phillips, & P.Williams) pp. 345-350, Woodhead, Cambridge.
9.
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Microporous sorbents produced by pyrolysis and gasification of hydrolytic lignin G.V.Plaksin, O.N.Baklanova, V.K.Duplyakin, V.A. Drozdov Institute of Catalysis of the Siberian Branch of the Russian Academy of Sciences, Omsk Department, Russia 644040, 54, Neftezavodskaya st., Omsk,Russia Tel.: +7 (8) 381-2-664411, Fax:+7 (8) 381-2-646156
ABSTRACT: Plant raw processing via acid hydrolysis is widely spread in Russia. Hydrolysis lignin is a large scale waste of this processing. In Russia hydrolysis lignin deposits mount to more than 1 million tons. Hydrolysis lignin contains 60-80% carbon and is a promising raw for production of adsorbents. We have studied the process of pyrolysis and gasification of hydrolysis lignin from the Krasnoyarsk Biochemical Plant. It has been found that as pyrolysis temperature increases from 400 to 1300 C, specific adsorption surface of pyrolysis product decreases from 297 m2/g AO 19.8 m2/g, and its micropores volume decreases from 0.17 cm3/gAO 0.01 cm3/g.The width of pores in carbonized lignin 2X, passes through an extreme depending on pyrolysis temperature. We have also studied the effect of burn-off on the texture of lignin adsorbent. As burn-off grow, adsorbent specific surface, micropores volume and average pores size increases. Chromatography was used to estimate the separation capacity of lignin adsorbent as applied to gas mixtures He-Ch. Maximum separation coefficient is found to correspond to adsorbent with pores 0.8-0.9 nm wide.
INTRODUCTION
Microporous carbon materials are widely used in adsorption processes for separating gaseous and liquid components [l]. Fossil peat and coal [2], polymers and resins [3], wood pulp and other plant raw materials [4] are widely used as raw to produce microporous carboncontainingmaterials. There is rather a broad assortment of plant raw materials used as precursors in carbon sorbent production: wood pulp, nuts shell, rice husks, corn cobs, wood bark, and that is not a complete list of carboncontaining raw materials. In Russia hydrolflc lignin (a waste of large scale wood processing) is a quite promising raw material to produce microporous carbons. Note that hydrolflc lignin dif€ers from ~ t u r a lignin l by polymerization degree and polymer structure. Lignm utilization is stimulated by the following factors:
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(1) At present Russian enterprises, working with wood pulp hydrolysis, annually produce more than 1 million tons of lignin as secondary waste [5]. A large quantity of lignin is concentrated in Siberia, but not more than 30% of lignin is efficiently used, mainly as a fuel. (2) Hydrolysis lignin has some paaicular features in comparison to other lignin obtained by various methods. It has a well developed porous capillary structure, and contains more carbon. Experiments show hydrolysis lignin to be a rather promising raw for carbon adsorbents production. Adsorbents of lignin origin are characterized by large pore volume (1.0-1.5 cm3/g) and bimodal pore size distribution with the main contributionof micro- and macro-pores [5 1.
Among other industrially evolved lignins hydrolysis lignin is the most condensed polymer owing to its production conditions. Acidic treatment at high temperature and pressure (16O-19O0C,0.8-1.0 Mlla) produces a strong three-dimensional polymer net. It contains many carbon-carbon bonds between aromatic nuclei. This spatial structure makes lignin hardly soluble in water and organic solvents. Structural forms in hydrolysis lignin may be destroyed only at pyrolysis or by very strong oxidizers [6]. In Russia many researchers focused their efforts on the development of technology for the synthesis of porous materials from hydrolytic lignin, and certain results have been achieved in the field. According to some publications, not only common sorbents production was tried to solve environmental problems [7,8], but also the synthesis of sorbents for special purposes, in particular for gas mixtures separation [91. Separation of mixture H2(He) - hydrocarbons is a well known chemical problem. This process is known to base on equilibrium sorption [lo] due to a different thermodynamic nature of adsorbate-adsorbent interactions. At present the most efficient separation of thisparticular gas mixture is achieved with uniform microporous adsorbents. Macro-pores are needed to provide a free access of gas to material bulk. This study focuses on the processes producing a microporous structure of carbon sorbents during a thermal treatment of hydrolytic lignin in inert and oxidizing media, and verifies adsorbents capability of separatmg He-hydrocarbon mixtures.
EXPERIMENTAL
Hydrolysis lignin, produced at the Krasnoyarsk Biochemical Plant, was used as raw to produce micro-porous sorbents. Hydrolytic lignin consists of lignin (61.3%), cellulose (13.3%), substances extracted by alcohol-benzene mixture (15%), ash -(3.3%), sulfur (0.4%). The summed pore volume of raw lignin was l.lcd/g, its specific surface was A B E12~m2/g [81. Prior to laboratory studies lignin was dispersed to a fraction of 0.1-0.25 mm. This powder and liquid medium were for paste preparation. Thus prepared paste was extruded through a die to produce granules 6 mm in size. Then lignin granules were dried at 105°C. Dried lignin granules were exposed to carbonization and graphitization in rotatable reactor providing reaction zone volume of 0.5 dm3.Reactor was equipped with electric heater, stabilizing reaction temperature with an accuracy of *2"C. Lignin granules were loaded into reactor at 20-25". Reactor heating to desired temperature was done with a constant rate of 1-4 'C/min. Lignin pyrolysis occurred in the inert gas atmosphere at 400-13OO0C. Pyrolysis time at adjusted temperature was constant for all samples 120
-
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min. Carbonized product gasification was performed in steam at 80OoC. Gasification time was 5-120 min. After gasification carbon material bum-off was estimated.
Analysis of texture and separation capability
Thus obtained carbon material was studies regarding its texture and adsorption propexties using Sorptomatic-1900 and Sorpty-1750 (“Fisons”, Italy). Sorptomatic1900 is an automated static vacuum device to measure full adsorption isotherms by volumetric method. Sorpty-1750 - is a static device for the express analysis of specific surface by a single point of nitrogen adsorption isotherm (in our case measurements were done at a relative pressure of nitrogen vapors PPo=0.178). The total volume of sample pores was calculated fiom nitrogen adsorption isotherms (77.4K) at P/Po=0.996, assuming a molar volume of nitrogen adsorption layer to equal 34.68 cm3/mole. Specific adsorption surface (ABET)for carbons was estimated at relative equilibrium pressures ranging within 0.05-0.33, assuming an adsorbed nitrogen molecule to occupy an area of 0.162 nm2in the filled monolayer. The Dubinin-Radushkevitch equation (TOZM theory) was applied to calculate the micropores volume. Aflinity coefficient for nitrogen was taken as 0.33, and adsorption phase density at adsorption temperatures was taken as 0.808 g/cm3 [ll]. Micropores size was estimated from adsorption energy Em defined from the DubininRadushkevitch equation using correlation: X= 10/Eo,(nm), where X is a semi-width of a slit micro-pore [121. Efficiency of H2(He)-CH,separation was estimated by chromatography. Adsorbent separation capacity was verified with a model mixture H2(He)-CH,, helium content being 60-80%. Mixture was separated on a column 3mm in diameter and 1 m long. Lignin adsorbent with 0.1-0.5 mm particles was used for the purpose. Argon is gas carrier. Flow rate of gas carrier was 40 ml/min. Argodmixture ratio was 1:40. As a criteria of separate sorbent capability was used the coefficient Kp calculated from chromatography pattern. Coefficient K,, was calculated by equation:
K, =
b Y
m1 +m2 where b is a distance between the He and CH, peaks, and ml and m2 are the peak widths.
RESULTS AND DISCUSSION
Effedof pyrolysis parameters The porous structure and specific surface of activated carbons are determined by precursor type [13] and pyrolysis parameters, i.e. temperature [14] and heating rate [ 15,161. Many papers are dedicated to the synthesis of active carbons based on lrgnincellulose materials of various types [17]. There are empirical dependencies of texture on thermal treatment parameters for carbon materials fiom various precursors of plant [13,14,18]. Models for cellulose fibers pyrolysis are suggested [ 151. Hydrolysis lignin is less studied. Major s&u&es in th~sfield were performed by the Russian researchers [19,201. It is known [21], that hydrolysis lignin differs from both native lignin and cellulose by its chemical composition and structure. In hydrolysis lignin there are many carbon1511
carbon bonds and secondary multi-nucleus aromatic structures [22]. Lignin samples, obtained at wood pulp hydrolysis contain aromatic nuclei linked to each other directly or through side chains. To some extent they may be considered as polymers with conjugated bonds. Hydrolysis lignin pyrolysis produces carbon with a yield of 40-50%, as referred to dry lignin mass, and liquid phase containing 50-55 % of phenols. Therefore, it is not correct to transfer regulations found for cellulose pyrolysis onto that of hydrolysis lignin. . With this regards we have tried to elucidate the effect of main carbonation parameters (temperature and heating rate) on micro-pores volume (Vd and size (2X), specific adsorption surface (ABET). Figure la,b gives parameters dependence on carbonation temperature (heating rate 4 "/min). As carbonization temperature increases from 400 up to 700°C at other pyrolysis parameters being constant samples specific adsorption surface and micro-pores volume reduce significantly Fig. l(a). Micro-pores size varies considerably from 1.48to 1.04 nm, revealing minimum values 2X=1.04-1.24nm (Fig. lb). At 700-900°C a sharp decrease of ABETand V, is observed, which becomes even more dramatic at a further increase of carbonation temperature up to 1300°C. ABETand V,, decrease by several times. Under these conditions pores width increases to 2.74 nm. Similar dependencies for pores volume and diameter were obtained by other researchers, studying the pyrolysis of polymer and lignocellulose materials [23,24]. However, there is no good explanationto this phenomenon yet. As a working hypothesis we may suggest the following. At 400-7OO0C two main processes occur simultaneously. One of them is a thermal destruction of threedimensional polymer structures formed at wood hydrolysis. Another process is a polycondensationof polymer cbah residues provided by the labile side propane chains and phenol rings with reactive groups. These processes produce carbon-containing layers formed chaotically. They bring no essential changes to specific surface and micro-pores volume, but cause a drastic change of micropores size. A further increase of carbonizationtemperature reconstructs the carbon-containing substance to more regular structures, decreasing the microporosity. In this case micropores size increases to 2X=1.8-2.7 nm with the growing carbonizationtemperature. Thus,for hydrolysis lignin carbonization temperature range 600-700°C is optimum regarding the smallest size of pores. However, we need further studies to confirmed the above said hypothesis and to reveal more details related to the mechanism of hydrolysis lignin pyrolysis, Since the processes of lignin carbonization have different rates, texture characteristics of materials produced depend on the rate of temperature growth (Table 1). Table I Effect of the heating rate (VP) in pyrolysis of hydrolysis lignin (T=7OO0C). VP,
Oc/min 1.o 2.2 3.3 4.0
AMT, m2/g 410 371 319 304
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Vmi,
cm3/g 0.24 0.18 0.17 0.15
2Xm 0.76 0.80
0.94 1.24
P) Ni
E
h Temperature, OC
2,6 2,4 3,O 2.8
,E
2.2:
g-
2.0-
'3
1.8-
a
1.8: 1,4 1.2 1,o
-
1J
b I
400
I
I
I
I
I
600
800
1000
1200
1400
Temperature, OC
Fig. I (a,b) Carbonization temperature influence on specific adsorption
surface, micropores volume (a) and micropores width (b).
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Speclfic surface and pore volume decrease with the growing heating rate results, and micro-pores width increases from 0.76 up to 1.24 nm. Therefore, most likely high heating rates prevail de-polymerization of hydrolytic lignin accompanied by removal of low molecular fragments. At low heating rates important become the slow processes of polycondensation and thickening of low molecular aromatic compounds required to form microstructures with the smallest pores. Effect of burn-off Gasification (activation) of lignin carbonized products, causing burn-off of various degree, was performed for a further development of microporous structure of lignin carbons. Table 2 shows specific adsorption surface (ABET), micropores volume (VA and micropores width (2X) of lignin carbon versus burn-off. Table 2 Effect of bum-off on the lignin adsorbent texture. Burn-off,% 18 39 48 57 82
:$,
ABET’
0.15 0.22 0.25 0.26 0.30 0.36
m2/g 304 444 526 526 653 749
2X,nm 1.24 0.80 0.86 0.90 0.90 1.26
Experimental data demonstrate a monotonous increase of micropores volume and specific adsorption surface with growing degree of burn-off. Average pore size at a loss of 18% decreases as compared with the nonactivated material, and then grows attaining 2X=1.26 nm. Probably, a higher content of small pores content in the material at a small degree of burn-off is caused by a partial removal of volatile and resinous substances providing pores opening. Then average pore size increases, when carbonaceous material is burnt away. Activation of lignin sorbents was compared with the results obtained by M.T.Gonzalez, et al. [16] for olive stones carbon activation. Independently of a raw material used (lignin or olive stones), at activation porous structure proceeds in a similar manner in both cases. Separation of model gas mixture He-CH4
Thus the texture of produced lignin sorbents indeed depends on thermal treatment parameters. Sorbent separation capacity on its turn depends on the texture parameters. Therefore, we tried to reveal correlation between the coefficient of He-C& separation and micropores width in then lignin adsorbent by experimental data simulation (see Fig 2). Apparently, lignin sorbents with 0.76-1.0 nm micropores show a rather high ability of He-C& (Kp)separation, which corresponds to an outlet helium content of not less than 99.9%.
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Thus, regarding the effect of thermal treatment of hydrolyhc lignin on the size of lignin carbon pores, the optimum regimes of thermal treatment produce sorbents with a pore size of 0.7-1.0 nm. A pilot batch of such microporous carbon was produced, its texture
2.5
-
1.0
-
0,s
-
9
0.6
0,s
1.0
1.2
1,)
1,6
1.8
2.0
2.2
Pore width, ZX, nm
Fig. 2 Correlationbetween the coefficient of He-Ch separation and width of micropores
parameters were estimated, and its capacity in the separation of He-Ch mixtures was tested. Test results show that this carbon has micropore size 2X = 0.78 nm, and it provides separation coefficient Kp=2.37. CONCLUSIONS
Hydrolysis lignin is a large scale waste of wood processing and also a promising raw for the synthesis of various carbon adsorbent. At present synthesis of adsorbents from hydrolysis lignin is not yet studied well enough. In the present paper we report about experimental results and discuss the revealed regulations of carbon adsorbents synthesis. The width of pores in adsorbents as a function of carbonization temperatures is found to be of extreme character. We suggest a working hypothesis explaining the pore width dependence. However, we need more experimentalresults to confirm our hypothesis. We have synthesized carbon materials from hydrolysis lignin and tested them in the separation of gaseous mixtures He-C&. The adsorbents show a helium concentration capacity of up to He no 95-99 %vol. They appear proved to be promising for application in the pressure swing adsorption (PSA) plants for H2 H He concentration. REFERENCES 1. Foley H.C. (1995) Carbogenic molecular sieves: Synthesis, properties and application.Microporous materials, 4,407.
2.
Chagger H.K., Ndaji F.N.,Sykes M.L. & Thomas K.M. (1995) Kinetics of adsorption and diffusional characteristics of carbon molecular sieves. Carbon, 33,
3.
Hatori H., Yamada Y., Shiraishi M., Nakata H. & Yoshitomi S. (1992) Carbon molecular sieve films from polimide. Carbon, 30, pp. 1-111.
1405.
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4. Braymer T.A., Coe C.G., Fa~risT.S., Gaffney T.R., Schorc J.M. & Armor J.N. (1994) Granular carbon molecular sieves. Carbon, 32, 445. 5 . Chudakov M.I. (1983) Industrial application of lignin. Moscow, Lesnaya Promishlennost. 6. Babkin V.A., Levanova V.R. & Isayeva L.V. (1994) Medicine preparation from waste hydrolysis industry. Chimiya v interesach ustoichivogo razvitia, 2,559. 7. Plachenov T.G., Akhmina E.I., Boikova G.I., Vorozhbitova L.N. & Sokolovskaya L.I. (1983) Carbon adsorbents from hydrolybc lignin Journal pricladnoi chimii, 56, 1296. 8. Schipko M.L., Baklanova O.N., Duplyakin V.K. & Kuznetsov B.N. (1996) Powdery and molded sorbents from lignite and hydrolytic lignin. Chirniya v interesach ustoichivogo razvitia, 4,467. 9. Timofeev A.F., Kolosentsev S.D., Gavrilov D.N. & Plachenov T.G. (1986) Molecular-sieve properties of carbon adsorbents on the basis of hydrolytic lignin. Journal pricladnoi chimii, 59, 1210. 10. Juntgen- H., Knoblauch K. & Harder K. (1981) Carbon molecular sieves: production from coal and application in gas separation. Fuel, 60,817. 11. Dubinin M.M. (1972)Adsorption andporosify, Moscow, VAKHZ. 12. Carrsco-Marin F., Lopez-Ramon M.V. & Moreno-Castilla C. (1993) Applicability of the Dubinin-Radushkevich equcation to CO;! adsorption on activated carbons. Langmuir, 9,2758. 13. Gonzalez J.C.,.Gonzalez M.T, Molina-Sabio M. & Rodriguez-Reinoso F. (1995) Porosity of activated carbons prepared from different lignocellulosic materials. Curbon, 33,1175. 14. Mackay D.M.&.Roberts P.V (1982) The Muence of pyrolysis conditions on yield and microporosity of lignocellulosic chars. Carbon, 20, 95. 15. Brunner P.H.&.Roberts P.V (1980) The significance of heating rate on char yield and char properties in the pyrolysis of cellulose. Carbon, 18,217. 16. Gonzalez M.T., Rodriguez-Reinoso F., Garcia A.N.& Marcilla A. (1997) C02 activation of olive stones carbonized under different experimental conditions. Carbon, 35, 159. 17. Wigmans T. (1989) Industrial aspects of production and use of activted carbons. Carbon, 27, 13. 18. Mackay D.M &.Roberts P.V (1982) The dependence of char and carbon yield lignocellulosic precursor composition. Curbon,20,87. 19. Stalugin A.B, Kondratenok B.M., Dudkin B.N., Lyubitova S.G., Mordvanyuk S .A. (1993) Some problems of carbonization of hydrolytic lignin and production of active carbons on its basis. Trudi Komi nauchnogo centra Ur 0 RAN, 129,113. 20. Ciganov E.A, Akhmina E.I. &.Galaudina V.V (1978) Formation of porous structure chars of hydrolitic lignin by pyrolysis. Chimiya drevesini, 5 ,97. 21. Gogotov A.F.&.Babkin V.A (1994) Lignin is a potential source of valuable lowmolecular compounds. Chimiya v interesach ustoichivogo razvitia, 2,507. 22. Obolenskaya A.V. (1993) Chemistry of lignin. Sanct-Peterburg, Lesotecknicheskaya Academy. 23. Verma S.K & Walker P.L. (1990) Alternation of molecular sieving properties of microporous carbons by heat treatment and carbon gasification. Carbon,28,175. 24. Hayashi J. (1999) Preparing molecular sieve carbon from palm oil shell Carbon, 37, 524.
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The Formation of Petrodiesel by the Pyrolysis of Fatty Acid Methyl Esters over Activated Alumina Boocock, D.G.B., Konar, S.K. and Glaser, G. Department of Chemical Engineering and Applied Chemistry, University of Toronto, Toronto, Ontario, Canada M5S 3E5
ABSTRQCT Fatty acid methyl esters are currently being manufactured and sold in many countries as biodiesel, a renewable petrodiesel substitute. The use of biodiesel is limited by the lowest temperature at which the particular product will pass from the fuel tank through the injectors to the combustion chambers. T h s temperature is variously represented by the cloud point (CP), the pour point (PP) and the cold filter plugging point (CFPP). Very recently in Europe, North America and elsewhere there is an interest in using waste fats and oils as the source of the fatty acids and triglycerides required to make the methyl esters. Unfortunately, many of these feedstocks contain saturated fatty acids, the methyl esters of which have higher CFPP’s than those of unsaturated acids. One strategy to address this problem is to cold-filter’ the ester, thereby removing the more solid components. This presents the problem of what to do with the ‘solid’ methyl ester fraction. It has been shown that when fatty acid methyl esters are pyrolysed over activated alumina at 400°C at a weight-hourly space velocity of 0.45, they are completely deoxygenated to form linear hydrocarbons’. These are completely compatible with petrodiesel with which they can be blended in all proportions. Because the blending of biodiesel with petrodiesel is one strategy for addressing the CFPP, the pyrolytic deoxygenation offers several options for the use of waste fats and oils as biodiesel precursors.
INTRODUCTION Liquid hels play a key role in modem lifestyles. Their liquid nature offers convenience in being transportable and easy to use. Diesel, spark-ignition and jet engines are the result of this convenience. Unfortunately, the hydrocarbons, which play the major role as liquid fuel are finite, and upon combustion contribute to the accumulation of carbon dioxide in the atmosphere. It is not surprising that alternate renewable liquid fuels are receiving attention. These include ethanol, methanol, fatty acid methyl esters and fast pyrolysis oils. All can be produced from renewable resources. Ethanol may be formed by fermentation of sugars and thereby indirectly from cellulose and starches. Methanol can be produced from carbon monoxide and
1517
hydrogen, formed from the gasification of biomass. Fast pyrolysis oils result from the fast pyrolysis of biomass and they are the least homogeneous and stable of the four liquids listed. Methyl esters are produced from vegetable oils and fats. They are the only liquids of the four groups, which can currently be substituted for a conventional fuel without any engine modification. In Europe, they are at this time made mainly fiom rape oil obtained from dedicated seed crops. In the USA, soybean would likely be the oil of choice. In Canada, either soy or canola could be used. There has been some discussion of using canola containing salt resistant genes. This could be grown on the high salt acreage currently unused in Alberta. More recently, waste fats and oils have become attractive substrates. They are not only cheaper, but are becoming even more so as their use in cattle feed diminishes. More and more of this material is becoming “waste”. In addition, more jurisdictions are banning such wastes from landfills in which case it can only be incinerated. These wastes are mostly fatty acids and triglycerides from which methyl esters can be formed under appropriate conditions. Unfortunately many of these wastes differ fiom vegetable oils in that they contain higher percentages of saturated fatty acids and triglycerides. As a result, the methyl esters produced fiom them have a higher cloud point (CP) and pour point temperatures as well as high cold filter plugging points (CFPP). These parameters all relate to the usability of the material as fuel at low temperatures. The cloud point and CFPP relate to the appearance of solid material in the fuel, whereas the pour point is the temperature at which the liquid just flows. In the methyl esters, solid usually appears before the pour point is reached. A typical methyl ester made from a waste fat such as tallow would have a CP of 1017”C, whereas the methyl esters of rapeseed have a value of approximately -2. This may be contrasted with the cloud point of 8°C for palm oil, which contains close to 50% saturated fatty acids. Therefore, the cloud points of methyl esters are potential problems for all biodiesel methyl esters. In fact once the technology is established for producing biodiesel ‘standard’material, cloud point is the most important concern. At the present time the problems associated with cloud points which are too high are addressed in three ways. These are (a) cold filtration, (b) the use of additives, and (c) the blending with conventional diesel fuel. The recent inclusion of B20, a 20% blend of methyl esters with diesel fuel, as an allowable fuel under the US EPACT program, is not arbitrary. At the 20% level, significant reductions in CFPP of the esters can be achieved. This reduction obviously is related to the individual cloud points of the diesel fuel and methyl esters separately. As an example, a beef tallow methyl ester (CFPP = +20), when blended at the 20% level with a petrodiesel having a CP of -3, had a resulting CP of 113. The reduction effect is not linear over the whole range and has the maximum benefit below 20%. The “winterisation” of methyl esters describes a process whereby the esters are cooled, and the solid material which results is fltered off from the liquid fraction’. This raises the obvious question of what to do with the solid portion. Several years ago we studied the pyrolysis of vegetable oils and fats in a trickle bed reactor over selected catalysts4. Surprisingly we found that the pyrolysis over a certain type of activated alumina at 400°C and a weight-hourly space velocity of 0.45 gave “mechanistic” yields of close to 100% with virtually complete de-oxygenation. The products were almost exclusively linear alkanes and akenes, which is to be contrasted with the results of other workers who, under similar conditions, obtained mixtures of oxygenates and aromatic compounds. The linear h e s and allcenes we obtained were completely compatible with conventional diesel fuels in terms of miscibility, but as expected exhibited a wider boiling point range. 1518
In a later study we identified the two major routes by which triglycerides were converted to hydrocarbons5 (see Fig. 1). The first starts with cleavage of the side chain, which is believed to involve a y-hydrogen transfer to the carbonyl group. This mechanism accounts for the formation of the thermodynamically unfavourable terminal l-alkenes. The second mechanism involves P-cleavage to form fatty acids6. Two fatty acids then combine to form a ketone which in turn undergoes the yhydrogen transfer mechanism to form a linear akene and the en01 of a methyl ketone. This en01 rearranges to the terminal aldehyde which then disproportionates to an alcohol and a fatty acid. The fatty acids then go through the ketone route while the alcohol dehydrates to an alkene. Unfortunately, the disproportionation step leads to deactivation of the catalyst. Because this step is part of the route from fatty acids, it is the latter which are indirectly responsible for the deactivation of the catalyst. This was confirmed by using fatty acids as the substrate. It is clear that fatty acid methyl esters should not be capable of producing fatty acids. This study describes the pyrolysis of soybean methyl ester under similar conditions as were previously used for vegetable oils.
T riglyce ride Figure I 0-elimination and 7-hydrogen transfer on a triglyceride molecule.
EXPERIMENTAL SECTION MATERLQLSAND ANALYSIS PROCEDURES Soybean oil was supplied by Procter and Gamble Inc., Toronto, Ontario, Canada. Methyl ester was made by the one-phase method employing a methanolloil molar ratio of 27:1, tetrahydrofuran as cosolvent and 1.0 wt.% sodium hydroxide based on the oil. The methyl ester accounted for 99.7% of the total ester present. The Alcan AA 200 activated alumina used was supplied by Alcan Chemicals, Brockville, Ontario, Canada. The catalyst's BET surface area is 27O-29O2/g, and its pore volume is 40 cm3/g. The pore distribution is binodal, with 66% of pore volume in pores ,less than 30 radius and the rest in larger pores. Its bulk density is 0.75 g/cm M a r e d spectra were recorded on a Nicolet model 598 IR spectrophotometer using the thin film method. Liquid samples were analysed using a Hewlett Packard 5880A series gas chromatograph (GC) equipped with a flame ionisation detector and a DB17 (30 m x 0.53 mm i.d.) fused silica capillary column. The operating parameters were as follows: detector temperature 225OC; injector temperature 225°C; temperature program, 5 min at 50°C; heated at a rate of 5"C/minto 210°C; held for 23
'.
1519
min. The gas flow rate was 59.6 was 59.6 mL/min, which included both the carrier
gas He and the makeup gas. Gas analyses were performed on a similar GC equipped with a flame ionisation detector and a 316 stainless steel column (2.44 m x 3.18 mm i.d.) with a Poropak type QS packing material. The operating conditions were as follows: detector temperature, 120°C; injector temperature, 120°C; temperature program, 1 min at 50°C; heated at a rate of 30"C/min to 150OC; hold for 20 min. Carrier gas (Ar) flow rate was 40.05 mL/min. Gas chromatography-mass spectrometric (GC-MS) analyses were done on a Hewlett Packard 5890 GC equipped with a J&W Scientific DB-5 capillary column (30 m x 0.250 mm i.d.) and using helium carrier gas at a volumetric flow rate of 1 mL/min. The linear flow velocity was 32 c d s . GC-MS was performed by the Department of Chemistry, University of Toronto, courtesy of Professor Tom Tidwell and Dr. Alex Young. Analytical hydrocarbon standards were supplied by Polyscience Corp., IL. Carbon and hydrogen elemental analyses were performed by Guelph Chemical Laboratories, Guelph, Ontario, Canada.
PYROLYSIS APPARATUS AND EXPERIMENTAL PROCEDURE The pyrolysis unit consisted of an insulated 3 16 stainless steel preheater tube (1.3 cm i.d. x 50 cm length) which extended 1 in. into a 316 stainless steel fixed bed tubular reactor (2.5 cm i.d. x 46 cm length), which was heated by a cylindrical block heater. Two type J (iron-constantan) thermocouple probes were used to both monitor the internal catalyst bed temperature and maintain a consistent reactor wall temperature in combination with a temperature controller. A syringe pump, condenser, vacuum adapter, receiving flask, nitrogen cylinder, and gas collection system were connected as shown in Figure 2. The reactor midsection was packed with 40 g of activated alumina, which was held in place by a circular stainless steel screen. The preheater and reactor were operated at 180-190 and 450"C, respectively. The entire process remained at n o m l atmospheric pressure throughout the run. Prior to a run,fresh catalyst was demoisturised at 450°C for a 3 h period by intermittently passing nitrogen over the catalyst bed. High demoisturisation was necessary in order to activate the catalyst. One hour of treatment was inadequate, and two hours was only barely adequate to activate the catalyst. Immediately before commencing liquefaction, a nitrogen environment was established in the pyrolytic system. A syringe pump (Sage Instruments (subsidiary of Orion Research Inc.), Model 355, Sage Instruments Inc., White Plains, New York) was used to inject liquid substrate (ranging between 30 and 100 mL volumes) from a 100 mL syringe into the preheater. From previous studies,3." the optimum feed rate which gave the highest organic liquid yield was 0.34 mL/min, which is equivalent to a 0.46 h-' weight hourly space velocity (WHSV) for crude canola oil. Consequently, this was also the rate employed for these studies. The preheater was angled so the injected substrate proceeded by gravity flow into the reactor. Substrates which were solids at room temperature (i.e. dodecanol) were melted and maintained as liquids by wrapping both the syringe and the preheater entrance area with heating tapes during injection. The pyrolysed product exiting the reactor was cooled with a water condenser and collected in one or more 100 mL receiving flasks immersed in an ice bath. Noncondensable products passed through the gas trap into the brine solution and were measured by brine displacement into a 500 mL, graduated cylinder. After all substrate was injected, an addition 8-15 min period was allowed for the residual substrate to
1520
exit the reactor. The liquid product and syringe were then weighed to determine product yield, and the spent catalyst was also weighed to obtain coke-deposition quantities. Hydrogenation of the pyrolysis product was carried out using a palladium on carbon catalyst. Viscosity was measured according to ASTM D445. The boiling point curve was measured according to ASTM D86. Heat of combustion was measured in a 1241 oxygen bomb calorimeter. Pour point was determined according to the ASTM D97 method.
Figure 2 Experimental set up: (1) 3 16 stainless steel feed preheater tube (1.3 cm i.d. x 50 cm length); (2) block heater containing a 316 stainless steel fixed bed reactor tube (2.5 cm x 46 cm length); (3) catalyst bed; (4) Type J (irodconstantan) thermocouple probe; (5) Type J (irodconstantan) thermocouple with temperature controller; (6) syrmge pump; (7) condenser; (8) receiving flask; (9)gas trap; (10) gas collection vessel; and (1 1) nitrogen cylinder.
RESULTS YIELD AND DEOXYGENATION
A total of 310.3 g of methyl ester were pyrolysed to yield 204.4 g of liquid product which corresponds to a 65.9% actual liquid yield. The theoretical yield based on a typical CI6,CI8composition of soybean oil, and the y-hydrogen transfer mechanism is approximately 76%. Therefore, the yeld based on this limitation is approximately 87%. This shows that random cracking of the ester chain does not play a sipficant role in the reaction. The idrared spectrum shows no carbonyl peaks at 1720 cm-I. Peaks appearing at 907 cm-' and 965 cm-' are probably absorptions due to terminal alkenes and tram glkenes, otherwise the spectrum looks like that of a typical alkane. Consistent with these results, elemental analysis gave 83 -7% carbon, 14.56%
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hydrogen with the balance (1.74%) presumably being oxygen. A typical diesel fuel (D2) analysed in the same way contained 86.7% carbon, 12.8% hydrogen, 0.22% sulphur and 0.05% nitrogen, the balance (0.86%) presumably being oxygen. UNSATURATION AND CHAIN LENGTH
The infrared spectrum indicated the presence of double bonds. The hydrogenation experiment showed one double bond for every 19.8 carbon atoms. The gas chromatograph of the pyrolysed product is shown in Fig. 3. C14 I
J
Figure 3 Gas chromatograph of product from methyl ester pyrolysis.
Approximately 50% of all carbon atoms are in chains of Clz or less (down to c6). In addition very few carbon chains are above CIS. Therefore many of the individual components must be alkanes with the balance being mostly alkenes. Figure 4 shows the gas chromatograph of the hydrogenated material. The dominance of the n-alkene peaks confirms that the majority of the pyrolysis product is linear material, either alkanes or alkenes. However, other minor peaks show that some nonlinear material is also present. SPECIFIC GRA VITY. VISCOSITY AND BOILNG RANGE
The specific gravity of the product was 0.815 at 23°C resulting in an API gravity of 42.2. The specific gravity was slightly lower than a number 2 diesel fuel (0.851). T h s was due to the significant quantities of c6, C7 and C9 material in the product. However, the specific gravity does fall in the range of typical diesel fuel specific gravities (0.813-0.852). The kinematic viscosity of the pyrolysis product (1.803 cSt at 39.6") was again lower than that of a typical number diesel fuel (2.6 cSt at 39.6"C). The prescribed range for number 2 diesel fuels is 1.3-4.1 cSt at 40°C. The boiling point curves for the pyrolysis product and number 2 diesel fuel are shown in Figure 5.
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16
Figure 4 Gas chromatograph of hydrogenated pyrolysed methyl ester product.
Figure 5 Distillation curves for a number 2 diesel fuel and pyrolysed methyl ester product.
Approximately 50% of the pyrolysis product boils below the initial boiling point of the diesel fuel, which is entirely consistent with the specific gravity and viscosity results. The pour point (-39°C) of the product was also considerably lower than a number 2 diesel fuel (-27°C). It is lower than the average value (-21°C) of diesel fuel in the Great Lakes and Eastern Region of Canada and similar to the average value (39°C) in the Western Region of Canada. . One purpose of the pyrolysis was to produce a liquid, whch could be blended with methyl esters, without raising the cloud point of the blend, and preferably lowering it. At a blend of 30% of soybean methyl ester with the pyrolysis product, precipitation of solids did not occur at 5°C. Soybean methyl ester normally throws down considerable solid over a period of time at this temperature.
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CONCLUSION The pyrolysis of soybean methyl ester at 400°C over activated alumina (WHSV = 0.45 h-') produces a high yield of hydrocarbon liquid which is comprised mostly of linear alkanes and alkenes. This material can be blended with methyl esters, thereby lowering the cloud point of the latter. This discovery presents one option for dealing with the solid methyl esters produced in the formation of "winterised" methyl ester biodiesel. REFERENCES 1. Rathbauer, J. (1994) Fatty acid methyl esters (FAME) as special winter fuel. In: Biomass for Energy, Environment, Agriculture and Industry, vol. 2, 8' E.C. Conference (Ed, by Ph. Chartier, A.A.C.M. Beenackers & G. Grassi. Pergamon Press. 2. Badal, C. & Woodward J. (1997) ACS Symposium Series 666, Fuel and Chemicalsfi.omBiomass, Chapter 10, pp. 173-208. 3. Lohrlein, H-P., Anggraini-Suss, A. & Krause, R. (2000) Possibilities and limits of the re-use of cooking oils with animal fat content as motor and heatingjkel, proceedings of the 1st World Conference and Exhibition on Biomass for Energy and Industry, Seville, Spain, June 5-9. 4. Boocock, D.G.B., Konar, S.K., Mackay, A., Cheung, P.T.C. & Liu, J. (1992) The production of alkanes and alkenes by the pyrolysis of triglycerides over activated alumina. Fuel, 71, 1291-1297. 5. Vonghia, E., Boocock, D.G.B., Konar, S.K. & Leung, A. (1995) Pathways for the deoxygenation of triglycerides to aliphatic hydrocarbons over activated alumina. Energy and Fuels, 9, 1090-1096. 6. Leung, A., Boocock, D.G.B. and Konar, S.K.(1995) Pathways for the catalytic conversion of carboxylic acid to hydrocarbons over activated alumina. Energy and Fuels, 9,913-920.
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Bio-Crude-OiVDiesel oil emulsification: main achievements of the emulsification process and preliminary results of tests on Diesel engine P. Baglioni, D. Chiaramonti, M. Bonini Consorzio Interuniversitario Sviluppo Sistemi a Grande Interfase, University of Florence, Via G. Capponi 9, Florence, Italy I. Soldaini Pasquali Macchine Agricole, Via Nuova 30, Calenzano, Florence, Italy G. Tondi Department of Energetics 5'. Stecco", University of Florence, Via S. Marta 3, Florence, Italy
ABSTRACT: In the context of an increasing use of Renewable Energy Sources, biomass derived pyrolysis oil is a very attractive solution: Bio-Crude-Oil (BCO) has potential to replace fuel oil or Diesel in many applications such as boilers, turbines and alternative engines for electricity production. However, technological development for BCO upgrading in order to improve its utilisation is still a problem to be solved. This paper describes the main results of a research project (supported by the European Commission, DG XII, JOULE Programme) aimed at the development of a low-cost physical-chemical and mechanical process for facilitating BCO utilisation in small Diesel engine units. This process is based on the preparation of an emulsion between BCO and Diesel oil. BCO and Diesel oil are not miscible, therefore a third component has to be added to obtain a stable emulsion. This third component is called emulsifier (or surfactant). It changes the interfacial properties of the system avoiding (or delaying) the emulsion's breaking. The developed emulsions - based on different BCOs in terms of feedstock and production facility - have been tested in a conventional small size Diesel unit (6.25 kW)at Pasquali Macchine Agricole (Italy). These tests were aimed at assessing combustion quality, operating performances and emission levels of the Diesel engine helled with the BCO/Diesel oil emulsion: some not structural modifications were made on the engine, mainly adding of components for insulating and cooling.
INTRODUCTION AND BACKGROUND Renewable Energy (RE) sources are currently insufficiently exploited in the European Union: the energetic potential is considerable, but currently they provide a small contribution of less than 6% to the EU's overall energy consumption. RE sources are expected, however, to grow considerably in the future, especially to comply with the commitments at both European and international level on environmental protection.
1525
They can significantly contribute to the Kyoto targets and EC “White Paper” (reduction of 8% of the GHG emissions between 2008 and 2012 and doubling of the RE share in the EU energy balance from 6% to 12% in 2010). REs exploitation is important not only from the environmental point of view: REs are indigenous sources and therefore can contribute to reducing dependency on energy imports (the EU energy imports are currently 50%, a figure that is expected to rise over the coming years if no action is taken, reaching 70% by 2020) and increasing security of supply; furthermore, the RE development can actively contribute to job creation and can be a key point in local (for instance rural) development. In this scenario, a fundamental role will be certainly played by Bioenergy. Technically, bioenergy could be developed in the EU and many other countries on a very large scale given the significant potential of the agro-forestry sector to quickly supply a huge amount of biomass resources: large areas of cultivated land could be dedicated in the near future to biomass crops for energy production. Sustained growth scenario
exajoules 1500
4
Surprise GeoiOcean Solar INewBlomaoa wind Nuclear
I
1 Hydro I Gas
rn
500
Oil & NGLs Coal Trad. Blamaos
0 1860
1880
1300 1920
1940 1960
1380
2000
2020
2040 2060
Fig. 1 “Sustained Growth Scenario” foreseen by Shell Oil Company.
Taking into account that biomass resources can be produced (in a sustainable way) by good practice with yields in the range of 5-10 Toeha per year, bioenergy contribution to the primary energy needs may be substantial: by 2010 - according to “White Paper” recommendations bioenergy should provide in the EU 135 Mtoe/year.
-
1526
This means that an energy production three times the 1995 amount of 45 Mtoe is considered to be a possible target for the year 2010. This additional capacity - 90 Mtoe/year - will require an estimated total investment of 84 billion Euro, with a C 0 2 reduction of 255 MTodyear in 2010. The potential offered by biomass in fact is widely recognised; for instance, Fig. 1 shows the “Sustained Growth Scenario” foreseen by Shell Oil Company (1 Exdoule = 23.8 Mtoe). For biomass conversion technologies, one of the most promising route seems to be the production of biomass derived oil (Bio Crude Oil, BCO) through pyrolysis, given its economic viability and simple integration into conventional energy systems. Pyrolysis is attractive because converting solid biomass and wastes into liquid products presents many advantages in transport (energy density is increased about four-fold), storage, handling, retrofitting, combustion and flexibility in production and marketing. This liquid is of moderate heating value, easily transportable, has potential to be used in thermal power plants, in modified Diesel engines or gas turbines. However, hrther technological development aimed at facilitating BCO utilisation is still necessary: one promising route for BCO upgrading is the production of an emulsion between BCO and conventional Diesel oil.
THE BIOEMULSION PROJECT The aim of the BIOEMULSION project is to develop a low-cost physical-chemical and mechanical process for improving operational properties and performances of pyrolysis oil in small and medium Diesel engine units. The project partners are: 0 0
0 0
CSGI, University of Florence (IT) PMA, Pasquali Macchine Agricole (IT) IEE, lnstitute fur Elektrische Energietechnik, Kassel University (DE) Ormrod Diesels (UK)
and, as project subcontractors, DEF, Department of Energetics University of Florence (IT), Bio-energy Research Group, Aston University (UK) and WIP (DE). The principal activities of the project are the preparation of the BCO/Diesel oil emulsion and the consequent experimental campaign aimed at assessing the operation of three Diesel engines fuelled by these emulsions. The use of a mixture of BCO and conventional Diesel oil in fact is relevant in both socio-economic (e.g. fuel cost, new job creation, rural development) and environmental terms (mainly C 0 2 and sulphur emissions reduction). The final objective of this research project is therefore to achieve a low-cost upgrading process of biomass derived oil for facilitating its utilisation in Diesel units for powerheat generation. The project is supported by the European Commission, DG XII, in the framework of the JOULE Programme. The present paper describes the activities accomplished up to now by the Italian partners (CSGI, *PMA and DEF) within the project. These activities can be summarised as follows: 0 0
0 0
BCO upgrading Development and optimisation of the BCO/Diesel oil emulsification process Emulsions production and characterisation Experimental campaign in a small Diesel engine at PMA factory
1527
The first task has been the provision of a sufficient amount of BCO, to be emulsified with the Diesel oil and hence tested in the engine. Four (4) types of BCO have been selected; these BCOs are different in terms of production facility and feedstock: 0 0 0
0
ENEL 1: produced by ENEL (Bastardo plant, IT), from Canadian oak ENEL:!: produced by ENEL (Bastardo plant, IT), from beech wood Dynal 00 1 : produced by Dynamotive (Vancouver, Canada), from California pine DynalOO9: produced by Dynamotive (Vancouver, Canada), from California pine
The BCOs produced by Dynamotive Technologies - types 1001 and 1009 - are both from California pine and have very similar characteristics. Table 1 shows the main physical-chemical properties of the selected BCOs.
BCO UPGRADING The BCO upgrading has been useful to reduce noxious fractions, to increase pH and to make the selected BCOs more similar to each other. This last point is important as it permits to adopt few classes of surfactants and/or additives and to develop and optimise only one BCO/Diesel emulsification process. However the developed emulsification method does not require BCO upgrading, reducing considerably the operational costs. On the basis of BCO characterisation (both raw and upgraded), finally BCOs referred as ENEL2 and Dyna1009 have been selected for the emulsification process: they have been chosen since they are more manageable and easier to be emulsified in comparison with ENEL 1, which is characterised by high viscosity and ash content, thus causing a low quality emulsion in terms of combustion.
EMULSIFICA TION PROCESS DEVELOPMENT AND OPTIMISATION Diesel oil and BCO is a two phases system since Diesel is insoluble in BCO and vice versa. They are not miscible. If the Diesel oil/BCO system must be used as fuel, a stable emulsion is necessary. In the simplest emulsion a phase (oil or water) is dispersed in the continuous medium (water or oil) in the form of droplets. In this case Diesel oil has been considered as the “oil” phase and BCO as the “water” phase because of its consistent water percentage. Three kinds of emulsions can be prepared according to the value of the BCO/Diesel oil ratio:
0
0
Water in oil emulsions (W/O) are obtained when up to about 45% by weight of BCO is added to the Diesel oil phase. Oil in water emulsions (O/W) are obtained when up to about 45% by weight of the Diesel oil is added to the BCO phase. Bicontinuous emulsions are obtained when the percentage by weight of the two phases is close to 50%.
In the first case, the BCO’s droplets dispersed in the Diesel oil (continuous medium) form the emulsion. In the second case Diesel oil’s droplets are dispersed in BCO (continuous medium). The third case’s description is more complex: in fact ‘‘theoretically’’ there are no droplets and Diesel oil and water phases are continuous and form a “bicontinuous emulsion”.
1528
Table I Physical-chemical properties of the BCOs. Physical-chemical properties Viscosity (cPoise): 25 OC Viscosity (cPoise): 35 OC Viscosity (cPoise): 45 O C
ENELl
ENEL2
Dyna-1009
70 1 36 1 150
9.8 6.7 4.1
46.2 27.1 14.1
5.79 44.06 48.68 0.08 co.0 1
5.94 34.96 58.70 0.11
6.3 1 43.52 50.07 0.07
AlkaWBCO (ppm): Na K Ca Ash weight (YOwlw)
103 339 1096 1.38
179 694 1561 0.28
6 8 123 0.03
AlkaWAsh (mg/g): Na K Ca
7.5 24.6 79.4
64 248 558
20.5 26.9 410
22.1[23.1]' 1.230 3.2 16942 18205
43.6[45.1] 1.165 3 .O 14577 15873
30.5[3 1 . I ] 1.175 2.6 16779 18155
Elemental analysis ("10w/w): H C 0
N S
Water weight (YOw/w) Density @ 20 OC (g/cm3): PH LHV (kJ/kg) HHV (kJ/kg)
Underlining that BCO and Diesel oil are not miscible' a third component must be added to obtain a stable emulsion. This third component is called emulsifier (or surfactant). It changes the interfacial properties (namely interaction potential between droplets) of the system avoiding (or delaying) the emulsion's breaking. The choice of an emulsifier is a step not completely random but it is driven by several empirical rules. These rules are based on the primary condition for the emulsifier to be efficient in a two phase, tivo-liquid emulsion. It has to be localised at the oil/water interface to a maximal extent. This criterion is exemplified in the HLB (Hydrophilic Lipophilic Balance) number. The HLB number for a certain molecule denotes the balance between its hydrophilic/lipophilic properties only; the number itself does not give any information about the stabilising efficacy of the emulsifier.
' Numbers in square brackets indicate water content values measured adopting the Karl Fisher method. Shaking a BCO/Diesel oil system, an emulsion is obtained but its stability is approximately only one miiiute at 25°C.
1529
The HLB number method is based on the fact that the emulsifier is optimal in a wateroil system in which the properties of the oil matches the surfactant. Hence, each wateroil combination is characterised by an HLB number. More practically, emulsifiers characterised by HLB number: 0 0 0
from 4 up to about 8 stabilise W/O emulsions; from 8 up to about 10 stabilise bicontinuous emulsions; from 10 up to about 18 stabilise O/Wemulsions.
Fig. 3 BCOIDiesel oil mixture (left) and emulsion (right) The Phase Inversion Temperature (PIT) or HLB temperature concept relates the emulsifier selection to the temperature at which an emulsion stabilised by a non-ionic emulsifier of the polyethylene glycol type changed from oil-in-water to water-in-oil with rising temperature. The tested surfactants were a large number (approximately one hundred), both of commercial type and chemically pure composition. Among them, cationic, anionic, zwitterionic and non-ionic stabilisers were tested. Moreover, for each of these classes of additives, polymeric and non-polymeric surfactants were tested. Additives presenting sulphur or nitrogen in their composition have been neglected, to avoid an increase in NO, and SO, gaseous emissions during the emulsion combustion. Preparation of the emulsions The water-in-oil emulsions have been prepared by adding the surfactant to Diesel oil and thereafter adding the BCO to the resulting mixture. Bicontinuous emulsions have been formulated by adding the surfactant to Diesel oil, and then mixing to BCO the resulting mixture. Oil-in-water emulsions have been formulated by adding the surfactant to BCO and thereafter adding to the resulting mixture the Diesel oil during emulsification. The emulsification process has been carried out using a homogenisation unit. The temperature during mixing has been preferably maintained between 60-65 "C and the emulsification process continued until a homogeneous single phase has been obtained. During the emulsifier choice phase, emulsions were prepared using a discontinuous process on a little amount of mixture (from 20 to 500 8). 1530
Having identified the emulsifier, a continuous process has been employed on greater amount of material (up to 10 kg). In this configuration, the tank was a heating case controlling the desired temperature in the continuous homogenisation unit. Such emulsions present a stability up to some months; emulsions stable even for one year have been produced. It is important to underline that emulsions can be prepared also at room temperature; in this case the emulsion stability decreases (about 20 days). However, the destabilisation of the emulsions, prepared at both high and room temperature, is not an irreversible process, since it is sufficient to gently shake the fluid to obtain again a homogenous system. The phase diagram of the-system BCODiesel oil\emulsifier has been investigated in order to identify the best w/w ratio in terms of homogeneity, stability, engine performance, and operational costs. Diesel oil
Fig. 4 Phase diagram for the BCO\Diesel oil\emulsifier ternary system.
Fig. 4 shows the phase diagram of the analysed system. The interesting region as far as regards the present work is located on the left of the dashed line (any concentration of BCO and Diesel oil but emulsifier <5% by weight). The circles are the regions that correspond to the 25% (W/O), 50% (bicontinuous) and 75% (O/W) of BCO by weight. These three regions have been studied more accurately. It has been noted that: 0 0 0
The emulsions are more stable than correspondent BCO. The emulsions were stable and opaque with no evidence of phase separation. Higher the emulsifier content, higher the viscosity and stability of emulsions. The optimal range to have an acceptable viscosity is between 0.5 and 2%. Emulsifier up to 4% has been added, in this case an additive (like n-octanol) must be employed to decrease emulsion viscosity.
153 1
0
0
0
The evaluated emulsions stability versus time at high temperature (approximately 70 "C) is about 3 days. The high BCO content emulsions had very high viscosity which would make them difficult to use in most applications. Recycling cycles of the emulsion in the feeding engine system through the injectors and pipes have not changed the emulsions properties.
The destabilisation of an emulsion goes through several consecutive and parallel steps before the final stage of separated layers is reached. As a first step, the droplets move due to diffusion (or stirring), and if the repulsion potential is too weak, they become aggregated to each other; flocculation has taken place. The single droplets are now replaced by twins (or multiplets) separated by a thin film. The thickness of the thin film is reduced due to the Van der Waals attraction, and when a critical value of its dimension is reached, the film bursts and the two droplets form a single droplet. Coalescence has occurred. In parallel with these phenomena, the droplets rise through the medium (creaming) or sink to the bottom (sedimentation) due to differences in density of the dispersed and continuos phase. The final result is a highly concentrated emulsion at the top or bottom of the container, and the increased number of droplets per unit of volume increases the flocculation rate in a most decisive manner. The flocculation and coalescence process lead to larger and larger droplets until, finally, a phase separation has occurred. Prepared emulsions tend to flocculate and finally (after one-two months) the flocks sediment and coalescence. It seems that coalescence process is very delayed by the emulsifier. In order to investigate the flocculation process, a study through an optical microscopy of the droplets and flocks has been employed. Unfortunately, a little difference in the contrast (brown against pale yellow) between droplets and continuous phase has not permitted a good determination of mean droplets dimensions. Probably, mean diameter should be few microns (1-5). EMULSIONS PRODUCTION AND CHARACTEHSATtON
Six (6) different emulsions with different weightlweight ratio (from 25 to 75%) and based on two different BCOs (ENEL2 and Dyna1009) have been produced for utilisation in the experimental campaign. The following tables report the principal chemical-physical properties of the emulsions derived from Dyna1009 and ENELZ. It can be noticed that the physical-chemical properties of the emulsions are strictly related to the corresponding properties of the initial fuels (that is Diesel and Bio Crude Oil). Other additives (conventional lubricating and amidic or aminic co-surfactants) have been added (up to 0.5% by weight) to the emulsions to improve the lubricating properties and decrease the' corrosive behaviour due to BCO. In order to have large amounts of BCO/Diesel oil emulsions, an apposite unit has been designed and manufactured. The constructed emulsifier system is able to produce L 50 lih of emulsion, since the existing laboratory facility was insufficient to the production of the amounts (several kilograms) necessary for the experimental campaign.
1532
Table 2 Physical-chemical properties of DynalOO9 emulsions.
Physical-chemical properties BC025 D1009 Viscosity (cPoise): 25 O C Elemental analysis (YOw/w): H
C 0
N S
BC050 D1009
BC075 D1009
9.79
25.41
n.a.
12.27 74.63 13.04 0.03 <0.01
10.50 61.82 27.6 0.05 <0.01
39.88 0.07
2
6 8 125.1 0.03
8.98 5 1.03
Alkali/emulsion (ppm): Na K Ca Ash weight (Yow/w)
41.1 0.0 1
4 5 83.2 0.02
Alkali/Ash (mg/g): Na K Ca
20.1 26.0 41 1
20.3 26.4 415
20.4 26.7 417
7.3 [ 7.71 0.923 3.O 35810 38486
14.9 [15.3] I .008 2.9 28347 30637
22.5 [23.1] 1.094 2.8 21558 23517
Water weight (YOwlw) Density @ 20 O C (g/cm3): PH LHV (kJlkg) HHV(kJ/kg)
3
EXPERIMENTAL CAMPAIGN The experimental campaign is aimed at assessing combustion characteristics, operating performances (e.g. efficiency, flexibility and hours of operation) and noxious emission levels of a small size Diesel fbelled with the BCO/Diesel oil emulsions. The engine used to perform the test on the emulsions is a mono-cylinder direct injection Diesel engine manufactured by the Italian company Lombardini. The engine is part of a generator set of 5.4 kW and it has been only slightly modified, because the aim of the project was to succeed in running an unmodified engine. However, some minor modifications have been necessary, mainly an adding of components (for insulating and cooling) aimed at optimising the operation of the injection system when the engine is run with the emulsion.
1533
Table 3 Physical-chemical properties of ENEL2 emulsions.
Physical-chemical properties BC025 ENEL2 BC050 ENEL2 BC075 ENEL2 9.43
28.3
n.a.
N
12.11 71.59 16.13 0.04
8.83 44.98 45.88 0.09
S
<0.01
10.09 56.64 33.05 0.07 CO.01
0.07
91 347 78 1 0.14
134 520 1171 0.2 1
64.3 248.6 557.1
65.0 247.9 557.9
63.8 247.6 557.6
10.7 [I 1.1 ] 0.92 1 3.2 35706 38348
2 1.6 [22.0] 1.003 3.1 27134 29335
32.4 [32.9] 1.084 3.O 19259 21 185
Viscosity (cPoise): 25 OC Elemental analysis (YOw/w): H C 0
Alkali/emulsion (ppm): Na K Ca Ash weight (YOw/w)
45 174 390
AlkalVAsh (mg/g): Na K Ca Water weight (YOwlw) Density @ 20 O C (g/cm3): PH LHV (kJkg) HHV (kJ/kg)
The main characteristics of the Lombardini engine (see figure 4) are reported in the following table. Table 4 Characteristics of the Lombardini engine. Model Cylinders Bore x stroke Displacements Max rotational speed Max power Max power for electrical generation Max torque Dry weight Injection hole’s diameter
1534
Lombardini 6LD400 1
86 x 68 mm 395 cm3 3600 rpm 6.25 kW 5.4 kW (a 3000 rpm) 19.6 Nm (a 2200 rpm) 45 kg 0.2 mm
Fig. 4 Lombardini 6LD400 engine
A large number of tests have been performed, using the emulsions prepared during the previous activities. The most used percentage was the 25% BCO. However, also emulsions with higher BCO content have been tested. Above all one test with 50% BCO Dyna 1009 gave promising results: the engine was able to run correctly for over 30 minutes; when the engine was switched to diesel oil its performance drastically decreased. Among all the tests executed at PMA, the most significant in terms of running time and performances are the four reported in table 5 .
Table 5 Main data of the engine tests Test number and fuel type
Warm Running Diesel oil DeveloUP with after-test ped Diesel oil emulsion Power
I I
1 IBC025 D1009 I 2 BC05D1009
3 BC025D1009 4 IBC025 Dl0091
min.1I 20 I
min.1I 78 I
25 32
53 116
301
min.1I 81 12
25
kW 3.1 3.2 3.2-3.3
151 3.0-3.1
811
tion.
tion
Iitresl
I
I
none I 1.51 water spray 4.51 water spray1 2.61 water jacket1 1.11
The first test was performed to evaluate the maximum time the engine could operate with 25% emulsions (used BCO: Dynal009). During the initial 4-6 minutes the engine ran correctly, without significant increase in smoke and emissions, only a little higher than Diesel oil. Then the engine operation suddenly started to get worse and in about 20 minutes the smoke became heavy black and the emissions were so high that it was better to detach the analysis probe. The engine continued running for 78 minutes, but it must be stated that during the last 15 minutes the engine worked very erratically. The engine was then serviced. The injector was completely stuck in its clearance and a notable enlargement of the holes occurred a fmal diameter of 0.45 mm was measured, the starting value being equal to 0.2 mm. All the injector items in contact with the emulsion were covered with a black sticky deposit. 1535
The second test was performed using a low content BCO emulsion (5%). In this test it was used a water atomiser spraying the water on the visible part of the injector to keep the fuel as cool as possible. Even the recirculation pipe has been cooled. The test was positive: the emissions deteriorate slightly with time and the engine ran for more than 50 min. very plain and without smoke increasing. The injector showed less sign of wear: black deposit were concentrated on the tip of the needle only, no blockage of the needle, holes only a little larger (0.25 mm). After this good result the 25% BCO emulsion has been tested with the cooling system. The third test lasted for approximately 2 hours, the longest one: it was stopped because the emissions became very bad after two hours of operation. However, during an initial period of about 30 minutes, the engine performances were quite constant. Afterwards, the engine operation started to get worse, first slowly and then more rapidly: the reason is the progressive enlargement of the injection hole, probably due to the erosion caused by polymerisationkarbonisation of the BCO particles occurring on the tip of the injector and of the needle. The final effect was an incorrect fuel injection in the combustion chamber and the consequent worsening of the engine performance. However, as in test no 2, the arranged cooling system avoided the blockage of the needle and delayed the enlargement of the injection holes. Further modifications on the engine were mainly in the direction of keeping the fuel as cool as possible, preventing BCO polymerisation and formation of particles not only sticky but also eroding. In fact, these particles, although they are soft, when accelerated 0 on the tip of injector probably erode the metal of the injector itself. to ~ 1 0 m/s Therefore, for the fourth test a water jacket has been developed to be installed around the tip of the injector, in order to cool as much as possible this zone, where the temperature is higher. This modification brought to a further slight improvement of the engine performance in comparison with previous tests: a smoother behaviour of the engine and a more regular decreasing of performances. The injector’s needle presented less sticky deposits than previous tests.
Fig. 5 Injector’s hole before and after emulsion running (test no 4) Fig. 5 and 6 shows the injection hole and the needle tip as new (hole’s diameter: 0.2 mm) and after 80 minutes of emulsion running (diameter: 0.45mm). The pictures have been taken by a camera with a microscope adapter. Finally, the next charts show CO and NO, emissions as well as exhaust gas temperature and engine efficiency vs. running time, in order to make an analysis comparing the different performances of the engine with the corresponding modification on the engine. 1536
Fig. 6 Needle tip as new and after emulsion running (test no 4)
The behaviour of the emissions can give an idea of what happens in the engine. The increase in CO emissions is mainly due to the poor injection (large fie1 droplets) and consequently a bad combustion. The NO, decrease versus running time is due to low temperature in combustion chamber caused by bad combustion. This also cause the higher temperatures measured in the exhaust pipe, since the combustion reactions, not completed in the combustion chamber, continue outside in the pipe. From the chart it is possible to note that the engine performance increased when the cooling system was adopted: for a long period of time the emissions remained constant until the engine operation got worse due to the deterioration of the injector’s components. Moreover, the good operation of the engine when fuelled by an emulsion with a low BCO content (5%) should be recognized: the measured parameters have been constantly positive and the deterioration of the injection system is limited in comparison to the other tests executed in similar conditions. Exhaust gas temperature
0
10
20
30
40
50
60
time (min)
Fig. 7 Exhaust gas temperature versus running time
1537
70
80
CO concentration
0
10
20
30
40
50
60
70
50
60
70
80
70
80
80
time (min)
NOx concentration
0
10
20
30
40
time (min)
Fig. 8 CO and NO, emissions versus running time
0,22 0.2 0,18
P .$ 0,16 F
0.14
0.12
0,1
0
10
20
30
40
50
60
time (min)
Fig. 9 Engine efficiency versus running time
1538
CONCLUSION The BCOIDiesel oil emulsification process has been developed and optimised, to obtain a fuel easier to be handled, stored and used in Diesel engine units. On the basis of this process, stable emulsions have been prepared: in some cases the stability reached even one year, thus representing a notable improvement in comparison with whole BCO, in general presenting a limited stability over the time. The produced emulsions have been tested on a 6.25 kW Diesel unit: the most interesting results have been achieved with emulsions having a BCO content not greater than 50% by weight. The performed tests pointed out the following issues 0
0
0
The deterioration of the injection system is probably due to erosion phenomena; The modifications on the engine aimed at cooling the fuel at the tip of the injector caused a considerable improvement of the engine performance; The problems related to the utilisation of emulsions as fuel are limited to the injection system: further tests with the. adoption of special coating materials are scheduled in the future.
In conclusion, the tests performed at PMA facility showed that the engine is able to bum this kind of fuel, however some problems have been encountered. Some of them have been successfully solved by insulation and cooling. The remaining problems, especially the erosion phenomena on the tip of the injector, are expected to be solved by means of surface treatments of the injector itself.
ACKNOWLEDGEMENTS
Tha authors wish to thank the persons and Institutions that worked andor are working at this project, in particular Prof. Bridgwater from Aston University, Graziano Crocetti from PMA, Klaus Gartner from IEE, Archie J. Webster from Ormrod Diesels. The research was funded in part by the European Commission in the framework of the Non Nuclear Energy Programme JOULE 111. A particular mention to Mario Frias, responsible on behalf of the European Commision for the Bioemulsion Project. REFERENCES I . Contract JOR3-CT98-0307. Midterm (July 1999), Progress reports (January 2000) 2. Leech J. (1997) Running a dual fuel Diesel engine on crude pyrolysis oil. In: Biomass gasification and pyrolysis. State of the art andfuture prospects (Edited by M. Kaltschmitt and A.V. Bridgwater), pp. 495-497. CPL Press, Newbury, UK 3. PyNE Newsletter, Bio-Energy Research Group, Aston University, UK. 4. A.V. Bridgwater et al. (May 1999) Fast Pyrolysis of Biomass: a handbook (Edited by IEA Bioenergy). CPL Press, Newbury, UK 5 . Proceedings of the “10’ European Conference & Technology Exhibition Biomass for Energy and Industry” (Edited by H. Kopetz et al. June 1998).
1539
Production Of Diesel Fuel Additives From The Rosin Acid Fraction Of Crude Tall Oil R. Coll’, S. Udas2,W.A. Jacob9 ‘Sewei de Tecnologia Quimica, Rovira i Virgili University, Av. dels Paibos Catalans, s/n, 43007 Tarragona (Spain) 2Department of Chemical Engineering, University of MissouriColumbia, W2033Engineering Bldg. East, Columbia, MO 65211 (USA)
ABSTRACT: This work aims to develop processes for converting the rosin acid fraction of crude tall oil into diesel fuel additives. Fractionation of crude tall oil produces an excess of rosin acids. This material is an abundant, inexpensive and chemically acceptable feedstock. Two components, abietic acid and dehydroabietic acid, comprise about 70% of the rosin acid fraction of tall oil. Hydrotreatment of these compounds allows obtaining saturated hydrocarbons that can be used as diesel fuel additives. In this paper, we report the first results obtained in a laboratory-scale batch reactor. Three different factors (temperature, hydrogen partial pressure and space velocity) were studied using two commercial NiMo catalysts. Operating conditions for each experiment were selected making use of factorial design-techniques. The resulting liquid products, mainly saturated tricyclic ring hydrocarbons, were analyzed to determine their suitability to be used as diesel fuel additives. The results are promising enough to encourage further research on this topic, which could lead to the future use of tall oil rosin acid as a source of fuels.
INTRODUCTION The increasing use of biomass-derived fuels is m a d y due to their environmental advantages in comparison with fossil fuels. Besides these advantages, it is widely accepted that biomass-derived fuels do not show any particular technical problem when they are used in standard engines. Ethanol, for example, has been successfully employed as a gasoline additive or gasoline surrogate. For diesel engines, the use of bio-diesel derived from esterification of vegetable oils has also achieved a considerable success. In this case, however, the process is still in an incipient stage. The main problem associated with all these attempts is the question of economic feasibility. One
1540
of the ways to overcome this difficulty would be the use of an almost-free raw material that could be obtained in large quantities from a single source. Tall oil fulfils all of these requirements. Tall oil is a natural product of pine trees that is isolated by means of the Kraft pulping process. It is composed of the ether extractable, non-lignin, non-cellulosic portion of the pine tree, and it must be fractionated via steam-vacuum distillation for commercial use. Fractionation of one metric ton of crude tall oil produces about 350 kg of rosin acids, 300 kg of fatty acids and 350 kg of distillated tall oil, head and pitch I. The economic value of tall oil is primarily derived fiom its fatty acid fraction. Tall oil rosin production has recently been driven by demand for its co-product, tall oil fatty acid. United States production of tall oil rosin in 1994 was 265,000 metric tons, exceeding consumption by nearly 50,000 metric tons. Therefore, the rosin acid fraction of crude tall oil is an under-used byproduct in abundant supply, with attractive characteristics as a precursor to high value chemicals. Two components, abietic acid and dehydroabietic acid, comprise about 70% of the rosin acid fraction of tall oil. Therefore, it is a relatively simple mixture, which is also a beneficial characteristic for its use as a feedstock. Despite its low cost and abundance, tall oil has only recently been considered as a potential fuel source. Initial studies have focused on the unfractionated material 2-8. Direct use of tall oil is not possible since it produces excessive corrosion on certain parts of the diesel engines; it has deficient rheological properties at low temperatures and produces unacceptable contamination of the lubricating oil and coking of the diesel engine '. Therefore, some kind of previous processing becomes necessary. Esterification and hydrotreatment are two interesting possibilities. Hydrotreatment has one important advantage: the process is already employed in refmeries so it can be applied without substantial modifications in producer plants. Moreover, the product has the same chemical characteristics as the regular present fuels, so no changes would be required in present-day engines.
An HO
bH
Abietic acid
Dehydroabietic acid
Fig. 1 Chemical structures of abietic acid and dehydroabietic acid.
1541
This work, however, does not deal with crude tall oil but with the under-used portion of fractionated tall oil, the rosin acids, which, as we mentioned before, show some advantages over crude tall oil. To our knowledge, this is the first time that the focus is on this fhction, although the work of Schuller’s group provides a relevant basis for exploration 9-11.
TECHNICAL APPROACH The chemical structure of all rosin acids is similar to those of abietic or dehydroabietic acids (‘Fig. Z’). To obtain fuel additives, carboxylic acid functionality must be removed and double bonds saturated. A two-stage process can achieve both objectives. The first step would consist of the cleavage of carboxylic groups (either thermally or catalytically) producing a hydrocarbon and COz. Then, hydrogenation of the unsaturated bonds with an appropriate catalyst (Ni, NiW, Pd or Pt,for example) should be performed. An alternate process involves the simultaneous hydrogenation of carboxylic groups and unsaturated bonds. Partial hydrogenation of carboxylic groups would yield oxygenated compounds (aldehydes or alcohols), as shown in Paths 1 and 2. These products can be interesting as additives with specific properties. Complete hydrogenation would produce a hydrocarbon structure (Path 3). R-COOH
+
H2
+ + +
R-CHO R-CH20H R-CH3
(Path 1) (Path 2) (Path 3)
At the same time as the carboxylic functionality is removed, double bonds and aromatic rings would probably be hydrogenated. This second process would require significant amounts of hydrogen, which should be taken into account when the economic feasibility of the process is evaluated. The characteristics of this process suggest the use of hydrotreating catalysts, most of them containing molybdenum and either cobalt (CoMo) or nickel (NiMo) on an alumina carrier. This process is simpler than the first one, although the extent of double-bond saturation would probably be lower. The higher simplicity of this second alternative, however, makes it more interesting to be studied. Gasoline-additive production would require one more step: theimal cracking or hydrocracking to reduce the molecular weight of the material. J n this work, however, we have just focused on the first part of the process, which intends to obtain diesel fuel additives.
EXPERIMENTAL Hydroprocessing reactions were carried out in a 500-mL batch reactor. The reactor is charged with 4 g of a commercial sulfided NiMo catalyst, supplied by Haldor Topsoe, Inc. and Criterion Catalysts, and 40 g of tall oil rosin (Unitol@ NCY), supplied by Arizona Chemical. The reactor is first purged with Nitrogen and then filled with
1542
hydrogen to 3 bar. After reaching 150°C, the hydrogen partial pressure is increased to the run value. The last part of the heating period is carried out at high pressure. No additional hydrogen is added during the experiment, so its partial pressure is decreasing as the reaction goes on. A stirrer is used to improve the reactive-catalyst contact. When the reaction is over, the reactor is cooled and then depressurized. Three operating variables have been analyzed: hydrogen partial pressure, temperature and reaction time. A factorial design of experiments was used to choose the conditions of each run. The levels of the three factors used during the runs are shown in ‘Table 1’. In the same table we can see the expressions employed to normalize the factors. Table I Operating conditions employed. Factor
LeveI -
Level +
x,
Temperature
350°C
400°C
(T (“C) - 375) / 25
H2partial pressure
100 bar
150 bar
(Pm(bar) - 125) / 25
2hr
4hr
(time (hr)- 3) / 1
Time
Product analysis The amount of liquid product was measured to find out the reaction yield for each experiment. Density and boiling-point distribution were analyzed by ASTM D97 and ASTM D2887 (simulated distillation) standard methods. Simulated distillations were converted to ASTM D86 distillations to allow calculating cetane indexes according to ASTM D4737. Liquid products were also analyzed by MS-GC to obtain their qualitative composition, and by elemental analysis (C, H, S and 0) to check the extent of hydrogenation.
RESULTS Liquid Yield Liquid yields, calculated as mass of product recovered divided by mass of rosin acid introduced in the reactor, range from 57 to 85 % (‘Table 2’). There are significant differences between the catalysts employed. Catalyst A yields amounts of liquid product markedly lower than those obtained with catalyst B. This suggests that the fust catalyst favors cracking reactions whereas the second one does not, at least at the operating conditions explored. For liquid yield, as well as for the other variables, the effects of temperature, hydrogen partial pressure and duration of the experiment have been calculated.
1543
Quantification of these effects was achieved by adjusting the response variables versus the normalized factors by the least-square method. The general form of these equations is:
Variablemeasured = A0 + A(T) X, + A ( P Hx~p H 2
+ A(time)
(Eq. 1 )
where AO, AV), A(PHJ and A(time) are the parameters adjusted by the least-square method andxi are the normalized factors. Observing the values of these effects (‘Table 3’), the differences between the two catalysts become clearer. For catalyst A, liquid yields are strongly affected by temperature (in a negative way) and by hydrogen partial pressure (in a positive way). Duration of the experiment does not seem to have a significant effect, at least for the times employed. For catalyst B, on the other hand, the results are just the opposite. Neither temperature nor hydrogen partial pressure affect significantly to liquid yield, whereas longer experiments produce lower amounts of liquid recovered, probably due to an increase in the extent of cracking. Correlation coefficients between these variables confirm this conclusion. These results are consistent with the idea that catalyst A favors cracking more than catalyst B does.
Table 2 Liquid yields for the two catalysts.
P
Exper.
Catalyst
T (“C)
EXPO 1
A
EXPO2
A
EXPO3 EXPO4
A A
350 400 350 400
100 100 150 150
2.0 4.0 4.0
EXPOS
B B B B
350 400 350 400
100 100 130 150
2.0 4.0 4.0 2.0
EXPO6 EXPO7 EXPO8
(bar)
time (h)
2.0
xT
+ + +
+
xPH2
+ + + +
&me
Yield
+ +
66 57 78 73
+ +
85 75 77 84
Product Composition We have performed two kinds of analyses in order to determine the chemical composition of the liquid products. First, we analyzed the elemental composition of the product. The raw material is composed of 76.2% of C, 9.6% of H and 14.2% of 0 (percentages in mass). S u l k was also analyzed, but it was not detected (detection limit = 0.05%). These percentages suppose an H-to-C ratio of 1.52. For the hydrogenated products, no oxygen was detected, except for experiment EXPO4 (1.25% in mass). Therefore, we can a f k n that carboxylic groups are either completely reduced to -CH3 or removed as COP.
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Hydrogen consumptions were about 200 NLkg rosin acid. The extent of double bond and aromatic hydrogenation in each experiment can be measured by comparing the product H-to-C ratios. Supposing that no cracking is produced and an abietic or pimaric acid structure, the maximum H-to-C ratio achievable would be 1.80. Results for catalyst A range from 1.40 (EXP02) to 1.72 (EXP03), with 1.59 as average. EXP02, as we saw before, yielded the smallest amount of liquid, surely due to a high extent of cracking. This is also the reason for the low extent ofhydrogenation in that experiment. Experiment ExP03, on the other hand, which yielded the hghest amount of product with this catalyst, also gave the highest extent of saturation. For catalyst B, conclusions are very similar. Hydrogenation is maximum at 350°C,150 bar and 4 hr, and minimum at 400°C and 100 bar. Liquid yield and extent of hydrogenation, hence, seem to be proportional.
Table 3 Effects of temperature, hydrogen partial pressure and reaction time. Catalyst
A0
A(T)
A(Pw)
A(tim)
A
68.4 80.3
-7.3 -1.8
14.3 0.8
-8.5
B
28.0 29.9
0.9 4.9
4.8 3.1
-1.4 1 .o
BPlO%(OC):
A B
179 164
-78 -81
-46 -19
-12 -21
BP 50 % ("C) :
A
B
296 284
-47 -80
-1 3
2 -9
BP 90 % (OC):
A B
346 351
-1 5 -24
-3 -7
-7 -6
C.I. :
A B
35.1 36.7
-5.8 0.0
6.8 4.4
-4.8 0.6
H/C:
A
1.59
B
1.58
-0.18 -0.10
0.14 0.10
-0.05 0.00
Yield (YO) :
B Density (OAPI ):
A
-2.3
The effects of the three operating variables (temperature, hydrogen partial pressure and reaction time) on extent of hydrogenation have also been quantified following the same procedure as for liquid yield (as well as for the other response variables studied). Results, shown in 'Table 3', reveal that the three operating conditions also affect in the same way to both liquid yield and extent of saturation. Temperature and duration of the
1545
experiment decrease hydrogenation and yield, whereas a high hydrogen partial pressure improves them. We have also analyzed the products by GC-MS. Identifying all the compounds present in the liquid product is not possible because a perfect separation is not possible by GC. However, these analyses are still useful to:
-
Observe the amount of C20 and C19 compounds. C20 are produced by reduction of the carboxylic group to -CH3. C19 are produced by cleavage of the carboxylic group (producing COz). Observe the presence of aromatics. Observe the characteristics of the cracked compounds, that is observe whether the tricyclic compounds are present or have been cracked, or what the amount of linear compounds present in the product is.
The MS-GC analyses reveal that most of the product is composed of C19 compounds. Therefore, we can a f f m that the main mechanism for carboxylic group removal is by producing COz. Tricyclic structures of the raw material are still preserved in the final product, especially in the experiments performed at 350°C.The products also show significant amounts of aromatic compounds. This fact points out that aromatic hydrogenation is not good enough. Therefore, future efforts should be carried out to improve this aspect of the reaction and produce better quality fuels. The use of more specific hydrogenation catalysts should help to achieve this purpose.
-100
200
300
400
S O
Boiling point (‘C)
Fig.3 Boiling point distribution for products with catalyst A.
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Boiling Point Distribution In 'Fig.3 - 4 , boiling point distributions for all the products are shown. The extent of cracking in each experiment is also easily deduced from these graphics. As we can see, there exist important differences between the various operating conditions employed. Experiments at 400°C produce significant higher amounts of cracked compounds, whch c o n f m the conclusions obtained from liquid yields. However, the characteristics of the products are not so different depending on the catalyst employed. The effects of operating conditions on boiling point distribution were calculated ('Table 3'). Temperature appears to be the most mportant factor in all the cases. Hydrogen partial pressure and, especially for catalyst B, reaction time also show significant effects on BP 10% (10% recovery temperature, in "C). All these variables affect cracking reactions in a negative way (decreasing the boiling points). The trends are the same for both catalysts.
3.00
100
FGGG
200
300
400
I
500
Boiling point ("C)
Fig. 4
Boiling point distribution for products with catalyst B.
Density Density is another variable that must be taken into account to study the applicability of our product as a diesel fuel. In general, diesel fuels usually have densities between 30 and 40"API (diesel fuel No. 2) or between 37 and 45"API (diesel fuel No. 1). Our product densities range between 25 and 35"API depending on operating conditions. Therefore, we can conclude that the product we obtain can fall into the specifications for diesel fuels.
1547
Cetane Index
Cetane index is an estimation of the cetane number, which measures the anti-knock properties of the fuel. This index is more used than the cetane number because it can be calculated by a four-variable equation (density, BP lo%, BP 50% and BP 90%) according to ASTM D4737 standard method, instead of the test engine needed to measure cetane number. In our case, the relatively low amount of product obtained in these initial experiments does not allow determination of the cetane number, so we have used the cetane index to evaluate and compare our products, an approximation perfectly valid at this initial stage of investigation. Results show that rosin acid produces fuels with low cetane indexes, between 28 and 42 for catalyst A and between 34 and 43 for catalyst B. According to ASTM D975 (Standard Specifications for Diesel Fuel Oils), diesel fuels must have a minimum cetane number of 40 (diesel no. 2). Therefore, some of our products seem to fall into specification for diesel fuels. However, it must be taken into account that cetane index is just an estimation of cetane number, and this estimation may be inaccurate. According to ASTM D4737 standard method, this equation possesses some limitations:
-
-
It is not applicable to fuels containing additives for raising the cetane number. It is not applicable to pure hydrocarbons, or to non-petroleum fuels derived from coal. It can be inaccurate if it is applied to residual fuels or crude oils
A study about different diesel fuels 12, for example, showed that cetane index underestimates cetane number when it is lower than 43, sometimes more than 5 units (C.N.=40, C.I. = 35.1). Therefore, the actual cetane numbers for our products can be hgher than the measured ones. In the future, producing higher amounts of product to measure actual cetane numbers appears to be necessary to c o n f i i their suitability as diesel fuel. Regarding the effects of the operating conditions, we can affirm that cetane indexes improve when hydrogen partial pressure is high. For catalyst A, they decrease at h g h temperatures and reaction times. For catalyst B, these last two effects are not significant.
CONCLUSIONS The research on the use of hydrotreated rosin acids as fuel additives is in an incipient phase of development. However, the first set of experiments reported in this work gives promising results that encourage continuing the work on this topic. Liquid yields' are high, especially when hydrotreatment is performed at 350°C and high hydrogen partial pressures. No coking or deactivation was observed. Further research is necessary to optimize operating conditions in order to maximize hydrogenation and to extend the applicability of the products for specific functions in the fuel. However, tall oil rosin appears to be a suitable source of diesel fuel and/or diesel fuel additives.
1548
ACKNOWLEDGEMENTS We would like to thank the University of Missouri Research Board for the economical support provided, Criterion Catalysts and Haldor Topsoe for catalyst samples, Arizona Chemical, NREL and Dr. Esteban Chornet for h s comments and suggestions.
REFERENCES 1. 1997 North American Pulp and Paper Fact Book 2. Sharma, R.K. and Bakhshi, N.N. Catalflc Conversion of Crude Tall Oil to Fuels and Chemicals over HZSM-5: Effect of Co-feeding Steam. Fuel Process. Technol. 27, 113-130, 1991, 3. Sharma, R.K. and Bakhsh, N.N. Upgrading of Wood-derived Bio-oil Over HZSM-5. Bioresource. Technol. 35,57-66, 1991. 4. Sharma, R.K. and Bakhshi, N.N. Upgrading of Tall Oil to Fuels and Chemicals Over HZSM-5 Catalyst Using Various Diluents. Can. J. Chem. Eng. 69, 1991. 5. Sharma, R.K. and Bakhshi,N.N. Catalytic Upgrading of Biomass-derived Oils to Transportation Fuels and Chemicals. Can. J. Chem. Eng. 69, 1991. 6. Wong, A. Tall Oil-based Cetane Enhancer for Diesel Fuel. P u b & Paper Canada, 96 (1 l), 37-40, 1995. 7. Stwnborg, M., Wong, A., Hogan, E. Hydroprocessed Vegetable Oils for Diesel Fuel Improvement. Bioresource Technol. 56, 13-18, 1996. 8. Liu, D.D.S., Monnier, J., Tourigny, G., W ,J., Hogan, E., Wong, A. Production of High Quality Cetane Enhancer from Depitched Tall Oil. Petr. Sci. & Tech. 16 (5&6), 597-609, 1998. 9. Severson, R.F., Schuller, W.H., Lawrence, R.V. Pyrolysis of Certain Resin Acids at 800°C. Journal of Chemical & Engineering Data 17 (2), 1972. 10, Severson, R.F., Schuller, W.H. The Thermal Behavior of Some Resin Acids at 400-500°C. Canadian Journal of Chemistry 50,2224,1972. 11. Takeda, H., Schuller, W.H., Lawrence, R.V. Thermal Behavior of Some Resin Acid Esters. Journal of Chemical & Engineering Data 13 (2), 1968. 12. Diesel Fuel Quality and Effects of Fuel Additives
1549
Preliminary Study on Fungicide and Sorption Effects of Fast Pyrolysis Liquids Used as Wood Preservative D. Meier', B. Andersons2,I. Irbe', J. Chirkova', 0. Faix' 'Federal Research Centrefor Forestry and Forest Products, Institutefor Wood Chemistry and Chemical Technology of Wood, 0 - 2 102 7 Hamburg, Germany 2 Latvian State Institute of Wood Chemistry, 27 Dzerbenes St., LV-1006 Riga, Latvia
Abstract: Pine wood samples were impregnated with fast pyrolysis liquid to study the fungicide and sorption effects. Three fungi were selected: Coniophora puteana, Poria placenta, and Lentinus lepideus. Impregnation was performed according to EN 113, and ageing procedures such as leaching and evaporation according to EN 84 and EN 73, respectively. The treated specimens without the ageing procedure did not suffer from fungal decay. After leaching, the samples lost approx. 5% of their initial weight. However, after evaporative ageing of the samples, the mass loss was below 1%. The alteration in the porous structure of the wood cell wall as a result of the impregnation and the ageing procedures, as well as after the action of fungi (Coniophora puteana and Lentinus lepideus) was investigated by the water vapours sorption method. The impregnation of wood with oil leads to the intensification of hydrophobic properties, which remain practicaly the same during evaporation (EN 73). As a result of leaching, the hydrophilic properties are restored incompletely, and the structure of the sample changes. Distinctions in the destructive action of Coniophora puteana and Lentinus lepideus have been established: from the viewpoint of structure, the activity of Lentinus lepideus relative to impregnated wood is more than twice that in the case of Coniophora puteana. After the leaching procedure (EN 84), the destructive action of Coniophora puteana increases dramatically, while the structural characteristics of leached samples upon contacting with Lentinus lepideus are not practically changed. Pyrolysis oil has shown a protective effectiveness against wood destroying basidiomycetes, so that it could be used as a wood preservative. However, some properties, such as stability in wood after ageing procedures, reduction of hydrophilicity and retention in wood, call for improvement.
1550
INTRODUCTION Creosote (tar oil) preservatives are used in industrial applications to treat ties, poles, posts, pilings etc. Creosotes and other tar oils are produced when a naturally occurring carbon-rich substance (coal, lignite or wood) is heated without air (1). The nature of creosote depends upon the type of coal used and the nature of the distillation processes. Typically, creosote is a complex mixture of over 200 compounds consisting mainly of polyatomic aromatic hydrocarbons (PAH) and phenolics. It may contain more than 30 different PAHs amounting up to 85 % (2). The wood-preservative properties of creosote depend upon many factors and, whilst the tar acids are good fungicides, they are also the components, which are most susceptible to loss from wood by volatilisation and leaching (3). Due to the relatively high toxicity of creosote based predominantly on the presence of benzo[a]pyrene (b[a]p), there exist special regulations and restrictions within the EU and in some member states such as Sweden, Denmark, Austria, Germany, and the Netherlands (4). In this study, the use of pyrolysis oil from wood was tested as an alternative wood preservative. In the initial laboratory evaluation, the toxic effects of the oil against wood destroying basidiomycetes were determined. The fungus Lentus lepideus was included as an especially resistant organism to tar oils.
MATERIALS AND METHODS Pyrolysis oil was obtained from pine wood fast pyrolysis in a fluidised bed reactor operated at Union Fenosa, Spain. Pine (Pinus sylvestris L.) sapwood blocks (20 x 20 x 5 mm) were vacuum impregnated with pyrolysis oil according to the modified European Standard EN 113 method (5), based on the agar wood block decay test. Additionally, ageing procedures such as the leaching (modified EN 84) (6) and evaporation (modified EN 73) (7) of impregnated wood blocks were carried out. The modifications of the Standard methods were based on the shortening of the test procedures. In all cases, the test duration was 6 weeks, the leaching of treated blocks lasted 2 weeks and evaporation 6 weeks before exposure to the fungal attack. A series of treated and control specimens was exposed to the brown rot basidiomycetes Coniophora puteana BAM Ebw. 15, Poria placenta FPRL 280 and Lentinus lepideus Ebw. 20 at 2 2 C and 70% RH. Treated check test specimens were placed in uninoculated culture vessels for calculation of the correction value. At the end of the test, the specimens were assessed and mass losses were calculated to determine any protective effectiveness of the given preparation. Changes in the structure and hydrophilic properties of the wood cell wall in specimens after the impregnation and ageing procedures samples were studied by the water vapours sorption method. The sorption-desorption isotherms measured on a vacuum sorption balance at 295 K were analysed by the comparative method in combination with the BET method (8). The accessible specific area A (m’/g), reflecting swelling capacity, the mass hydrophilicity a ( m M l g ) and the surface concentration of the hydrophilic centres a (groups/nm2)was determined. If the structure of the samples is identical, the comparative graphs are linear. The deviations from linearity indicate that the cell wall of the sample under study contains pores which are absent in the control sample. The sizes of such pores are calculated from the theory of capillary condensation (8), based on the location of the bend region corresponding to a definite PRO value on the isotherm. The bends on the comparative
1551
desorption-adsorption graphs unequivocally indicate the presence of pores of definite size in the sample.
RESULTS AND DISCUSSION EFFECTS ON FUNGI For each fungus and screening method, four replicates were taken and the mean values of results are shown in Table 1 - 3. The blocks were treated both with distilled water (0 concentration) and pyrolysis liquid in concentrated form. As can be seen from the Tables 1 3, uncorrected and corrected mass losses for waterimpregnated wood had a positive value and exceeded 30 % for each test fungus. On the contrary, uncorrected mass losses for oil-treated specimens showed negative values indicating an increase in mass. Negative values were obtained also for check (correction) test specimens. Since they were not placed on fungi, the mass variations of these specimens allowed determination of the changes in mass resulting from other factors than fungal attack. The corrected mass losses of the treated specimens were obtained in mathematical calculations and might show irrelevant theoretical values. These theoretical values are attributed to the results of the modified EN 113 method where, in fact, no mass losses of the treated blocks were observed. According to EN 113, especially for creosotes, the increase in the mass of the treated check test specimens may be very high. Therefore, the use of the correction values may produce erroneous theoretical losses in the mass of the treated specimens ( 5 ) . To avoid these misassessments, the treated specimens should also be carefully investigated for any visible attack by the fungus. Our investigations confirmed that the treated specimens without ageing procedure did not suffer from the fungal decay. The screening test with a leaching procedure revealed some loss of mass after the fungal attack. Oil-treated and leached specimens after attack by C. puteana, L. lepideus and P. placenta showed mass losses of 5.5%, 5.1% and 25.9%, respectively. Regarding the first two fungi, mass losses of about 5% seemed to be believable, because some visible signs were noticed after splitting of the specimens. The highest mass loss of P. placenta-attacked blocks evidently had an erroneous value. Some small traces of decay were noticed but they could not be as high as 25.9% and exceed the mass loss of the corresponding control specimens (24.2%). These results indicate that pyrolysis oil, is leachable to some extent, hence, it could promote further fungal decay. The screening test with an evaporative ageing of wood after the fungal attack indicated a negligible mass loss (less than 1.0%) i.e. O.O%, 0.1% and 0.8% for C. puteana, L. lepideus and P. placenta, respectively. A visual assessment of the treated blocks confirmed the calculated results. Wood blocks were intact and showed no signs of decay. Mass losses of the corresponding untreated control specimens reached acceptably high values (Tables 1 3). The virulence of fungal strains was controlled using untreated specimens in separate culture vessels. Mass losses of these specimens were 5 1.2%, 56.2% and 45.5% for C. puteana, L. lepideus and P. placenta, respectively. This is in-good agreement with the EN 113 demands where the minimum mass loss for virulence specimens should be 20%. Moisture contents of the treated and control specimens are shown in Tables 1 - 3 as the lowest and highest values of the specimen batch. Changes in moisture content were determined on a basis of the final dry mass. In general, specimens treated with water
-
-
1552
had a moisture content comparable with that of the control. On the contrary, wood treated with pyrolysis liquid had a comparatively high moisture content after-modified EN 113 and EN 1131 EN 73 tests, i.e. ca 80-120% depending on the fungus. These elevated hygroscopic properties resulted in an extra moistening of wood and the accumulation of liquid in the culture medium. The liquid around the specimens could affect the fungal activity. The screening tests with a leaching procedure presented different results. In this case, when a part of oil components was leached out, wood moisture did not exceed 60% depending on the fungus. The specimens were covered with a sparse mycelium and some loss in mass was observed. Retention of the oil in wood comprised approximately 315 kg/m’. This is a higher amount in comparison with creosote preservatives used at present, which are introduced into wood up to 200 kg/m for ground contact use or in a marine environment (9). Further investigations will be carried out to improve the properties of pyrolysis products for their potential use in wood preservation.
1553
P
m
c
0 preparation
0 preparation
EN 113/ EN 84
EN 113/ EN73 35.1 -34.0
37.1 -10.8
36.9 -36.5
Uncorrected mass loss, %
0.5
-0.1 -3 1.9
- 16.3
35.2 0.0
36.6 5.5
35.0 43.5
36.7 44.6
74.6-94.7 86.0-1 10.9
64.4-80.6 36.1-59.8
61.7-78.8 87.6-107.9
35.4 44.5
37.6 10.2* (0.0)
-0.7 -46.2
treated. Y O
%
%
%
control
82.5-87.7 65.5-73.0
68.5-82.6 75.9-90.8
64.1-73.9 62.0-74.2
Moisture
Mass loss of control,
Corrected mass loss,
Correction value,
*Irrelevant theoretical value due to the calculation of the correction value
0 preparation
Concentration, %(dm)
EN 113
Procedure
Table 1. Summary of results for Scots pine sapwood blocks treated with pyrolysis oil and attacked by the brown-rot fungus Coniophorn puteana.
c
ul ul
v1
0 preparation
0 preparation
EN 1131 EN 84
EN 1131 EN 73
35.2 -32.6
34.9 6.3
32.4 -25.3
Uncorrected mass loss, %
-0.1 -31.9
0.5
- 16,3
-0.7 -46.2
YO
Correction value,
35.3 0.8
34.4 25.9*
33.1 20.9* (0.0)
Corrected mass loss. %
31.8 36.4
24.8 24.2
24.1 33.4
Y O
Mass loss of control,
*Irrelevant theoretical value due to the calculation of the correction value
0 preparation
Concentration, %(dm)
EN 113
Procedure
72.1-116.4 70.6- 103.2
91 .O-122.0 46.7-54.8
68.8-97.1 79.9-102.9
65.4-81.2 63.4-69.4
61.9-103.8 5 1.7-67.4
71.8-84.9 65.3-78.9
Moisture % treated. control
Table 2. Summary of results for Scots pine sapwood blocks treated with pyrolysis oil and attacked by the brown-rot fungus Poria placenta.
Q\
vI
c VI
0 preparation
EN 1131 EN 73 -33.8
39.7
-1 1.2
33.1
40.6 -32.1
%
Uncorrected mass loss,
-0.1 -31.9
0.5 -16.3
-0.7 -46.2
Correction value, Yo
39.8 0.1
32.6 5.1
41.3 13.5* (0.0)
Corrected mass loss, %
38.6 42.6
36.4 29.5
40.3 38.3
Mass loss of control, %
*Irrelevant theoretical value due to the calculation of the correction value
0 preparation
preparation
0
% (dm)
Concentration,
EN 1131 EN84
EN 113
Procedure
60.0-75.6 85.7- 109.3
59.8-69.1 24.4-35.4
58.1-72.0 81.1-126.1
treated.
%
58.7-69.5 58.0-65.5
55.3-73.2 49.7-56.2
53.7-68.1 55.1-75.6
control
Moisture
Table 3. Summary of results for Scots pine sapwood blocks treated with pyrolysis oil and attacked by the brown-rot fungus Lentinus lepideus.
EFFECT ON SORPTION Some isotherms of water vapours sorption-desorption are presented in Fig. 1. In the first cycle of sorption, the isotherms have a pronounced stepwise character, and the values of water vapour sorption are unusually low for wood specimens. This is connected with the strict drying conditions of the specimens (378K at a rapid rise in temperature) after the exposure procedure, which resulted in the fixation of the nonequilibrium structure. After saturation with water vapours and the subsequent vacuum treatment at 295K during the procedure of the sorption experiment, the second cycle of isotherms was carried out, and this state was regarded as an equilibrium one. Thus, the analysis of the specimens structure was performed on the basis of the sorptiondesorption isotherms of the second cycle. In pine wood impregnated with water according to EN 113, pores with r > 2.1 nm are practically absent, and the filling of the pores proceeds without the capillary condensation, which is testified by the linearity of the comparative desorptionadsorption plots. As a result of impregnation with oil, in comparison with water impregnation, a dramatic (by 37%) decrease in the accessible surface and the mass concentration of the hydrophilic groups occurs (Table 4). Although the total volume of the pores tends to decline (from 0.15 to 0.12 cm3/g),a minor volume of pores in the range r = 2.1 - 4.8 nm appears (Fig. 2,A).
Table 4. Structural characteristics of pine wood specimens after impregnation with water and pyrolysis oil (EN 1 13), leaching (EN 84), and evaporation (EN 73)
Of interest is the comparison of specimens leached according to EN 84 after impregnation with oil with those impregnated with water according to EN 113. Such a comparison makes it possible to evaluate the degree of leaching of oil components. As can be seen from Table 4, an increase in A (by 30%) and a (by 28%) values of the oilimpregnated samples after leaching (EN 84) occurs. Therefore, a considerable part of the more light oil components is leached from the specimen with water. However, the structural characteristics do not reach the corresponding values for the blank specimens washed only with water, and the A and a values remain lower by about 20 %. In this
1557
case, a considerable change in the cell wall structure occurs, i.e. very narrow pores (radius about 0.5 nm) and rather a large pore volume in the radial range 1.0-1.3 nm tends to appear. This is unequivocally indicated by the bends in the comparative plots (Fig. 2,B). As a result of the blowing (EN 73) of pine wood impregnated with oil, no dramatic changes in the sorption properties and structure are observed (Table 4): only a minor increase in A occurs, and the surface becomes more hydrophilic as a result of the weathering of more volatile hydrocarbon components of oil. After the exposure of the impregnated wood specimens to C. puteana and L. lepideus, the hydrophilic properties of wood increase, i.e. A and a values increase by 6.9% and 9.4% in the case of C. puteana, respectively, and by 15.7% and 18.8% in the case of L. lepideus, respectively. The increase in a values in both cases indicates the appearance of new hydrophilic centres in wood. Obviously, it occurs owing to a partial destruction of wood substance by the microorganisms. From the above mentioned data it follows that the destructing action of microorganisms on wood impregnated with pyrolysis oils in the case of L. lepideus is almost twice as intensive as in the case of C. puteana. According to comparative plots, both cultures cause some decrease in the volume of the wider pores (Fig. 2,A). However, in the case of L. lepideus, in,contrast to C. puteana, a minor volume of pores in the radial range 1.5-2.1 nm appears. After exposure of impregnated and leached samples to C. puteana and L. lepideus, the picture is being changed: in the case of C. puteana, they exhibit a dramatic increase in A (by 38.5%) and a (by 42%), respectively. Moreover, some increase in a is observed. A considerable alteration in the structure, i.e. a new pore volume in the radial range 1.8-2.5 nm (mesopores) appears (Fig. 2,B, 3,A). Obviously, the increase in the accessible surface of the specimen impregnated with oil after leaching facilitates the destruction process upon the contact with this microorganism. On the contrary, in the case of L. lepideus, only minor changes occur in the A and a values (a decrease by 1.9% and 4.3%, respectively). The analysis of the structure indicates a decrease in the volume of wider pores and an increase in the pores volume in the radial range 1.3-1.5 nm (Fig. 2,B, 3,A). Hence, the oil components removed in the EN 84 procedure, inhibiting the considerable activity of L. lepideus relative to impregnated wood, play some role in their vital activity. The dramatic increase in the activity of C. puteana after leaching may be explained by facilitating the diffusion of the fungicide in wood cell wall. The specimens exposed to L. lepideus after the weathering procedure, (by EN 73) do not practically differ in terms of structural characteristics from the specimens not subjected to this procedure (Table 4). In the case of C. puteana, some growth in the hydrophilic properties is observed, mainly owing to the appearance of a new surface (the A value increases almost twice), which is important for this culture from the viewpoint of facilitating diffusion, as has been suggested earlier. As follows from Fig. 3,B, the specimen exposed to C. puteana is practically identical to the specimen before exposure, while, in the comparative plot of the specimen exposed to L. lepideus, there are bends indicative of a porous structure distinct from that of the non-exposed specimen.
1558
CONCLUSIONS The impregnation of wood with pyrolysis oils (by EN 113) decreases its hydrophilic properties mainly owing to the decrease in its swelling ability. The surface concentration of hydrophilic centres decreases negligibly. In the evaporation procedure (by EN 73), in which the light components of oil are removed, there is practically no the change of cell wall structure, moreover, the concentration of the surface hydrophilic groups somewhat increases. In the leaching process (EN 84), water-soluble oil components are removed, and the specific surface accessible to water as well as mass hydrophilicity increase. L. lepidius has a destructive action upon oil-impregnated wood, which is twice as high as that in the case of C. puteana, i.e. in particular, specific surface is increased by 7% and 16%, respectively. In this case, the hydrophilic properties of the surface are changed not only at the expense of the loosening of the structure, but also owing to the increase in the surface concentration of the hydrophilic centres. The activity of C. pureann relative to the sample after the leaching procedure tends to increase dramatically, i.e. specific surface and mass hydrophility tend to increase by 40% at some simultaneous increase in the concentration of the surface hydrophilic centres. At the same time, the activity of L. Iepideus relative to this sample is negligible, i.e. specific surface is increased only by 2%. The activity of both the cultures relative to the sample subjected to the evaporation procedure (EN 73) are close, i.e. the destructive action of L. lepideus is only by 10% stronger, in comparison with the case of C. pureana. Pyrolysis oil has shown a protective effectiveness against wood-destroying basidiomycetes, so that it has a potential for use as a wood preservative. However, some properties of the oil need to be improved, particularly, the stability in wood after ageing procedures, the reduction of a high hydrophilicity and a high retention of the oil in wood.
I559
A
120
160
Q 140 100 120
60
-
100
--B w I
I
m
m
-
80
m
m
80
40
40 20
0
0
0.2
0,4
0,6
0,8
1
0
0,2
PPO
c
0,6
0,4
0.8
PIP0
D
140
200
160
?
120 160
100
-a
-B
1
140
-a
60
120
g 100
I
m
60
60 80
40
40
20 20 0
0
0
0,2
0,4
0,6
0,8
0
1
0,2
0,6
0,4
PIP,
PIP,
Fig. 1. Isotherms of water vapours sorption-desorption at 295K on pine wood after the procedure: A - water (EN 113); B - oil (EN 1 13); C - oil (EN 113) / C. puteana; D - oil (EN 113/EN 84) / C. puteana (bold line - 1'' cycle; labels - 2"dcycle: empty labels - sorption; filled labels - desorption)
1560
0,8
1
A
I2O I00
80
--F cn
-
60
m
40
20
0
20
10
30
50
40
70
60
80
90
B 160 140 120
0 1
0
20
40
60
80
100
120
140
a (mug)
Fig. 2. Comparative plots of the desorption vs. sorption on the specimen: “oil (EN 113)” (+) and “oil (EN 113) - C. puteana” (A)(A); “oil (EN 113/EN 84)” (+) and “oil (EN 113IEN 84) - C. puteczno” (A)(B)
1561
A 160 140 120
-
=i 100
ul E 9
80
60 40 20
0 0
20
40
80
60
100
120
a (mglg)
0
10
20
M
40
50
60
70
80
90
100
a (m&)
Fig. 3. Comparative plots of the water vapours desorption ’on the wood specimens: “oil (EN 113/EN 84 - C. puteana)” (A)and “oil (EN 113EN 84) - L. lepideus.” (+) to “oil (EN 113/EN 84) (A); “oil (EN 113/EN 73) -C. puteana” (A)and “oil (EN 1131EN 73) L. Zepideus.” (+) to “oil (EN 113/EN 73) (B)
1562
REFERENCES 1. 2.
3. 4. 5.
6. 7.
8.
9.
Milton F.T. (1995) The Preservation of Wood. Minnesota Extension Service. Official Journal of the European Communities, October 26,1999. Richardson B.A. (1 978) Wood Preservation. The Construction Press Ltd, Lancaster, England. EU Directive 94/60/EG. European Standard EN 113: 1996. Wood Preservatives. Method of Test for Determining the Protective Effectiveness against Wood Destroying Basidiomycetes. CEN, Brussels. European Standard EN 84: 1990. Wood Preservatives. Accelerated Ageing of Treated Wood Prior to Biological Testing: Leaching Procedure. CEN, Brussels. European Standard prEN 73: 1987. Wood Preservatives. Accelerated Ageing of Treated Wood Prior to Biological Testing: Evaporative Ageing Procedure. CEN, Brussels. Gregg, S.J. & Sing K.S.W. (1982) Adsorption, Surface Area and Porosity, Academic Press, London. Willeitner H. & Liese W. (1992) Wood Protection in Tropical Countries: a Manual on the Know-How. Deutsche GTZ GmbH.
1563
Fractional Vacuum Pyrolysis of Biomass for High Yields of Phenolic Compounds H. Pakdel, J. N. Munvanashyaka and C. Roy Department of Chemical Engineering, Universite' Laval, Sainte - Foy, Que'bec Canada G l K 7P4
ABSTRACT: Phenolic compounds with potential industrial applications are major pyrolysis products of biomass lignin. Fractional pyrolysis allowed the evolution of chemical compounds such as phenols to be monitored as a finction of temperature. Due to the less restrained pyrolysis conditions, phenolic compounds were produced in high yields by fractional pyrolysis in five steps. The pyrolysis oils were analysed and the phenolic compounds were quantified. The feedstock moisture and volatile components evolved during the first step below 200 "C. The pyrolysis oil obtained in the second step (200-275 "C) was composed of oxygenated compounds which were mainly produced by the degradation of cellulose and hemicelluloses. Low molecular weight carboxylic acids and aldehyde derivatives were abundant in the second pyrolysis step; they tended to polymerize with phenolic compounds if they were not quickly remo.ved from the reactor hot zone. The pyrolysis oil concentrated with guaiacol and syringol-derivatives was obtained by the rupturing of p-0-4 (guaiacylglycerol-P-arylether) bonds during the third step in the temperature range of 275-350 "C. As the pyrolysis temperature increased, ArO-CH3 bonds cleaved to form alkyl phenols and diols in the 350-450 "C temperature range in the fourth step. Polycyclic aromatic hydrocarbons were produced at temperatures above 450 OC by rupturing of the Ar-0 bonds in the final step. Methanol detected in the final step was formed by demethoxylation reactions.
INTRODUCTION In wood pyrolysis, it is known that several parameters influence the yield of pyrolytic oil and its composition. Among these parameters, wood composition, heating rate, pressure, moisture content, presence of catalyst, particle size and combined effects of these variables are known to be important'. The thermal degradation of wood starts with free water evaporation. This endothermic process takes place at 120 to 15OoC,followed by several exothermic reactions at 200 to 25OoC, 280 to 32OoC, and around 4OO0C, corresponding to the thermal degradation of hemicelluloses, cellulose, and lignin respectively2. In addition to the extractives, the biomass pyrolytic liquid product represents a proportional combination of pyrolysates from cellulose, hemicelluloses, 1564
and lignin when these compounds are separately pyrolysed3. Pyrolysis of different wood components produces different classes of compounds. Cellulose for example, produces acetic acid, gases, water, sugars and occasionally a small amount of furans and phenols4. Hemicelluloses on the other hand principally produce acids, furans and sugars. Lignin is biosynthetically constructed by copolymerization of phenylpropanoid monomers namely coumaryl-, coniferyl- and sinapyl alcohols5 which is principally decomposed to phenols and aromatic hydrocarbons during pyrolysis6. In addition to pyrolysates, free organic compounds or extractives present in wood are also released during pyrolysis, making pyrolytic oils a very complex product. Pyrolysis under vacuum of aspen poplar wood chips exhibited an active zone of decomposition occurring between 240 and 300°C'. A 58 % by wt. yield of pyrolytic oil was obtained at 350OC which contained miscellaneous compounds such as phenols, carboxylic acids, alcohols and esters7. The vacuum process removes the reaction products before they are further decomposed'. The yield of relatively volatile phenols (monolignols in this paper) depends on the pyrolysis materials. The yield of phenols varied from 1 to 3.7 % when different biomass samples were subjected to vacuum pyrolysis'. Phenolic compounds typically possess anti-diarrheal and anti-motility properties', germicidal activity", herbicidal effect and antiseptic properties". They have also been used in the tanning of leatheri2, as dyes13, as a thermal insulating materialI4, as food aroma and liquid smoke15. In some cases, the smoke flavour has been attributed to the presence of phenols and sometimes to a single component. The phenols forming the "liquid smoke" which is used to smoke foods are well studied but their role and extent of flavouring have been clarified recently16. The present study was undertaken to determine the evolution of the main monolignols at different temperatures by fractional vacuum pyrolysis of birch wood.
MATERIAL AND METHODS
Birch (Betula papyrifera) wood chips used in this work were obtained from Scierie John Lewis, a woodmill in the province of Quebec, Canada. The sample was composed of 53.5 YOof sapwood and 46.5 % of inner and outer bark. The moisture content was 8.5 % as determined by placing a known quantity of sample in an oven at 102 f 3°C untiI a constant weight was reached. The elemental composition of the sample was 5 1.3 % C, 6.2 YOH, 0.7 % N and 41.8 YO0 (by difference). Its ash content was 2.0 %. FRACTIONAL VACUUM PYROLYSIS A 800 g sample of the milled wood was pyrolysed in a batch reactor under vacuum (run # G72). A detailed description of the batch pyrolysis reactor system used in this work
has been described elsewhere17. Two vacuum pumps in series were used to achieve a total pressure of 0.7 kPa and three dry-ice-in-limonene condensers (-72°C) were used to trap the pyrolysis vapours. When the maximum pyrolysis temperature was reached, it was held for 1 hour prior to cooling to room temperature. After the pyrolysis in each step, the system was kept under nitrogen until the next pyrolysis step was started. Each pyrolysis step was carried out by using the solid residue from the previous step. Table 1 shows the different pyrolysis steps and product yields. The pyrolysis oils which were
1565
recovered in the different traps were stored in sealed vials in the refrigerator (4°C) for M e r chemical analysis.
Table I Fractional pyrolysis conditions and yields ( run # G72). Pyrolysis step
5
Temperature range ("C) 30-200 200-275 275-350 350-450 450-550
Total One step
25-550
# 1
2 3 4
Oil yield (wt. %, anhydrous wood basis)
6.02 18.14 27.28 9.47 1.48
62.39 63.43
PYROLYSIS OIL CHEMICAL CHARACTERIZATION
The identification and quantification of the phenols were carried out using gas chromatograph coupled to a mass spectrometer (GCMS) and using standard phenols. The pyrolysis oil was acetylated before analysis. A 200 to 300 mg sample of well mixed pyrolytic oil, together with 5 mL of acetic anhydride and two drops of pyridine, were heated in a sealed vial in a water bath at 60°C for 90 minutes. The derivatized solution was then eluted over approximately 2 g of silica gel with 100 ml of 80 % CH2C12 in hexane solution. The eluate was concentrated under vacuum. A 100 pl of a solution of 25 mg anthracene dissolved in 10 ml of ethylacetate was added to the pre-concentrated solution as an internal standard. Three sets of derivatized standard phenols diluted in 1:1, 1:2 and I :3 volumes of ethylacetate were used for the determination of the absolute response factors of acetylated phenols with respect to anthracene for quantitative analysis. The GCMS analysis was performed on a HP-5890 gas chromatograph with split injection at 290°C. The column was a 30 m x 0.25 mm i.d. HP5-MS fused silica capillary with 0.25 pm film thickness from Hewlett Packard. Helium was the carrier gas with a flow rate of about 1 ml min-'. The GC initial oven temperature was held at 50°C for 2 min, then programmed to increase to 2 10°C at 5°C min-' and then to 250°C at 10°C min-'. The oven temperature was held at 250°C for 10 min. The'end of the column was introduced directly into the ion source of a HP-5970 series quadruple mass selective detector. The transfer line was set at 270°C and the mass spectrometer ion source was set at 250°C with 70 eV ionization potential. A volume of 1 pl of sample was injected into the GC using a HP-7673 automatic sampler. Data acquisition was done with a PC base G1034C Chemstation software and a NBS library data base. The mass range of m/z = 30-350 Dalton was scanned every second. The identification of targeted compounds were based on NBS library mach values typically over 50 %. Furthermore, the identifications were confirmed by matching their mass spectra and retention times with the standard authentic compounds.
1566
RESULTS AND DISCUSSION
FRACTIONAL PYROLYSIS Compositional analysis of the pyrolysis oils at various steps allowed five distinct pyrolysis steps to be established in the temperature range of 200-550 "Cas illustrated in Figure1. All five oil hctions were analysed by GCMS. In the first step, about 6.02 YO volatile compounds, which contained 99.9 % moisture, evolved at a temperature lower than 200 "C. That step involved mainly the sample drying process and partially the evaporation of wood natural extractives. Volatiles, hydrophilic and lipophylic compounds like terpenes and carboxylic acids have been detected earlier during the drying of fresh wood at about 200°C'8. Table 2 lists the main components which were produced in the first step.
30 25
s
20
-$
s 15 0
s 10 5
0 30-200
200-275
275-350
350-450
450-550
Temperature OC Fig. I Product distribution following fractional pyrolysis (wt. %, anhydrous wood basis). The pyrolysis oil collected between 200 to 275°C in the second step was 18. I4 % by wt. on an anhydrous wood basis and corresponded to the hemicelluloses degradation". In general, hemicelluloses decompose to yield firans and its derivatives as well as a series of aliphatic carboxylic acids. Hardwoods are composed of pentosanbased hemicelluloses, while hexosan-based hemicelluloses are found in softwoods. Hardwood produces more low molecular weight carboxylic acids and furans than softwood upon decomposition". In contrast, hardwood contains a lower quaptity of fatty acids than softwood. C1 - C6 carboxylic acids, fatty acids (e.g. CIS), firan and derivatives were identified in this fkaction. In the second pyrolysis step, a little rearrangement of the lignin polymer occurred which was based on the formation of small amount of phenolic compounds ( 0.395 % by wt., anhydrous wood basis). The main components of this fraction are listed in Table 2. Betulin and lupeol were abundant in this fraction. Furthermore, levoglucosan which has been reported earlier as a major compound in the cellulose-derived pyrolysis oils', was identified in the second 1567
and third fractions. Upon further heating, the fragmentation is followed by dehydration, disproportionation, decarboxylation and decarbonylation reactions. Several phenolic compounds as well as aromatic and cyclic hydrocarbons, aliphatic and cyclic alcohols, ketones, aldehydes, acids, esters and furans associated with the recondensation reaction of cellulose degradation products have been reported earlier4. A pyrolysis liquid yield of 27.28% was obtained in the third step in the temperature range of 275350°C. This step corresponds mainly to the lignin decomposition where C-C bonds cleaved to form radicals, followed by recombination reactions. A 2.176 wt. % mixture of volatile phenols on an anhydrous initial feed basis was identified in this fraction. Guaiacyl and syringyl compounds like guaiacol, 4methylguaicol, syringol, syringol, 4-methylsyringol, isoeugenol, allyl- and propenylsyringol were the main phenolic compounds. The guaiacol and syringol concentration leveled off at about 37OOC and 350°C respectively. Catechol and its derivatives (catechol, 3-methyl- and 4-methylcatechol, 3-methoxycatechol and resolcinol) were also identified in this fraction. The main components of this fraction are listed in Table 2. The oil yield in the fourth step represented 9.47 wt. YOof the initial anhydrous sample. In the fourth step, 1.845 wt % of volatile phenols on an anhydrous initial feed basis were produced. The content of catechol and its isomers were high, being the most abundant compounds, whereas the guaicyl and syringyl content were low in the fourth step. The main components of this fraction are listed in Table 3. C-0 bond cleavage of the remaining lignin residue in the temperature range of 450-550 OC, mainly resulted in the formation of some polycyclic aromatic hydrocarbons (PAHs) as listed in Table 3. The total yield of hydrocarbons produced in the final step was 1.48 wt. % on an initial anhydrous sample basis. The residual phenolic compounds represented < 0.1 % by weight in the final step. The principal components of this fraction are listed in Table 3. Interestingly, the contamination of pyrolysis oil with PAHs can be reduced without losing a significant mass of the pyrolysis oil if the pyrolysis reaction is stopped at approximately 450 OC. Table 2 Pyrolysis oil composition as a function of pyrolysis temperature. 25- 200 OC Acetic acid Hydroxypropanone Pentanol Hexanal 3-furaldehyde Hexan- 1-01 Pentanoic acid Hexan-3-01 Ethylfuraldehyde Hexanoic acid Furanone Methyldihydrohran Heptanoic acid trans-linaloloxide Methy lheptadienone Heptanol
275 - 350 OC Hydroxytetrahydrofuran Propanol 3-methylpropanol Pentanoic acid 4-meth y lcyclopentanone 3-furaldehyde 2-firaldehyde 2-furanmethanol 5-methyl furanone Tetrahydro-2,Sdimethoxyhran 2-hydroxycyclopent-2-en-1-one Cyclopentanedione Methylfuraldehyde 2-c yclopentenone Methylmethoxypropenoate Furanone Hexanoic acid 2-hydroxycyclopenten-1-one Dimethylfuranone Methylfiranone
200-275 OC Hydroxytetrahydrofiran Hy droxypropanone Hexanal Furaldehyde Acetic acid Furaldehyde Dihydromethylpyran Propanone Pentanoic acid Hexanoic acid (SH)-furan-2-one
1568
25- 200 "C Octanol 1,4-dimethoxybenzene Heptan-3-one Hexacosane Heptacosane Methylbetulinate Methy loleanolate
200-275 "C 3-methylcyclopentanedione Dimethylhydroxyfuran-3-one Guaiacol Maltol Octan-3-one 4-hydroxybenzeneethanol 4-methylguaiacol Catechol 1,6-anhydro-a-Dglucopyranose 4-(2-propenyl)phenol 3-methoxycatechol 4-ethy lguaiacol 4-methylcatechol 5-acetoxymethy lfuraldehyde Syringol Eugenol Vani 11in Cis-isoeugenol 4-methylsyringol trans-isoeugenol 4-propylguaiacol Acetovanillone Levoglucosan 2-propiovanillone Hydroxypheny lbutanone 4-allylsyringol cis-4-propeny lsyringol Syringaldehyde Trans-4-propeny lsyringol 1,2-dimethoxybenzene Trihydrox yphen y lpentanone Hexadecanoic acid Heneicosane 9,12-octadecadienoic acid Octadecanoic acid Docosane Tricosane Tetracosane Pentacosane Hexacosane Heptacosane Octacosane Nonacosane Lupeol Betulin
1569
275 - 350 "C 5-Methylfuraldehyde Phenol 1,2-cyclohexanedione 3-meth ylcyclopentanedione Dimethylcyclopentenone o-cresol m-cresol 2,5dimethy lhydroxyfuranone Guaiacol 2-2-eth y lcyclohexanone ' Maltol 3ethyl hydroxycyclopentenone 2,5dimethy lhydroxyfuranone 4-ethylphenol Octanoic acid 4-methylguaiacol Catechol 3-methoxycatechol Naphthalene 4-ethylguaiacol Hydroquinone 4-methylcatechol 1,2,4-benzenetriol 1-(3methoxypheny1)ethanone Syringol Eugenol 3,4-dimethoxyphenol 4-propy lguaiacol 4-ethylcatechol I ,2,3-benzenetriol Vanillin 4-ethylsyringol Levoglucosan I -(2,6-dihydroxy-4methoxypheny1)ethanone 4-ally lsyringol cis-4-propeny lsyringol Syringaldehyde trans-4-propenylsyringol 1-(4-hydroxy-3,5-
dimethoxypheny1)ethanone (2,4,6-trihydroxy-3methylpheny1)butanone Hexadecanoic acid
25- 200 OC
200-275 "C
275 - 350 "C Octadecadienoic acid Methyloctadece-9-noate
Table 3 Pyrolysis oil composition as a hnction of pyrolysis temperature.
350 - 450 OC 350 - 450 "C Toluene Dihydroindenone Octene Tridecene Octane 4-methylcatechol 2-methylpropanol 1-methylnaphthalene Ethylbenzene Tridecane 1,2-dimethylbenzene 2-methylnaphthalene 1-(2-hydroxy-5Styrene Nonene methylpheny1)ethanone Nonane 1,3-dirnethoxybenzene Methylcyclopentenone 2,3-dihydroindenol Ethylmethylbenzene Syringol Phenol 4-methylcatechol Decene 3,4-dimethoxyphenol Trimethylbenzene Decanoic acid Benzofuran 4-ethylcatechol Decane Tetradecane 4-methylguaiacol 1,4-dimethylnaphthalene Trimethylbenzene 7-dimethylnaphthalene lndene 4-methylsyringol o-cresol 1,2-dimethylnaphthalene m-cresol 1,6-dimethyInaphthalene Hexanoic acid Pentadecene 2-ethyl- 1,3Pentadecane Dimethylbenzene Trimethylnaphthalene Guaiacol Hexadecene, Hexadecane Undecane 2-methylnaphthalenol 2-methylbenzofuran Heptadecene 2-ethylphenol Heptadecane Dimethylcyclohexanedi 1 -methyl-7-( 1one methylethy1)naphthalene 2,4-dimethylphenol Otadecene 1-methylindene Fluorenol 2,5-dimethylphenol Octadecene 4-ethylphenol Octadecane Octanoic acid Dihydromethylphenylbenz1,4-dimethoxybenzene ofurane Dodecene Nonadecene Methy lguaiacol Nonadecane Dodecane Hexadecanoic acid Catechol Eicosene
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450 - 550 "C Dimethylnaphthalene Pentadecene 2-methylbiphenyle Acenaphthene Pentadecane 4-methylbiphenyl Dibenzohran 1,4,6-trimethylnaphthalene 1,6,7-trimethylnaphthalene Hexadecene Fluorene Dimethylbiphenyl Fluorenol 2-methy lfluorene o-hydroxybiphenyl Azulene Ethyldimethylmethylfluorene Dimethylethylazulene Phenanthrene Anthracene 2-phenylnaphthalene 2-methylanthracene 1-phenylindene 2-phenylnaphthalene Eicosane 3,6-dimethylphenanthrene Pyrene Heneicosene Heneicosane Fluoranthene Benzo[a]fluorene Benzo[c]fluorene Methylpyrene o-terphenyle Triphenylene Benz[a]anthracene Dihydromethylbenz[alanthracene
350 - 450 "C 350 - 450 "C Trimethylphenol Eicosane Dimethylbenzohran Heneicosene 4-ethyl-2-methylphenol Heneicosane 4-propy lphenol 9,12-octadecadienoicacid 3-methylcatechol Tetradecane 2,3,6-trimethylphenoI 2-methylanthracene Octanoic acid 2-phenylnaphthalene 4-ethylguaiacol
450 - 550 "C
EVOLUTlON OF THE PHENOLIC COMPOUNDS
At a pyrolysis temperature lower than 200 "C no phenolic compounds were detected. During the second pyrolysis step, 0.395 % phenolic compounds (anhydrous feedstock basis) including phenol and cresol (low quantities), allylphenol, guaiacol, 4ethylguaiacol, eugenol, isoeugenol, vanillin, ethylvanillin, syringol, 4-methy1, 4-ally1 and propenylsyringol were detected (Table 4). The presence of these phenols indicated the initiation of lignin degradation and disproportionation reactions following the drying of the wood in the first step. Similar compounds have been reported earlier during pyrolysis of a lignin model compound in the 230-260°C temperature range2'. These compounds indicate the rupturing of bridge C-C bonds in aromatic compounds or P-ether or P-arylether bonds (p-0-4). These two former bonds possess a low energy of dissociation in the range of 217-235 kJ/mo122.The dissociation of P-ether bonds was accelerated as the pyrolysis temperature exceeded 300 "C and guaiacyl compounds were formed. The formation of guaiacol and derivatives ended at ca. 400°C or below. The evolution of phenols reached a maximum in the third step in the temperature range of 275-350 "C (see Table 4). Normally, the natural lignin contains no catechol unit, which leads to the conclusion that catechols are the secondary and guaiacol is the intermediate pyrolysis products which were formed at high temperatures. P e t r ~ c e l l iemphasized ~~ that the secondary decomposition reactions of guaiacol to catechol at high temperature were favoured at a long holding time. The cleavage of the 0-C bond in methoxy groups of guaiacol at high temperature reduced the guaiacol and increased the dihydroxybenzene content. In this work, catechol was detected at about 300°C and its concentration was increased considerably and ended at 500"C, the temperature at which the pyrolysis was almost over. Pyrolysis of 4-ethylguaiacol yielded 4-ethylphenol by cleavage of the 0-C (alkyl) and 0-C (aryl) bonds24.Similarly, methyl-, dimethyl- and vinylphenols originated from guaiacol intermediates. It has also been reported that phenol is a secondary pyrolysis product of guaiacol and derivativesz3. Phenol, cresol, ethylphenol, dimethylphenol, propylphenol and methylcatechol were the main components of the oil fraction in the fourth step. Traces of hydrocarbons such as toluene and styrene were produced by the dehydroxylation of phenols in this fraction. The maximum phenolic compounds which could be produced and identified was about 2.5 wt. % on an anhydrous feedstock basis during the one-step pyrolysis approach in the temperature range of 25-550 "C. Interestingly, 4.4 wt. % phenols were identified following the fractional pyrolysis approach in a lower temperature range of 25-450 "C. One can conclude that the removal of volatile phenols during the fractional approach prevents some secondary thermal decomposition reactions which occur during the one-
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step pyrolysis process. These results are in agreement with the data reported earlier by Vohler and Schweers". According to Avni et a1.26,at a low pyrolysis temperature, the decomposition of lignin leads to the release of C02, water and hydrocarbon gases from aliphatic and methoxy side chain groups. Aldehyde functional groups tend to decompose to CO. At high temperatures, thermal cracking reactions and rearrangements of lignin sub-units lead to the evolution of H2 from aromatic hydrogen as well as additional CO from the tightly bonded oxygen functionalities such as diary1 ether and phenols in lignin. This information reveals that lignin is very sensitive to temperature. The lignin hnctional groups are easily decomposed or modified and C-C and C-0 bonds were cleaved within the wide temperature range of 200 to 400°C. However, it is believed that a pyrolysis process can still be designed to efficiently degrade the lignin polymeric structure into simple phenols without modifying its original structure by preventing dehydroxylation, decarbonylation, demethoxylation and dehydrogenation reactions which occur during the thermal degradation process. Catechol and its derivatives and alkylphenols were very abundant in the fourth step due to the decomposition of the primary phenols by demethylation reactions which require a higher dissociation energy (about 356-414 kJ/mole) than the P-arylether bonds. Ceylan and Bredenberg*' have shown that 29-40% of guaiacol could be transformed into catechol after 2 h of pyrolysis at 305-345 "C. Similarly, 4ethylguaiacol has been shown to transform to 4-ethylguaiacol by pyrolysis28. In rather restrained pyrolysis conditions, phenolic compounds produced at low temperatures were decomposed as they were exposed to high temperatures in the onestep pyrolysis (Table 4). Moreover, phenolic compounds in the acidic medium polymerize with aldehydes to form resins. The result was a progressive decline in the phenolic content. Under both fractional and one-step pyrolysis conditions, the oil yield was about the same (Table 1). Figure 2 illustrates the evolution of both total phenols and pyrolysis water as a function of temperature. The evolution of water at temperatures above 200 OC was indicative of the dehydration reaction of oxygenated compounds during pyrolysis. The abundance of phenolic compounds declined during . the dehydration reaction at temperatures above 350 "C. A similar result has been reported earlier in the literature29. J
*
200
250
300
350
400
450
500
24
550
Temperature ( O C )
Fig. 2 Evolution of phenols and water as a function of pyrolysis temperature.
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Table 4 Evolution of phenols as a function of pyrolysis temperature ("C). Compound
Yield on an anhydrous feed basis (wt. %) 200-275 275-350 350-450 450-550 Total Phenol 0.082 0.004 0.023 0.1 13 0.004 o-cresol nd. nd. 0.080 0.002 0.082 nd. m-cresol nd. 0.052 0.00 1 0.053 p-cresol 0.026 0.002 0.001 0.102 0.072 2,5-xylenol nd. 0.01 1 nd. 0.000 0.012 2,3-xylenol nd. nd. 0.028 0.001 0.029 Xylenol nd. nd. 0.017 nd. 0.0 17 4-ethylphenol nd. nd. 0.035 nd. 0.035 Guaiacol 0.25 1 0.017 0.04 1 nd. 0.309 nd. nd. 0.010 3,5-xylenol nd. 0.0 10 4-allylphenol 0.010 nd. nd. 0.0 10 nd. 0.108 nd. 4-methy lguaiacol 0.01 1 0.097 0.001 0.01 1 0.197 0.539 0.001 Catechol 0.33 1 nd. nd. 0.064 0.064 nd. Hydroquinone 0.024 0.078 nd. 0.048 Resorcinol 0.005 0.1 16 nd. nd. 4-ethylguaiacol 0.027 0.089 0.014 0.197 nd. 0.005 0.178 4-methylcatechol 0.425 nd. 0.025 Syringol 0.03 1 0.370 0.066 0.200 0.270 nd. 4-methylcatechol 0.005 nd. 0.012 nd. nd. 0.012 Vanillin nd. 0.058 nd. 0.058 nd. Methy lresorcinol nd. 0.075 nd. Dimethylcatechol nd. 0.075 0.105 nd. nd. 0.0 17 0.088 Dimethylcatechol 0.184 nd. 0.014 0.143 0.028 4-methy lsyringol 0.185 nd. nd. 0.1 10 0.075 Methoxycatechol 0.063 nd. nd. 0.063 Dimethylresorcinol nd. 0.0 19 nd. nd. 0.0 19 Dimethylresorcinol nd. nd. 0.020 0.0 12 0.101 0.069 Isoeugenol nd. 0.01 1 0.057 0.085 0.017 4-ethylsyringol nd. 0.098 nd. 0.067 0.03 1 Propylguaiacol 0.195 nd. 0.128 0.067 1,2,4-benzenetriol nd. nd. 0.132 0.1 19 nd. 0.012 Allylsyringol nd. 0.1 17 0.0 17 0.100 nd. Syringaldehyde 0.171 nd. 0.06 1 0.1 10 1,3,5-benzenetriol nd. 0.195 nd. 0.065 0.130 nd. Propenylsyringol nd. 0.063 0.056 0.007 Methoxyresorcinol nd. Total *nd. : not detected.
0.395
2.176
1.845
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0.010
4.426
25-550 0.151 0.089 0.064 0.144 0.0 14 0.050 Nd. 0.054 0.182 Nd. Nd. 0.107 0.246 Nd. 0.034 0.090 0.103 0.1 13 0.108 Nd. Nd. 0.027 0.069 0.089 Nd. Nd. Nd. 0.069 0.029 0.067 0.034 0.093 0.048 0.273 0.161 nd. 12.504
CONCLUSION
The thermal decomposition of wood under vacuum in a fractional pyrolysis approach can simplifL the pyrolysis oil composition and facilitate the separation and purification of phenols produced during the pyrolysis. The evolution of pyrolysis water and phenols by fractional vacuum pyrolysis of birch wood bark exhibited a similar trend. The major volatile phenolic compounds were obtained in the temperature range of 200 to 350 "C following free radical C-0 and C-C bond cleavage. A pyrolysis temperature higher than 350 "C was detrimental for the recovery of syringyl compounds which tended to decompose to catechol and its derivatives. Guaiacol and its derivatives were produced at a low temperature in the range of 250-350 "C followed by syringol and its derivatives. The transformation of the methoxy group to catechol was accelerated at temperatures higher than about 350 "C. Finally, the aromatic hydrocarbons were produced by dehydroxylation and demethoxylation reactions at temperatures higher than 450 "C. Under vacuum, the maximum yield of volatile phenolic compounds with 4.4 wt. % was achieved in the temperature range of 270-400°C following a fractional pyrolysis approach. REFERENCES
1. Pakdel H., Roy C. & Lu X. (1996). Effect of Various Pyrolysis Parameters on the Production of Phenols from Biomass. In : Developments in Thermochemical Biomass Conversion. ( Ed. by A. V. Bridgwater & D.G .B. Boocock), pp.509-522. Blackie Academic & Professional, London. 2. Fengel D. & Wegener G. (Eds) (1984) Wood: Chemistry, Ultrastructure and Reactions. Walter de Gruyter, Berlin. 3. Elder T. J . (199 1) Pyrolysis of Wood. In : Wood and Cellulosic Chemistry. ( Ed. by D. N. S. Hon & N. M. Shiraishi), pp. 665-702. Dekker Inc. New York. 4. Russell J.M., Miller R.K & Molton P. (1983) Formation of Aromatic Compounds from Condensation Reactions of Cellulose Degradation Products. Biomass, 3, 4357. 5 . Monties B: (1989) Lignins. In : Methods in Plant Biochemistry. Vol. 1 (Ed. by J. B. Harbome and M.P. Dey), pp. 113-157. Academic Press. 6 . Alen R., Kuoppala E. & Oesch P. (1996) Formation of the Main Degradation Compound Groups from Wood and Its Composition During Pyrolysis. J. Anal. Appl. Pyrolysis, 36, 137-148. 7. Roy C., de Caumia B., Brouillard D. & Menard, H. (1985) The Pyrolysis of Aspen Poplar. In : Fundamentals of Thermochemical Conversion. (Ed. by R. P. Overand, T. A. Milne & L. K. Mudge), pp. 237-255. Elsevier Applied Science Publication, New York. 8. Shafizadeh F. (1984) The Chemistry of Pyrolysis and Combustion. In: The Chemistry of Solid Wood,.(Ed. by R Rowell), chapter 13. ACS, Washington D.C. 9. Ogata N. B. T. & Shibata T. (1993) Demonstration of Antidiarrheal and Antimotility Effects of Wood Creosote. Phalmacol. 46, 173-180. 10. Guha G., D. Das P.D. Grover, & Guha, B.K. (1987) Germicidal Activity of Tar Distillate Obtained from Pyrolysis of Rice Husk. Biol. Wastes. 21, 93-100. 11. Azhar L.P., Levin E.D. & Sokolova, N.A. (1972) Antiseptic Properties of Oil Obtained from Lignin Sedimentation Tar. Zzv. VUZ, Lesnoi Zh. 15 , 1 18-20.
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12. Ratner M.E., Smetanina S.S., Kovalev V.E. & Korotova, O.A. (1979) Preparation of Syntans Based Phenols of Soluble Resin and Wood-resin Pyrolysate. Izv. Vyssh. Uchebn. Zaved., Lesn Zh, 72-76. 13. Rijkuris A., Biseniece S. & Sergeeva, V.N. (1978) Composition and Properties of Tar Formed During Thermolysis of Lignocellulose. Preparation of Azo Dyes from Thermolysis Tar. Khim. Drev, 1:68, CA. 90: 153480. 14. Garalevicius R., Medzevicius V., Roshchupkin V.I., Faintsimer R.Z., Yakolvev, D.A. & Romanauskas, A. (1 978) Composition for Preparing a Heat-Insulating Material. USSR Patent SU 339173. 15. Guillen M.D. & Ibargoitia M.L. (1998) New Components with Potential Antioxidant and Organoleptic Properties, Detected for the First Time in Liquid Smoke Flavoring Preparations. J. Agric.and Food Chem. 46, 1276 - 1285. 16. Wittkowski R., Ruther J., Drinda, H. & Rafiei-Taghanaki F. (1992) Formation of Smoke Flavor Compounds by Thermal Lignin Degradation. In: Flavor PrecursorsThermal and Enzymatic Conversions. (Ed. R. Teranishi, G. R. Takeoka & M. Guntert), pp. 232-243. ACS symposium series 490, American Chemical Society, Washington DC. 17. Pakdel H., Couture G. & Roy, C. (1994) Vacuum Pyrolysis of Bark Residues and Primary Sludges. Tappi J.,77,205-2 1 1. 18. Fagernas L. (1993) Formation and Behaviour of Organic Compounds in Biomass Dryers. Bioresource Technology, 46,7 1-76. 19. Goldstein 1.S . ( 1980) Organic Chemicals from Cellulose. In: Organic Chemicals from Biomass. (Ed. I. S. Goldstein), pp. 101-124. CRC Press, Boca Raton. 20. Sarkanen K.V., Connors W.J., Johanson L.N. & Winslow, P. (1980) Thermal Degradation of Krafi Lignin in Tetralin. Holzforschung, 34,29-3 1. 21. Fiddler W., Doer R. C., Wasserman A. E. & Parker W.E. (1967) Thermal Decomposition of Ferulic Acid. J. Agric. Food Chem., 15, 757-761. 22, Benson S.W. (Ed) (1976) Thermochemical Kinetics, 2nd edn. Wiley, New York. 23. Petrocelli P. F. & Klein T. M. (1985) Simulation of Krafi Lignin Pyrolysis. In : Fundamantals of Thermochemical Biomass Conversion. (Eds R.P. Overend, T.A. Milne & L.K. Mudge), pp. 257-273. Elsevier Applied Science Publisher, London. 24. Connors W.J., Johanson L.N., Sarkanen K.V. & Winslow P. (1980) Thermal Degradation of Krafi Lignin in Tetralin. Holzforschung: 34,29-37. 25. Vohler W. & Schweers H.M. W. (1975) Utilization of Phenol Lignin. Appl. Polym. Symp., 28,277-274. 26. Avni E., Suib L.S. & Coughlin W.R. (1985) Free Radical Formation in Lignin During Pyrolysis. Holzforschung, 39, 33-40. 27. Ceylan R. & Bredenberg J.B. (1982) Hydrogenolysis and Hydrocraking of the Carbon-Carbon Bond. 2. Thermal Cleavage of the Carbon-Oxygen Bond in Guaiacol. Fuel, 61, 377-382. 28. Masuku, S.P. (1991). Thermal Reactions of Bonds in Lignin. IV. Thermolysis of Dimethoxyphenols. Holzforschung, 45, 181- 190. 29. Samolada M.C., Stoikos T. & Vasalos I.A. (1990) An Investigation of the Fictors Controlling the Pyrolysis Product Yield of Greek Wood Biomass in a Fluidised Bed. J. Anal. Appl. Pyrolysis, 18, 127-14 1.
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ACKNOWLEGEMENTS
The authors would like to thank Dr. Annette Schwerdtfeger for her technical assistance in the preparation of this manuscript.
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Production of Hydrogen from Biomass-Derived Liquids S. Czernik, R. French, C. Feik, and E. Chornet Chemistryfor BioEnergy Systems Center, National Renewable Energy Laboratory, 161 7 Cole Boulevard, Golden, CO 80401
ABSTRACT Renewable biomass can become an attractive feedstock for producing hydrogen especially because of its essentially zero net C02impact. Unfortunately, hydrogen content in lignocellulosic biomass is only 6-6.5% compared to almost 25% in natural gas. Therefore, on a cost basis, producing hydrogen by the biomass gasificatiodwater-gas shift process cannot compete with the well-developed technology for steam reforming of natural gas. However, an integrated process, in whch biomass is partly used to generate hydrogen and partly to produce more valuable materials or chemicals can be an economically viable option. In the presented concept hydrogen is produced from lignocellulosic biomass by pyrolysidsteam reforming or by steam-aqueous fractionatiodsteam reforming processes that also generate valuable co-products. Effluents from other biomass processing technologies such as transesterification of vegetable oils can also be attractive feedstocks for the production of hydrogen. T h s work focused on the second stage of the process, catalytic steam reforming of various biomass-derived liquids. We employed a fluidized bed reactor configuration with commercial nickel catalysts developed for processing natural gas and naphtha. The hydrogen yields obtained approached or exceeded 90% of the values possible for stoichiometric conversion. INTRODUCTION Hydrogen is the most environmentally friendly fuel that can be efficiently used for power generation. At present, hydrogen is produced almost entirely from fossil fuels such as natural gas, naphtha, and inexpensive coal. In such case, the same amount of C 0 2 as that formed from combustion of those fuels is released during the hydrogen production stage. Renewable biomass is an attractive alternative to fossil feedstocks because of essentially zero net C0 2 impact. However, the hydrogen content of lignocellulosic biomass is only 6-6.5%, compared to almost 25% in natural gas, and on a cost basis, producing hydrogen by a direct conversion process such as gasificatiodwater-gas shift cannot compete with the well-developed technology for steam reforming of natural gas. Vegetable oils have a better potential for producing hydrogen than lignocellulosic materials but their high costs make the process economics not favorable either. Only an integrated process, in which biomass is partly used to produce more valuable materials or chemicals with the residual fractions
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utilized for generation of hydrogen, can be an economically viable option. The concept of our approach to producing hydrogen from biomass is shown in Figure 1.
Catalytic steam reforming
Hz
Water gas shift
c02
+
Steam
7 4
w
Transesterification
Food processing
Plant and animal fats
Figure I. Biomass to hydrogen - Process concept
In earlier papers’” we proposed a method, which combines two stages: fast pyrolysis of biomass to generate bio-oil and catalytic steam reforming of the bio-oil to hydrogen and carbon dioxide. This concept has several advantages over the traditional gasification technology. First, bio-oil is much easier to transport than solid biomass and therefore, pyrolysis and reforming can be carried out at different locations to improve the economics. For instance, a series of small-size pyrolysis units could be constructed at the sites where low cost feedstock is available then the oil transported to a central reforming plant located at a site with existing hydrogen storage and distribution infrastructure. The second advantage is the potential for production and recovery of higher-value co-products from bio-oil that could significantly impact the economics of the entire process. The lignin-derived fraction can be separated from bio-oil and used as a phenol substitute in phenol-formaldehyde adhesives4or converted
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to cyclohexyl ethers (fuel additives5) whlle the carbohydrate-derived fraction is catalytically steam reformed to produce hydrogen. An alternative concept for producing hydrogen that we are developing combines steam-aqueous fractionation of biomass with catalytic steam reforming of the lowervalue hemicelldose-rich liquid by-product whlle cellulose and lignin components would be used for other applications. The concept can be extended to steam reform other biomass-derived liquids that are by-products from existing technologies. In this work we studied “crude glycerin”, a by-product from a bio-diesel plant based on transesterificationof vegetable oils. In previous years we demonstrated, initially through micro-scale tests’ then in bench-scale fixed-bed reactor experiment^^*^ that bio-oil model compounds as well as bio-oil carbohydrate-derived fraction can be efficiently converted to hydrogen. Using commercial nickel catalysts the hydrogen yields obtained approached or exceeded 90% of those possible for stoichiometric conversion. The carbohydrate-derived bio-oil fraction contains substantial amounts of non-volatile compounds (sugars, oligomers) which tend to decompose thermally and carbonize before contacting the steam reforming catalyst. Even with the large excess of steam used, the carbonaceous deposits on the catalyst and in the reactor freeboard limited the fKed-bed reforming time to 3-4 hours. The limitations of the fned-bed reactor were even more obvious for processing the whole bio-oil or the hemicellulose-rich solution from steam-aqueous fractionation. The hydrogen yield obtained from the whole oil was only 41% of that stoichlometrically possible and the reforming duration was less than 45 minutes. For this reason we decided to employ a fluidized bed reactor configuration that can overcome some limitations of the fixed-bed unit for this application. Though carbonization cannot be completely avoided, the bulk of the fluidizing catalyst remains in contact with the liquid droplets fed to the reactor.
EXPERIMENTAL MATERIALS Bio-oil used for t h ~ sstudy was generated from pine sawdust using the NREL fast pyrolysis vortex reactor system6. The oil comprised 47.7% carbon, 7.4% hydrogen, and 44.8% oxygen with water content of 26.7%. It was separated into aqueous (carbohydrate-derived)and organic (lignin-derived) fractions by adding water to the oil in a weight ratio of 2: 1. The aqueous fraction (75% of the whole oil) contained 21.8% organics (CH1.2500.55)and 78.2% water. Steam-aqueous fractionation of poplar wood was performed at the University of Sherbrooke, Canada using a continuous Stake I1 unit7. This treatment led to solubilization, after washing, of 30% of the biomass into a hemicellulose-rich aqueous solution, which contained 32.1% of solutes of the elemental formula CH1.3600.67. These solutes were mostly oligomeric pentosan and a small amount of dissolved ligtw. “Crude glycerin” samples were obtained from West Central Co-op bio-diesel plant in Ralston, Iowa. Transesterification of vegetable oils with methanol produces a mixture of bio-diesel (methyl esters of fatty acids) and glycerin. Glycerin settles down at the bottom of a separation tank while bio-diesel forms the top layer. “Crude glycerin“ is a very viscous liquid, only partially miscible with water. Its elemental composition includes 54.7% carbon, 9.9% hydrogen, and 35.5% oxygen, which suggests that the phase separation in the tank was not very clean and the liquid is a
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mixture of glycerin (55%) with methyl esters of fatty acids (45%). C1 1-NK, a commercial nickel-based catalyst used for steam reforming of natural gas and naphtha, was obtained from United Catalysts and ground to the particle size of 300-500~. FLUIDIZED BED REFORMER The bench-scale fluidized bed reactor is shown in Figure 2. The two-inch-diameter inconel reactor supplied with a porous metal distribution plate was placed inside a three-zone electric furnace. The reactor contained 150-200g of commercial nickelbased catalyst ground to a particle size of 300-500~.Superheated steam was used as fluidizing gas as well as a reactant in the reforming process. Steam was generated in a boiler and superheated to 750°C before entering the reactor at a flow rate of 2-4 g/min. Liquids were fed at a rate of 4-5 g/rnin using a diaphragm pump. A specially designed injection nozzle supplied with a cooling jacket was used to spray liquids into the catalyst bed. The temperature in the injector was controlled by coolant flow and maintained below the feed boiling point to prevent evaporation of volatile and deposition of nonvolatile components. HeatcExchangers 2” Fluid
Reactor Mass Flow Controller Mass Flc
I Orifice Flow
- Liquid Pump
Collecto Collector andscale
&per heater
Meter
1
MTI Gas Chromatograph
Mass Flo Controlle
Figure 2. Fluidized bed reformer system The product gas passed through a cyclone that captured fine catalyst particles and any char generated in the reactor, then two heat exchangers to remove excess steam. The condensate was collected in a vessel whose weight was continuously monitored. The outlet gas flow rate was measured by a mass flow meter and by a dry test meter. The concentrations of COz, CO, and CH4 in the reforming gas composition were monitored by a non-dispersive *a-red analyzer (NDR Model 300 from California Analytxal
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Instruments) and that of hydrogen by a thermal conductivity monitor TCM4 manufactured by Gerhard Wagner. In addition, the gas was analyzed every 5 minutes by an on-line MTI gas chromatograph, which provided concentrations of hydrogen, carbon monoxide, carbon dioxide, methane, ethylene, and nitrogen as a function of time of the test. The temperatures in the system as well as the flows were recorded and controlled by the OPT0 data acquisition and control system. Total and elemental balances were calculated as well as the yield of hydrogen generated from the feed. RESULTS AND DISCUSSION
BIOMASS PYROLYSIS OIL - AQUEOUS FRACTION The steam reforming experiments on aqueous extract of the pine bio-oil were camed out at the temperature of 800°C and 850°C. The steam to carbon ratio was in the range of 7-9 whle methane-equivalent gas hourly space velocity GclHSV was 1200-1500 h-'. During the experiments at 800°C a slow decrease in the concentration of hydrogen and carbon dioxide and an increase of carbon monoxide and methane in the gas generated by steam reforming of the carbohydrate-derived oil fraction was observed. These changes resulted from a gradual loss of the catalyst activity, probably due to coke deposits. As a consequence of that, the yield of hydrogen produced from the oil fraction decreased from the initial value of 85% of stoichiometric (3.24 g of hydrogen from 100 g of feed) to 77% after 12 hours on stream. If a water-gas shift reactor followed the reformer the hydrogen yields would be 94% and 84% respectively. At 850°C the formation of char and coke was much lower or their gasification by steam was more efficient than that at 800°C. During over 90 hours of uninterrupted reforming of the bio-oil carbohydrate-derived fraction, the product gas composition remained almost constant and only a small decrease in the concentration of hydrogen was observed, as presented in Figure 3. 80 70
60 50
30 -~ 20 10
~
-
07
0
1000
2000
3000
4000
5000
6000
Time, min
Figure 3. Reforming of bio-oil aqueous extract. Gas composition as a function of process time.
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HEMICELLULOSE-RICH AQUEOUS SOLUTION
The steam reforming experiments were carried out at the temperature of 800-850°C with the hemicellulose solution feed rate of 240-300 g/h and steam flow of 140-180 g/h. This corresponds to the methane equivalent space velocity Gc,VHSV of 1000 h" and the molar steam to carbon ratio S/C of 9.5-14. The product gas composition is shown in Figure 4. For over five hours the reforming gas composition was almost constant. However, after 2.5 hours on stream at 800°C the amount of gas generated started to decline. Consequently, the hydrogen yield, which at the beginning reached 85% of the stoichiometric potential decreased to 55% (Figure 5). Increasing the temperature to 850°C and the steam flow (S/C=14.2) resulted in the improvement of the catalyst performance. 80 , 1 70
H2
-
60 -
I
50
""
I
I
0
100
50
150
200
250
300
350
Time, min
Figure 4. Reforming of hemicellulose aqueous solution. Gas composition as a function of process time. 100
95
90
4.-
85
0
c
cn $
5 0
j;
80 75
70
8
65
60 1
55
v 1 w - ! A 7
50
0
50
100
150
200
250
300
350
Time, rnin
Figure 5. Yield of hydrogen from steam reforming of hemicellulose solution.
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Mass balances indicated that 85% of carbon from hemicellulose was converted to C02and CO in the first phase of the experiments. The remaining 15% could thus form char entrained from the system and coke deposits on the catalyst surface, which would explain the loss of its activity. The activity of the catalyst used for reforming was restored by steam gasification of the deposits and the catalyst was reused in the next experiments showing almost the same efficiency. Though the carbon to gas conversion reached again 85%, the hydrogen yield was in the range of 75-80% (85% initially) of the stoichiometric value and more methane was observed in the produced gas than during the tests using fresh catalyst. This indicates some permanent loss of the catalyst activity. In general, hemicellulose is more difficult to reform than bio-oil aqueous fraction due to a higher content of oligomeric material that tends to carbonize during the process. Therefore, it requires a higher steam to carbon ratio to reduce carbon deposits on the catalyst.
CRUDE GLYCERIN FROM BIODIESEL PRODUCTION “Crude glycerin” is a high viscosity liquid and, therefore, it had to be preheated to facilitate pumping and atomizing (the whole feeding line was maintained at 60-80°C). The liquid was fed at a rate of 78 g/h (GclVHSV = 1600 h-’) and steam at a rate of 145 g/h, which corresponds to the molar steam to carbon ratio of 2.3. The experiments proceeded very smoothly with only occasional fluctuations in the liquid feed rate resulting from a non-complete homogeneity of the feed. The concentration of the major gas products was constant during the run time but a gradual increase in methane production was noticed (Figure 6). The process performance measured as the yield of hydrogen did not decrease signtficantly during several hours on stream. The overall mass balance closure was close to 100% at the beginning of the tests then decreased to 9596% after four hours on stream. Similar closure was also observed for elemental balances of carbon, hydrogen and oxygen. 2500
I
500 00
50
1 50
100
200
250
300
Time, min
Figure 6. Methane concentration during steam reforming of “crude glycerin”
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The hydrogen yield oscillated around 77% of the stoichometric potential, whch was 23.6 g per 100 g of feed. It could be sigmficantly higher if more steam were used in the reaction. Conversion of CO in the gas through water-gas shift to C02 and H2 would increase the hydrogen yield to 95% of that theoretically possible. These promising results suggest that a low-value by-product from bio-diesel production could become a viable renewable raw material for producing hydrogen. An integration of these two technologies could significantly improve the economics of both processes. SUMMARY AND CONCLUSIONS
1. Biomass can be a valuable resource for producing hydrogen if utilized in an integrated process that also generates higher value co-products. Following t h s strategy we have considered three process options: fast pyrolysis/steam reforming, steam-aqueous fractionatiodsteam reforming, and transesterification of vegetable oils/steam reforming. 2. Bio-oil from pyrolysis or its aqueous, carbohydrate-derived fraction, hemicellulose-rich solution from steam-aqueous fractionation, and glycerin from bio-diesel production can be catalytically steam reformed to generate hydrogen using commercial nickel-based catalysts. 3. The hydrogen yield obtained in a fluidized bed reactor from the aqueous fraction of bio-oil was about 85% of the stoichometric value, which corresponds to almost 6 kg of hydrogen from 100 kg of wood. 4. Hydrogen yield from the hemicellulose solution was about 70% of the stoichiometric potential. Lower performance was due to the higher content of oligomeric material, which is more difficult to reform. 5. The hydrogen yield from “crude glycerin“ was 18 g per 100 g of the feedstock, which corresponds to 76% of the stoichiometric potential. If the steam reforming were followed by a water-gas shift process or a higher amount of steam were used in the reforming stage, the hydrogen yield could increase above 90%. 6. Catalysts are readily regenerated by steam or C02 gasification of carbonaceous deposits. 7. The process needs to be optimized to determine conditions that allow for maximum yields of hydrogen and minimum coke formation.
REFERENCES Wang, D., Czemik, S.,MontanC, D., Mann, M., and Chornet, E. (1997) Biomass to hydrogen via pyrolysis and catalytic steam reforming of the pyrolysis oil and its fractions. I&EC Research, 36, 1507-1518. Wang, D.; Czernik, S., and Chornet, E. (1998) Production of hydrogen from biomass by catalytic steam reforming of fast pyrolysis oils. Energy & Fuels, 12, 19-24. Czernik, S., French, R., Felk, C, and Chornet, E. (1999) Fluidized bed catalytic reforming of pyrolysis oils for production of hydrogen. In: Proceedings of the Fourth Biomass Conference of the Americas, (Ed. by R.P. Overend and E. Chornet), pp. 827-832. Elsevier Science Ltd., Oxford. Kelley, S. S., Wang, X.-M., Myers, M. D., Johnson, D. K., Scahill, J. W. (1997) Use of biomass pyrolysis oils for preparation of modified phenol formaldehyde resins. In: Development in Thermochemical Biomass Conversion, (Ed. By
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Bridgwater, A.V. and Boocock, D.G.B.), pp. 557-572. Blackie Academic & Professional, London. 5 . Shabtai, J.S., Zmierczak, W., and Chornet, E. (1997) Conv6rsion of lignin to reformulated gasoline compositions. In: Proceedings of the Third Biomass Conference of the Americas, (Ed. by Overend, R.P. and Chornet, E.), pp. 10371040. Elsevier Science Ltd., Oxford. 6 . Diebold, J.; Scahill, J. (1988) Production of primary pyrolysis oils in a vortex reactor. In: Pyrolysis Oils from Biomass: Producing, Analyzing and Upgrading; (Ed. by Soltes, E. J. and Milne, T. A.), pp. 31-40. American Chemical Society, Washington, D.C. 7. Heitz, M., Capeck-MCnard, E., Koeberle, P.G., GagnC, J., Chomet, E., Overend, R.P., Taylor, J.D., and Yu,E. (1991) Bioresource Technology, 35,23-32.
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Co-Firing of Bio-Oil with Simultaneous SO, and NO, Reduction R.H. Venderbosch', B.M. Wagenaar', E. Gansekoele', S . Sotirchos2, H.D.T. Moss3 I BTG Biomass Technology Group BV, P.O. Box 217, 7500 AE Enschede, The Netherlands Foundation for Research and Technology - Hellas, Institute of chemical engineering and high temperature chemical processes, P.0. Box 1414, GR - 265 00 Patras, Greece DynaMotive Europe Ltd., 33 Waterloo Road, Bedfort, UKGB - MK40 3PQ, United Kingdom
ABSTRACT Calcium-enriched bio-oil (CEB) can be used for flue gas desulfurisation in coal and waste combustion chambers. It is produced by mixing biomass derived fast pyrolysis oil with calcium oxide. The aim of the proposed project is to develop a technology i) to produce calcium-enriched bio-oil with a calcium content of 13 wt.%, and ii) to test the CEB in a combustion chamber by co-firing it with a sufur-containing fuel. In tlus paper the production method of CEB will be elucidated, and small-scale experiments related to CEB spraymg will be presented. Finally, co-combustion experiments of a sulfur-containing fuel with CEB in a small flame tunnel (20 kW*) will be reported. INTRODUCTION Industrialised nations impose strict limits on emissions of pollutants such as SO, and NO,. Guidelines referred to by the World Bank Pollution Prevention and Abatement Handbook for fossil fuel based thermal plants of 50 MW, or greater, aim at a 40% emissions reduction in NO,. In the USA, for example, the 1990 Clean Air Act Amendments require a 2-stage drop in the level of pollution. It aims at a reduction in SO2 emissions for coal and oil fired boilers by 10 million tons and for NO, by 2 million tons by the year 2000 (vs 1980 levels). National standards for SO2 became increasingly stringent throughout the 1980's and 1990's: a range of national emission standards for sulfur emissions in almost 40 countries is given by Soud (3). Technologies to control emissions of SO2 and NO, include flue gas desulphurisation (FGD) for SO2 control and NO, emissions abatement and control by primary measures, or by flue gas treatment. Flue gas desulfurization can be classified into the following six main categories:
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wet scrubbers; spray dry scrubbers; 0 sorbent injection processes; 0 dry scrubbers; regenerable processes; and 0 combined SO2 / NO, removal processes An overview of these technologies can be found on the website of the IEA Coal Research reviews (1) on FGD technologies (including low cost retrofit options) and from the reviews provided by Soud (2, 3). In the present paper more details will be presented about sorbent injection techniques. 0
SORBENT INJECTION Depending on the location (and temperature) where the sorbent is injected, there are four types of injection techniques: furnace sorbent injection (750-1250OC); economiser sorbent injection (30O-65O0C); duct sorbent injection (approx. 15OOC); and hybrid sorbent injection (a combination of two or more of the above). The simplest flue gas desulfurisation technology is furnace injection, where a dry sorbent is injected into the upper part of the furnace to react with the SOz in the flue gas. The finely grained sorbent is distributed quickly and evenly over the entire cross section in the upper part of the furnace in a location where the temperature is in the range of 750-1250°C. Commercially available and cheap limestone (CaC03) or hydrated lime (Ca(OH)2) is used as sorbent. The sorbent reacts with SO2 and O2 to form CaS04, Below 750OC the reaction rate is too low. At temperatures over 1250°C the surface of the sorbent will be sintered, and the structure of the pores will be destructed, reducing the active surface area. The major part of SOz-removal takes place w i h 1 to 2 seconds. For Ca(OH)2, removal efficiency of up to 50% are reported at a Ca:S ratio of 2. For CaC03, the removal efficiency will be even lower. Fine sorbent particle size (-3pm) and an even distribution of the sorbent over the cross-section of a boiler significantly improves the process performance. Bjerle and others (1993), for example, reported results of laboratory scale experiments showing that SO2-removal efficiencies > 95% can be acheved with furnace sorbent injection with particle size <3 pm. At lower temperatures hydrated lime, Ca(OH)2, can be injected into the flue gas stream near the economiser zone (300-650°C). In this temperature range, CaC03 can be formed, which is undesirable because it not only consumes sorbent but pore closure also blocks the access of SO2 to the active sorbent surface. Carbonation significantly increases with reaction temperature, and therefore, the flue gas duct process where the temperature is about 15OoC, may be more effective. Th~sprocess yields S02-removal efficiencies of approx. 80% in actual commercial installations if small particles with an open pore structure are applied. The aim of the present work is the development of a Calcium Enriched Bio-oil (CEB) that it is suitable to be co-fired with a sulfur containing fuel in a combustor to simultaneously reduce SO2 and NO,. Where the capture of NO, is important, CEB has greater flexibility. Additionally, CEB will reduce the higher N20-concentrations associated with this technology. The data below have been abstracted from the IEA Coal Research publication entitled “Air Pollution Control Costs for Coal-Fired Power Stations” (1995). 1587
Table I Investment and operating costs for various de-SO, and de-NO, processes Technology Wet FGD Dry FGD Selective Cat. Red. Selective Non Cat. Red. Gas reburn Low NO, burner (LNB) Overfire air (OFA) LNB - OFA CEB (7.3% Ca) CEB (1 1.O% Ca)
Capital cost US$/kWe
Avg 15yr cost US$/t
180 - 260 140 - 170 100 - 150 10 - 25 14- 18 20 30 10 - 20 30 - 45 2-5 2-5
460 - 580 430 - 480 1800 - 2000 926 - 1352 1200 - 2400 629 - 889 439 592 - 1065 726 562
-
SOz%
Red.
NO,% Red.
90 90
90 90
80 30 - 50 60 30 - 60 10 - 25 40 - 60 40 40
Table 1 demonstrates that the use of CEB can be of commercial interest. It may provide N0,/S02 control at costs well below those of conventional emission control systems and imply significant savings on up-front capital expenditure. Additionally, it has the advantage of reducing the production of greenhouse gases by substitution of fossil fuels. CALCIUM ENRICHED BIO-OIL CEB is produced by reacting biomass fast pyrolysis oil with CaO. The desulphurization technology is applicable both for coal and for heavy fuel oil fired combustion chambers. The installed coal-fired power station capacity in Europe approximates 75,000 MW,, but this is decreasing. The sulphur content of the coal ranges from 0.5 to 3.0 wt.% depending on the mining site. On the other hand, waste and orimulsion firing is increasing in capacity.
PREVIOUS WORK First experiments using CEB were carried out in a 90 k W h Down-Fired Coal Combustor at Penn State (4, 5, 6, 7) These experiments showed that it can be sprayed with slightly modified spray nozzles. Atomisation of a CEB-like material yielded droplets < 40 pm for 70 to 75 wt.% of the material, with all droplets < 170 pm. A 5 wt.% Ca-containing material was tested in the combustion chamber, showing a significant SOz and NO, reduction potential. The SOzreduction appeared to be a strong function of the temperature of the combustor, and values of up to 94 % were reported at a Ca:S ratio of 1. The optimum wall temperature for SO2 capture was observed to be around 1150°C, above which the performance dropped. The NO, reduction appeared to be a strong function of the CO concentration, which in turn depends on the oxygen concentration. The location of the introduction of additional air for complete combustion was an important factor to balance NO, reduction and to minimise CO emissions. A reduction in NO, levels of up to 44 % was observed at Ca:S -1.
PRESENT WORK In the present work, CEB with a calcium content of 13 wt.% (higher than in the 1588
previous work referred to above) will be tested as a flue gas desulhsation agent in a small (25 kW*) and a large combustion (7 MWm)facility. The work discussed in h s paper focuses on the demonstration of CEB, and includes: - Production of bio-oil and optimisation of the oil; - Production of CEB and its characterisation, and - Demonstration of CEB to remove SO2 in a small scale 25 kWh flame tunnel.
SMALL SCALE PRODUCTION OF BIO-OIL AND OPTIMISATION Upon heating biomass to about 650°C in the absence of oxygen, bio-oil is produced with a possible yield of > 75 wt.% of the feedstock material. This oil contains high concentrations of aldehyde- and carboxyl-groups, whch can react with lime to produce the CEB. The calcium uptake of the bio-oil is proportional with the aldehyde and carboxyl concentration of the bio-oil. Both are dependent on the pyrolysis process conditions. The first specific objective was to maximise the amount of aldehyde and carboxyl groups in the bio-oil, and to produce sufficient and reproducible amounts of bio-oil under the best optimal process conditions. The optimisation of the bio-oil was aimed at adjusting the operating conditions of the fast pyrolysis process to maximize the concentration of these reactive groups in the bio-oil, whle maintaining a hgh overall oil yield. Figure 1 shows the main part of the small-scale production facility for small amounts of the oil. It consists of the biomass feeder, the reactor and a bio-oil collection system. Before each experiment, a batch of sand was preheated inside an electrical furnace to about 600"C, where after it is mixed with cold sawdust in the bottom of the cone. The produced vapours could be immediately removed from the hot reactor, and collected in several water-cooled vessels.
I
Cyclone
BioOil
recei-
Figure I A schematic representation of the small-scale pyrolysis plant.
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Bio-oil can react with lime to Calcium Enriched Bio-oil (CEB) according to: carboxylic acids + Ca(OH)2 + Ca-carboxylates + Ca- (hydr0xy)phenoxides 0 phenols + Ca(OH)2 + Ca-carboxylates + alcohols aldehydes + Ca(OH)2 + Ca-polyol (aldol) complexes 0 aldehydes + Ca(OH)2 + Ca-polyol (ketal) complexes 0 ketones + Ca(OH)2 + Ca-carboxylates + alcohol 0 esters + Ca(OH)2 0 carbohydrates + Ca(OH)2 + Ca-polyol complexes The optimum conditions for these reactions is depending on the type of reaction taking place, and an experimental procedure much be undertaken to investigate the ‘best’ CEB preparation conditions. It seems reasonable to define the oil quality in terms of the functional groups present, mainly acids, ketones and aldehydes, and phenols. The ketodaldehyde and acid groups can be easily measured by techniques derived for liquid smoke. Details have been presented elsewhere (12). The results showed that the keton content increased with temperatures, while the acid content only slightly decreased. This indicates that higher pyrolysis temperatures are beneficial for the bio-oil to be used in the CEB. Fortunately, in the temperature interval of 450 to 6OOOC the oil yield is almost independent of the temperature. The operating temperature at which the bio-oil should be produced for CEB production is up to 600OC.At 6OOOC and in this small-scale setup, oil yields were up to 70 wt.% (on wood as received). For straw, the oil yields were significant lower in comparison (approx. up to 60%), most probably due to the higher ash-content of straw in comparison with wood, catalyzing secondary cracking reactions.
0
SMALL-SCALE CEB PRODUCTION
A stirred tank reactor enabling continuous measuring of pH, temperature and torque (at pre-set stirrer speeds), was used to produce the first small samples of CEB. Oil was slowly added to slaked lime at various operating conditions, viz. the reaction temperature, the stirrer speed, the total reaction time,the bio-oil addtion rate and the bio-oiVslaked lime ratio. In a typical CEB production experiment it was shown that the point of addition of bio-oil was clearly marked by a significant rise in temperature and drop in pH. The experiments demonstrated that slaked lime had a high capacity for biooil, and that calciumhydroxide could be added far in excess of the theoretical acid content without the pH increasing above 12. This indicates that ‘excess’ calciumhydroxide reacts further with the other hctional groups in the bio-oil. The CEB product was assessed visually (by colour), pH and water content, and settled for several days.. Several batches of 20 kg, comprising a mixture of both the upper and lower fraction, were produced to be used later in the small-scale combustion trials. Characteristics of the CEB are presented in Table 2. BENCH-SCALE INVESTIGATION OF CEB DECOMPOSITION /SULFA TION SO2 removal experiments with CEB were performed in a thermogravimetric analysis (TGA) set-up to simulate the process during its combustion. The sample was first heated in a mixture of 70% C02 in N2 to produce CaC03, followed by the decomposition of CaC03 to CaO under N2 and 02.This calcined material was sulfated using a mixture containing 3000 ppm SO2, 12% 02,and the balance N2.
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Table 2 Parameter Bio-oil fraction Water content PH Calcium content Viscosity
CEB characteristics value up to 50 up to 40 11-12 13 20°C - 3430 50°C - 1900 100OC- 1290 1210 up to 10
wt.% wt.%
wt.%
CP kg/m3 MJkg
Density Caloric value
/
IAt
--a-
GreerLimesbne Iceland Spar
Figure 2 Comparison of the behavior of CEB with 60 pm Greer Limestone and Iceland Spar at 750°C Although not shown in this paper, at 750 OC almost complete sulfation of the CEB material was observed, whereas conversions at 850°C were lower. Obviously, the temperature during the decomposition has a very strong influence on the performance of the sorbent. It may be that exposure to high temperatures increases the rate of sintering, which reduces the porosity and the surface area of the sorbent. The sulfation at 750°C is compared to that of similarly treated limestone particles (see Figure 2: open symbols). A strong decrease in weight at the start of the experiments with CaC03 can be clearly noted, which is due to the calination of CaC03 to CaO. For the limestone material, the decrease in weight due to the calcination reaction is not compensated anymore by sulfation. The CEB material exhibited a higher sulfation conversion than calcium carbonate solids particles. Shown in Figure 2 by the closed symbols is the increase in weight of the samples after sulfation of the calcined material. For complete 1591
conversion of Ca(OH)2 to CaS04, (l+AW/Wo) = 1.84, and for CaC03, (l+AW/Wo) = 1.36. To establish if the effect of the temperature on the sulfation is due to changes in the pore structure, the CEB was analysed with mercury porosimetry and nitrogen sorption. Results indicated that pores in the material treated at 75OOC had a smaller surface area and volume than in the materials treated at 850°C. The latter has more pore volume for pores close to 100 nm. However, on the basis of the sorption results, one would expect the material treated at 850°C to show a higher overall reactivity. This is characterized not only by a higher porosity but also by a larger fraction of small pores and a larger surface area. One explanation for the lower conversion for the material treated at 85OOC can be that the large pores produced at these conditions are accessible only through pores of smaller size. This is demonstrated also by the fact that a large fraction of mercury is trapped within the pore structure, and relatively broad hysteresis loops are encountered in the nitrogen sorption results. At 110°C, CEB forms a structure with a very high porosity of approx. 90-95%. This material already exhibited pores in the size order of a few hundreds microns, resulting in the so-called 'popcorn-effect' yielding a larger effective surface area. Because plugging of the pores is the main cause for the incomplete utilization of CaO-based materials, the high porosity of CEB decomposed material is a clear indication of the potential of CEB as a d e s u l h a t i o n sorbent. PRODUCTION OF LARGE QUANTITIES OF BIO-OIL
Several tonnes of wood derived bio-oils are required for the production of sufficient quantities of CEB for the combustion experiments in a small (25 kWh) and in a large (7MWm) combustion facility. These oils are produced in a 200 kg/hr pilot plant, shown in Figure 3. Details of the set-up are presented elsewhere (8).
Flue gas
@ ,
s1
E2
I Oil
Figure 3 A schematic flow diagram of the proposed 2 t/hr fast pyrolysis system
The pyrolyser is a patented rotating cone reactor. Biomass particles (#1) are fed near the bottom of the pyrolysis reactor (Rl) together with an excess flow of hot heat carrier material such as sand, where it is being pyrolysed at 500 to 600OC. The 1592
produced vapours (# 14) pass through several cyclones (not shown) before entering the condenser (Cl), in which the vapours are quenched by re-circulated oil (# 18). The pyrolysis reactor is integrated in a circulating sand system composed of a riser (the sand and char, stream 3, is transported by air, # 4), a fluidized bed char dombustor (IU), the pyrolysis reactor (Rl), and a down-comer. In this concept, char is burned with air (# 10) to provide the heat required for the pyrolysis process. Oil is the only product (# 19), as non-condensable pyrolysis gases (# 22) are combusted and flared (SI). Oil yields up to 75%, based on the feedtsock as received, could be obtained (see Figure 4: the line only represents a trendline). In the early oil samples, water contents of 25 to 35 wt.% were produced, while currently the water content has a consistent value of 24 to 26 wt.% with an ash content below 0.1 wt.%. As of date, almost 20 tons of oil has been produced, while approx. 3 tons oil has been made to produce the CEB material.
50 400
500
600
Temperature ('C)
Figure 4 The oil yield as a function of the pyrolysis temperature (gas phase residence time < 1 s) for several biomass feedstock types. The line represents a trend. CEB PREPARATION FOR LARGER-SCALE COMBUSTION TRIALS The experimental facility to produce larger amounts of CEB is shown in Figure 5 . Batch-wise, approx. 1 ton of CEB can be produced in one week. The recipe to produce CEB is to mix Ca(OH)2 with water to produce so-called slaked lime at the optimised conditions referred to above. This slaked lime is pumped into a stirred stainless steel reactor, and bio-oil is added over a period of about 1.5 hours. The mixture was slowly heated to 7OoCover 30 minutes and kept at this temperature for a further 2 hours. The CEB to used for the large-scale testing was stored at ambient conditions. The mixture will be thoroughly mixed before spraying and co-combustion.
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Slaked Lime Water
Slaked Lime + Water
Bid1
Water
ca enrichedma1 storage
L
A-
Figure 5 The CEB production facility DEMONSTRATION OF CEB IN A SMALL EXPERIMENTAL SET-UP Due to the abrasive character of the CEB together with its high viscosity, spraying of CEB is possible only if a suitable atomiser is applied in combination with an appropriate CEB pump. Various atomisers and oil pumps were tested for this purpose, after a first selection on basis of the physical properties of the feed (viscosity, density, solids content, moisture, surface tension, etc.), and the required spray characteristics (droplet size and its size distribution, spray angle, spatial distribution, power consumption, etc.). The atomisation system finally selected was tested with respect to the droplet diameter. A simple technique was used, in which the atomiser was positioned approx. 25 cm above a liquid collection vessel. The produced droplets were collected in a specific type of oil, yielding a dispersion of CEB droplets as shown in Figure 6. By means of a microscope the droplet diameter could then be easily determined.
Figure 6 Droplets observed in the oil-CEB mixture after spraying. 1594
The droplet size distribution could be measured with high accuracy and reproducibility by collecting the droplets in a microscope glass in oil. A droplet sue distribution (Figure 6) could be easily constructed from the photographs.
0
25
(a&)
15
100
Figure 7 Droplet size distribution upon spraying CEB at a flow rate of 0.37 kg/hr, cEB'2 N m 3 h for three different spray cap sizes: small (-), medium (0) and large (A).
Results can be presented as the fraction of particles smaller than a certain value versus the measured particle diameter in a cumulative F-curve (Figure 7). From these figures, an average droplet size of approx. 22 pm can be calculated for all nozzles and at the conditions presented. Apparently, larger nozzles do not influence the droplet size distribution. For two CEB flow rates, @cEB,and two gas flow rates for CEB atomisation, @&,EB, similar plots are presented in Figure 8. Representative values for the (number) average droplet diameter for CEB range fiom 30 to 50 elm, the lower values being observed at the higher air and CEB flow rates. Contaminant removal with CEB
CEB will be used for coal and waste combustion, and not primarily for heavy fuel oil. However, the projected large-scale test 7 MWh facility can only use heavy fuel oil. Therefore, this oil was used to test it on a small scale. A small flame tunnel with separate air and liquid fuel injectors was constructed to demonstrate the potential of CEB to recover SOz.Figure 10 shows a photograph of the set-up, with the control panel at the right hand, and the flame tunnel in the middle. The three injection ports for CEB material are visible at the fiont side. Through the opening in the middle (the nozzle has been removed), the flame from diesel combustion is clearly visible. In all experiments, CEB is sprayed perpendicular to the fuel stream. A portable Mass Spectrometer has been purchased to analyse the exit gases from the flame tunnel experiments on SO,.
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100
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50
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150
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Figure 8 Droplet size distribution after spraying CEB for a) two gas flow rates (1 and 2 Nm3/hr)and a ~ ~ ~ 4kghr . 3 and 7 b) two CEB flow rates (0.18 and 0.37 kg/hr) and @air, CEB=2 Nm3/hr.
1596
VENT
FLAME TUNNEL
A
300
V
A'
300
Figure 9 A schematic presentation of the 20 kW, flame tunnel. Indicated at the right side of the flame tunnel are the CEB injection ports.
Figure 10 Photographs of the flame tunnel, with the control panel at the right hand side, and the flame tunnel in the middle. The burner is located at the top (not shown). At the front side, the three injection nozzles for CEB material are visible. Through the opening in the middle, the clear flame fiom diesel is visible (right). 1597
Results SO2removal
For the first preliminary experiments, a 'standard' fuel oil burner system was purchased. Due to several reasons, however, a new pump and atomisation system was required, viz. a dedicated fuel pump and a different spray nozzle. In this new set-up, primary air for a proper fuel atomisation and secondary air for the combustion could be controlled separately. Furthermore, pumping and combustion of heavy fuel oil (sulphur content of 3.35 wt.% and a viscosity of 180 cP) was possible only, when mixed with diesel in a ratio of 60:40 (diese1:heavy fuel oil). Besides, the flame tunnel had to be preheated to approx. 750°C using ordinary diesel or ethanol as a start-up fuel. When a steady state operation was observed (after approximately 10 minutes), the diesel flow to the flame tunnel could be replaced by a mixture of heavy fuel oil and diesel (40:60), and the emissions of SO2 are recorded (approx. 400 ppm). Then CEB is injected and subsequently the SO2 levels are recorded every 30 s. Usually, almost instantly a sharp decrease in the SO2 concentration was observed, due to the reaction of the Ca with SO2. The experiments were carried out at various fuel flow rates, ranging fiom 0.5 to 3 kg/hr, and at a CEB flow rate corresponding to Ca:S ratios fiom 0 to 3. 100
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..............................................................................
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Figure 11 The SO2-reduction versus the time while combusting the 60:40 diesellfuel oil with injection of CEB. Operating conditions: @,a=1.5 kg/hr, @&,*=2.5Nm3/hr, @,h,,=20 Nm35r, @,h,c~~=2.2N m 3 h and Ca:S=l. Su& content oil approx. 1.36 wt. Figure 11 shows a typical result for the SO2-reductionwhen CEB (corresponding with Ca:S = 1) is injected in the flame of the diesellfuel oil mixture. On the average, the sulphur reduction was 55 to 60%. This relatively low value is due to the high fuel and air flows (leading to a short residence time of approx. 1.35 s), and the low Ca:S ratio. In real practice, however, residence times of several seconds are more realistic. The fluctuation in the results shown is due to the inevitable pulsation of the CEB flow rate. Results for different Ca:S ratios have been plotted in Figure 12. CEB apparently gives a SO2-reduction of nearly 100% at Ca:S=2. Compared to the injection of Ca(OH)2,an increase in the SO2-reductionis obtained. In Figure 12 the theoretical line for complete SO2-removalis given while assuming that only the following reactions are taking place: -+ CaO + H 2 0 Ca(OH12 1598
+ CaS04 CaO + SO2+ !4 O2 One of the explanations for this difference between the sulfation results for CEB and Ca(OH)2 may be that the effective particle size of the active components for the CEB material is smaller than those for the Ca(OH)2 material. This effect is already demonstrated in Figure 8 by varying the air for atomisation of CEB. Apparently, and due to a finer atomisation of the CEB material, the effective particle size decreases, resulting in an improved SO2-reduction.
0
1
2
3
4
5
6
Ca:S ratio (moYmoI) Figure 12a SO2removal as a function of the ratio Ca:S for CEB (-) and Ca(OH)2(0). Operating conditions: Ohe1=0.75kg/hr hel; O&,p=2.75Nm3/hr,aak,,=14 Nm3/hr, and Oak,c,3=2 Nm3/hr. The lines presented are theoretical lines calculated for mass transfer controlled conditions for effective particle diameters of 2.5 and 3 pm respectively, whereas the dashed line represents the theoretical limit. 100
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Ca:S ratio (moVmoI) Figure 12b SO2removal versus Ca:S for Q ) a i r , ~ ~ ~N= m l ' h (-) and Q)air,CEB=2 Nm3/hr (0)Operating . conditions: Ohel=1.5kg/hr fuel; a&,p=2.5Nm3/hr, and @&,s=13 Nm3/hr.The lines presented are theoretical lines calculated for mass transfer controlled conditions for effective diameters of 2.5 and 5 pm respectively, whereas the dashed line represents the theoretical limit. 1599
Results NO, removal The NO, concentrations were measured using a combination of the MS analysis system and a conventional fluw gas analyser. With CEB NOx-reduction levels of 18 and 38 % could be realised at a Ca:S ratio of 1 and 2 respectively. In comparison, a reduction of only approx. 12% was realised when spraying water in the flame, at a similar water-tooil ratio comparable with the CEB experiments.
0
1
2 Ca:S ratio (moVmol)
3
4
Figure 13 NO, removal as a function of the ratio Ca:S for CEB (-) and water (dashed line). Operating conditions: (Ph,l=0.75 kg/hr fuel; @,,k,p=2.75Nm3/hr, @,k,,=14 Nm3/hr, and <Pai,CEB=2 Nm3/hr.The lines presented are trendlines.
Theoretical calculations Only a small amount of Ca-particles and a hort residence time are required for the complete removal of SO2. In combination with the high temperatures in the combustion chamber (> 8OO0C),it seems reasonable to state that the removal of the sulfur is a mass transfer controlled reaction. For simplicity, it is assumed that the reaction of CaO with SOz is first order in SO2,and zero-order in Ca and oxygen: CaO + SO2 + % O2+ CaS04 Assuming that the reaction rate, represented by k, (m3$(m3,s)) is limited by the external gas-to-solid interface mass transfer to spherical particles, kr=Qa=k;6/dP yielding: ac k E V, - kgaaE,Vr- -~ 6k,~,V, -=
4" 4" +"dP The unknown parameters are the characteristic particle diameter, dp, the gas-toparticle mass transfer coefficient, kg, and the solid hold-up E,. For the very small particles and the very dilute two-phase system considered, the mass transfer number can be estimated by Sh=kgd@=2 (Ram and Marshall, 1952). The diffusion coeflicient, D, can be estimated by Fuller et al. (9) The solids hold-up equals GJp,u, (lo), with G, as the solids flux in kg/(mZrs),p, the particle density with us approximating the gas velocity u,. Values for these parameters are listed in Table 3.
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Table 3 Values and relation used for the calculations Parameter D L
4 Gs Sh Ps
us
’
Difhion coefficient Length Flame Tunnel Effective particle diameter Solids flux Sherwood number Density solids Solids velocity
kg/m3
Value Ref. 9 2-10-6 1.35 1-10 OCEB/( 1/4*n*DtA2) 11 2 1000
m/S
U,
m2/s m p
kg/m2s
The conversion data suggest that the initial S02-absorption rate is controlled .by
mass transfer of SO2 from the gas phase to the active surface of the Ca-particles, but later by diffusion rates inside the particle. The effective diameter of the calcium particles range fiom 2 to 5 pm, depending on the spray conditions and on the Cacontaining material (CEB or Ca(OH)2). Most likely, for CEB the CaO-efficiency is 1. In comparison, using lime instead of CEB yields CaO efficiencies < 1. FUTURE WORK
Combustion test with SO2 removal are ongoing, together with tests to demonstrate the NO,-reduction potential. Future tests will be carried out to reveal if the Ca in the CEB is bonded to organic groups or present in the form of suspended calciurnhydroxide. Tests will also be carried out in a 7 MWB heay fuel oil combustion facility. CONCLUSIONS
Wood derived fast pyrolysis oil and slaked lime produce a Calcium-Enriched Bio-oil (CEB) capable to remove SOz while combusting sulphur-containing hels. The bio-oil required is produced in a 200 kg/hr fast pyrolysis plant, while CEB is made in batchwise operation by slowly mixing slaked lime with bio-oil at controlled conditions. The prepared CEB material has pores in the size order of a few hundreds pm,and this m a y result in the so-called ‘popcorn-effect’ yielding a larger effective surface area. As plugging of the pores is probably the main cause for the incomplete utilization of CaO-based materials, the hgh porosity of CEB is an indication of the potential of CEB as a desulfiuization sorbent, as has been shown in TGA experiments. Co-combustion experiments are carried out in a small 25 k W h flame tunnel, combusting a mixture of diesel and heavy fuel oil (60:40 resulting in 1.34 wt.% S). An essential parameter in SO2-removalis shown to be the atomisation of the CEB material. Higher gas and or CEB flow rates improved the atomisation characteristics, yielding high S02-removal> 60 % at Ca:S=l . Compared to Ca(OH)2, injection of CEB during simultaneous heavy fuel oil combustion gave higher SO2 removal efficiencies. Further experiments should reveal whether this is due to the higher reactivity of the CEB, or to better atomisation. The conversion data suggest that the SOz-absorption rates are controlled by mass transfer of SO2 fiom the gas phase to the active surface of the Ca-particles. The effective diameters are 2 to 5 pm., depending on the spray conditions and on the Ca-
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containing material (CEB or Ca(OH)2). Conversion rates are controlled by external mass transfer rate at the start of sulfidation, but then by diffusion inside the particle. ACKNOWLEDGMENT
The financial support of the EU under contract JOR3-CT97-0179 and of NOVEM under the EWAB program is gratefully acknowledged. REFERENCES
1. http://www.iea-coal.org.uk 2. Soud H.N., 1994, FGD installations on coal fned plants, IEA Coal Research, ISBN 92-9029-239-3 3. Soud H.N., 2000, Developments in FGD, IEA Coal Research, ISBN 92-9029-339X 4. Pisupati S.V., Clark D.A., Hill M.A., 1996, Evaluation of Simultaneous SO2 and NO, Reduction potential of BioLimeTM, Pennstate Energy and Fuels Research Center, 1996 5. Pisupati S.V., Clark D.A., Hill M.A., 1997, A study on the Efect of Atomization Characteristics of BioLimeTM Simultaneous SOz and NO, Reduction Pennstate Energy and Fuels Research Center 6. Oehr K., Zhou J., Simons G., Whjtowicz M., 1997, Simultaneous SOz and NO, control with BioLimeTM derived from biomass pyrolysis oil, in Developments in Thermochemical Biomass Conversion (ed. by A.V. Bridgwater and D.G.B. Boocock), 1477-1481 7. Zhou J., Oehr K., Simons G., Barras G., Put B., 1997, Simultaneous SO2 and NO, control using BioLimeTM, in Biomass Gasification and Pyrolysis: State of the Art and Future Prospects (ed. by M. Kaltschmitt and A.V. Bridgwater), 490-494 8. Wagenaar B.M., Venderbosch R.H., Carrasco J., Strenziok R., van der Aa B., 2000, Scaling-up of the Kotating Cone Technology for Biomass Fast Pyrolysis, in: 1st World Conference and exhibition on Biomass for Energy and Industry. 9. Fuller E.N., Schettler P.D., Giddings J.C., 1966, A new method for predicting of binar gas phase diffucion coefficients, Ind. Engng. Chem., 58, 18 10. Kunii D., Levenspiel O., 1990, Fluidization Engineering, J. Wiley & Sons, New York 11. Ranz W.E., Marshall Jr. W.R., 1952, Evaporation from drop: part 1 and part 11, Chem . Engngn. Prog., 48, p.141, p.173 12. Meier D., 1999, New Methods for Chemical and Physical Characterization and Round Robin Testing, in: Fast pyrolysis of Biomass: A Handbook (edited by A.V. Bridgwater et d.),CPL Press, Berkshire, UK, ISBN 1 872691 07 2, p. 91-101
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Improving Charcoal Kiln Performance Fundamental Studies Have a Role?
- Do
M.A. Connor Department of Chemical Engineering, University of Melbourne, Vic. 301 0, Australia
ABSTRACT: Though not the subject of much research at present, wood carbonisation or charcoal making is still one of the most widely used thermochemical biomass conversion processes. Existing charcoal kilns are comparatively inefficient and many researchers have sought to find ways of improving luln efficiency. Prospects for using fundamental research to enhance kiln efficiency and improve luln productivity and profitability are explored. Processes within kilns can be divided into microscale processes and macroscale processes. The former category encompasses processes that take place in and around individual wood pieces. A good deal is known about these processes but it is concluded that practical considerations limit the economic benefits that further research into these processes can bring. Much less is known about the latter category, which includes large-scale heat transport, mass transport and reaction processes within the wood stacked in the kiln.A better understanding of the hndamentals of these processes has a much better chance of contributing to improved luln designs and operating procedures.
INTRODUCTION Recent research into thermochemical biomass conversion processes seems to have been concerned mainly with areas such as gasification and rapid pyrolysis. As a result, research inputs into other areas such as carbonisation or charcoal production have diminished greatly. Yet charcoal making is still one of the most important of the thermochemical biomass conversion processes, affecting the lives and livellhoods of millions of people. The global significance of charcoal can be gauged from the large amount consumed worldwide - in 1996 this was estimated to be around 100 million tonnes annually [ 11. What is more, despite the widely held view that charcoal making is a declining industry, global demand for charcoal continues to grow [1][2]. The biggest single user of charcoal remains the Brazilian iron and steel industry. In 1990 thls industry was estimated to have an annual charcoal consumption rate of around 5 million toe (tonnes of oil equivalent) [3]. However, the amounts used can vary markedly from year to year. This is reflected in the figures for Brazilian industry as a
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whole, which reportedly used over 1 1 million tonnes of charcoal in 1991 but only 7.3 million tonnes in 1992 [l]. Although industry is a significant user of charcoal, the major users are people living in Third World towns and cities. For households and small-scale commercial enterprises, charcoal has distinct advantages over wood for both cooking and heating [4]. These advantages have made it a preferred fuel in many urban centres across the developing world, including the Middle East, Southeast Asia, East and West Africa, South and Central America and the West Indies [4]. Virtually all the charcoal made in these regions - other than that used for specialised industrial purposes - is made in lulns. A wide range of kiln types is in use. Some are quite primitive, like the simple earth and vegetation covered kilns used by itinerant charcoal burners. Others, such as the beehive kilns used by Brazil’s iron and steel industry, are comparatively sophisticated [ 5 ] [6]. Perhaps most widespread, and typical of the hlns used by smaller commercial enterprises in the developing world, are traditional earth or earthmound kilns like those to be found in Zambia [7] and Sudan
PI. The nature of the carbonisation process is such that the charcoal removed from a kiln weighs a lot less than the original wood charge. For kilns of the earthmound type, charcoal yields are around 23% to 27% (based on the original mass of oven-dry wood), or as low as 13% when based on the original air-dried wood mass [7][8]. These values are well below the theoretically attainable maximum of 44-55% [9]. The inefficiencies of kilns, and the implications these have as far as deforestation rates are concerned, have prompted many researchers to seek ways of increasing charcoal yields. One approach has been to use larger or more sophisticated kilns. The benefits of this approach are well illustrated in a study [8] comparing a traditional Sudanese earthmound kiln with a metal kiln of the type developed by the Tropical Products Institute [lo]. In this study, the metal kiln gave a 33% yield, as compared to 27% for the earthmound type. Further improvements can be achieved by going to large, top of the range kilns such as the beehive kiln or the Missouri kiln but reported yields rarely exceed 35% for charcoals with at least a 70% fvred carbon content. Yield values, such as those given above, are given much prominence in the scientific literature on charcoal. In fact, scientific papers often give the impression that charcoal yield is the key parameter of concern to the charcoal industry. In practice, this is not the case. Researchers may be preoccupied with charcoal yield but commercial charcoal producers are much more concerned about profitability. A good yield can certainly contribute to a high profitability but other factors are of equal or greater importance. Some of these factors may be largely unrelated to the processes that take place within the kiln, examples include the capital cost of the kiln and its ancillary equipment, and local labour costs. Such costs can be the determining factor when it comes to choosing what type of kiln to use. In a study carried out in Montserrat, the economics of four quite different small to medium sized kilns were analysed [l 11. By far the most profitable of these proved to be the comparatively primitive, local ‘coal pit’ type of kiln. The low capital and labour costs of this kiln, together with its ten year lifetime, made it far more economically attractive than more.short-lived, costly and sophisticated kilns offering only a marginal improvement in charcoal yield. However, other factors affecting profitability are much more closely linked to what happens inside the kiln. One of these is the marketability of the charcoal. If charcoal is to command a high price, its properties have to conform to users’ requirements. Iron and steel plants, for instance, may require a charcoal with a higher fixed carbon content
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and a higher strength than is needed by domestic consumers. However, these plants may also tolerate a fines content greater than that acceptable to domestic users. The latter generally prefer a lumpy, dense, slow-buming charcoal with a good resistance to breakage during transport and handling. Domestic users also dislike charcoals that emit a pungent smoke or give off showers of sparks while burning. Some of the above characteristics are determined in part by the tree species from whch the wood charge comes. Charcoal makers are well aware of this and are selective in their choice of woods. However, other charcoal characteristics are more dependent on what takes place inside the luln. For example, prolonging the carbonisation period is known to enhance the charcoal yield as well as reducing the incidence of craclung (and hence the percentage of fines) [ 121. It should be evident from the above discussion that maximising the profitability of a charcoal making enterprise is not a straightforward process. Charcoal quality has to be maintained, yet at the same time the carbonisation period must not be too long or productivity will suffer. Heating rates must be high but not high enough to give too many fines. Capital, operating and labour costs also have to be taken into account, as do disparities in selling price for charcoals of differing quality. In the case of large-scale plants using wood of consistent quality, experience alone may enable a good compromise between all these factors to be achieved. Such a situation exists in Brazil where the large producers source much of their wood from large Eucalyptus plantations. Wood harvested from these plantations arrives at the lulns as pieces of uniform length, reasonably similar in diameter, and with consistent physical and chemical characteristics. This uniformity makes it possib1e;over time, to refine luln designs and operating practices and come up with a set of procedures that gives a satisfactory product in an acceptable time and at an affordable cost. However, kiln operators elsewhere generally lack access to a uniform wood supply or else see advantage in being able to switch between a variety of different wood sources as prices and availability alter. Such a policy may enable savings in wood costs, but it also makes it harder to establish appropriate operating procedures. What is needed is some way of predicting how luln operating procedures should be modified to allow for changes in wood characteristics. To do h s requires a better understanding of the fundamentals of kiln processes than we have at present. Some of these fundamentals are fairly well understood but others are poorly known.It is the aim of this paper briefly to review these fundamentals and discuss the extent to whch, through better understanding, producers using established kiln types might be helped to achieve better yields and greater profitability when using a variety of wood types and sizes.
KILN PROCESSES
Charcoal kiln processes can be divided into two groups, referred to here as microscale processes and macroscale processes. The former group encompasses those processes occurring close to or withm individual wood particles. The latter group includes larger scale processes such as the convective transport of heat through the wood stacked in the kiln. When reviewing these processes it is important to remember that the production of charcoal in kilns is itself a batch process rather than a continuous one. For each batch there are four stages [4]:
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The loading stage, during which as much wood as possible is stacked in a predetermined pattern inside the kiln. The carbonisation stage, during which a portion of the charge is burnt; the heat released is used to drive off moisture and raise the remainder of the wood charge to a temperature high enough for charcoal to form. The cooling stage, during which the charcoal is allowed to cool to a temperature low enough to prevent it catching fire when exposed to the air. The unloading stage, when the charcoal product is removed from the kiln. If profitability is to be maximised, all four of these stages need to be properly managed. The first and fourth stages are purely mechanical and it might seem unnecessary to cover them in a review of process fundamentals. This is true for the fourth stage but not for the first; this is because the structure of the woodpile formed when loading the kiln has a strong influence on gas and heat transport during subsequent stages. The stage of greatest importance is obviously the second, since it is during this stage that all the processes accompanying charcoal formation occur. As its name implies, the third or cooling stage does involve some heat transfer but the processes occurring are much more straightforward than those in the second stage. MICROSCALE PROCESSES Heat transfer, mass transfer and reaction processes all play a role in the transformation of a piece of wood into charcoal. Past research has shown that the carbon laid down within a carbonising particle is a product of both primary and secondary decomposition reactions. The larger part of the carbon is formed during the initial (primary) pyrolytic decomposition reactions that occur when the wood is heated. During these reactions the organic components of the wood break down to yield a mixture of gases, condensable vapours, and a solid carbonaceous residue known as char. Significant further amounts of char can also be produced during secondary reactions. These are reactions undergone by volatile products of the primary reactions as they move from the interior of the carbonising particle to its surface [12]. Both the primary and the secondary decomposition reactions produce char and hence have the potential to affect charcoal yield and hence kiln profitability. Factors known to influence the nature, extent or rates of these reactions are reviewed below. Wood origin
It was mentioned earlier that charcoal makers are well aware that woods from different tree species yield charcoals of differing quality. These differences can be linked to differences in wood anatomy and composition. There are well known and wellcharacterised anatomical and chemical differences between softwoods and hardwoods, for example [13], and softwood charcoal is widely regarded as inferior to that from hardwoods. Significant differences can also exist between the charcoals and charcoal yields from superficially similar species of the same genus. This has been shown in the case of Eucalyptus species [14][15]. There can even be differences in charcoals made from trees of the same species when these come from regions with dissimilar climates or soil types. Informed and careful charcoal makers no doubt make some allowance for most of the above differences by segregating their wood stocks on the basis of species or
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source. A more difficult problem is how to allow for the differences that exist between woods from different parts of the same tree. The heartwood and sapwood of many tree species can be quite different in structure [15], as can the so-called earlywood and latewood layers. In addition, wood from areas subjected to non-uniformly distributed tensile or compressive stresses (as happens where branches intersect the main trunk) can differ significantly from wood that has developed in a more evenly stressed environment. Segregating wood pieces from the same tree would seem quite impracticable. The only option open to charcoal makers appears to be to choose a set of kiln operating conditions that gives an acceptable result for all the structural types present. Wood composition
Given the diversity of tree species from which the wood used in charcoal kilns could be drawn, chemical composition might be expected to be a factor of some importance in charcoal making. Heightening h s expectation is the fact that the major organic components of wood contribute unequally to char formation, with the bulk of the char coming from the lignin fraction [ 16][171. However, because the relative proportions of the major organic components vary little between species - see Table 1 below variations in these components appear of only minor consequence to charcoal makers. Table 1 Composition of softwoods and hardwoods [ 131.
Component Cellulose Hemicelluloses Lignin
Mass % in Softwoods
Mass % in Hardwoods
42 f 2 27 f 2 28 f 3
45 f 2 30f5 20 f 4
One wood component that could be expected to influence thermal decomposition patterns is its mineral matter content. Introduction of various inorganic compounds into wood is known to modify wood behaviour during heating. This change in behaviour has been exploited by makers of fire retardants. These chemicals, largely inorganic in nature, suppress the formation of volatiles in wood during heating and promote char formation. Although these introduced compounds have a well-characterised impact on wood decomposition, it has yet to be demonstrated that the mineral matter naturally present in wood has any effect on levels of char formed during pyrolysis. This apparent lack of q a c t is rather surprising since it has been shown that leaching of various salts from agricultural residues does lead to changes in decomposition behaviour [ 18][191. Nevertheless, on the basis of the evidence available, it would seem that ash content is another property that can be disregarded by charcoal makers - except, of course, as it affects the calorific value, ash levels and burning characteristics of the final charcoal product. The wood component likely to be of greatest interest to charcoal producers is water. Freshly cut wood can have a moisture content in excess of 50% and even air-dried ,wood has an equilibrium moisture content of 10 to 15%. In kilns, the heat required to drive off thls moisture comes from the burning of a fraction of the wood charge. It follows that the greater the wood moisture content, the more wood has to be burnt and the lower is the overall charcoal yield. The longer drylng time will also increase the
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length of the carbonisation cycle, which could be expected to M e r reduce hln productivity. Despite these apparent disadvantages, it is reported that a charcoal producer in the Marianas found it better to use green wood (which entailed firing its metal kiln for a longer time than normal) since this gave a better quality charcoal that was easier to market [20]. This practical observation ties in with theory, which suggests that using wood with a higher moisture content could lead to an increased charcoal yield (based on wood dry weight). This is because an extended drying period will prolong the time wood spends at lower temperatures during pyrolysis, and this can favour decomposition pathways promoting char formation [21]. However, whether or not an increased char yeld always eventuates is not clear. A fairly recent paper did review published work on the impact of moisture content on biomass behaviour during pyrolysis but the information available was found to be conflicting and contradictory [22]. Based on the evidence presently available, and despite the findings of the charcoal producer in the Marianas, the safest and most profitable approach would seem to be to do what most charcoal producers normally do and keep the wood moisture content as low as is practicable. Wood structure and physical properties As discussed earlier, of the char formed during thermal decomposition, a significant fraction can be laid down as a result of secondary reactions. These reactions occur as the volatile products of primary decomposition migrate out of the decomposing wood. Evidence is accumulating that suggests that the extent of secondary char formation depends on how long volatiles remain within the woodchar matrix before escaping into the gas phase. It follows that any property that affects migration rates through the woodchar matrix could influence the overall charcoal yield. Density is one such property. It might seem logical that when the same wood constituents are more closely packed, as must be the case in woods of higher densities, the rate at which volatiles can diffise to the wood surface should fall. Woods of a hlgher density might therefore be expected to give greater charcoal yields. Support for this hypothesis comes from a study showing that higher density woods from species such as oak and hickory gave higher yields of charcoal than lower density woods from species like cottonwood [23]. However, a study of ten Eucalyptus species in Brazil showed no correlation at all between charcoal yield and density [14]. Similar conclusions were reached in a study of various Australian and South-East Asian hardwoods [ 151. However, the latter study did show that higher charcoal yields were obtained from higher density sections of a single wood sample of reasonably uniform structure. This implies that density may well influence charcoal yield slightly but that its d u e n c e is often overshadowed by other factors. In practice, these findings are probably of little relevance to charcoal producers. Producers are well aware that their customers generally prefer a dense charcoal and that a dense wood is needed to make such a product [ 111. This means that denser woods are in any case being targeted by charcoal makers. In the above rather simplistic discussions on density, the arguments put forward implicitly assume that wood is a homogeneous, isotropic material. Of course, in reality wood has a non-uniform, anisotropic, vascular structure. A low permeability wood could be expected to delay the escape of volatile pyrolysis products for a longer time than more permeable woods, increasing the opportunities for carbon-depositing secondary reactions to occur. This suggests that permeability could well be a better
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predictor of charcoal yield than density - provided, of course, that the permeability of the wood is representative of that of the carbonising woodchar matrix through which the volatiles make their escape. Whether or not this is the case is not known at present and fbrther research into permeabilities could provide some useful ihsights. Permeabilities have been measured for a wide range of tree species and span six orders of m a p t u d e [24]. This suggests that if there does exist a correlation between permeability and charcoal yield, th~sshould be easy to c o n f i i In practice it is not that simple. Because of local structural differences, samples of wood from the same tree can show permeabilities that vary by an order of magnitude or more [25]. Moreover, permeabilities are markedly anisotropic; in a study of beechwood, for example, longitudinal, radial and tangential permeabilities were found to be in the ratio of 13000: 0.03: 0.02 [24]. From a practical viewpoint, the longitudinal permeability is likely to be the determining value. A recent study of this parameter for a number of Eucalyptus species showed that, for the group of species studied, longitudinal permeability was a much better predictor of charcoal yield than density [ 151. However, when the study was broadened to take in a wider range of tree species, a much poorer level of correlation between permeability and charcoal yield was found [ 151. The usefulness of the above findings to charcoal makers is uncertain. Few would have the equipment, the expertise or the time to make accurate measurements of permeability. The need to make such measurements could be partly overcome if a reference set of permeability (and density) values for commonly used woods were available. Even then, predicting charcoal yield from such values remains a chancy business. There is evidence that woods with a high density and a low permeability will generally give a high charcoal yield, and vice versa, but exceptions do occur. It is also probable that related woods with similar properties will give similar yields, but again, nothmg is certain. One set of properties over which charcoal makers do have some control is the sue and shape of the wood pieces loaded into the kiln. Many practical considerations affect the way wood pieces are cut and shaped before being put in the kiln. The first of these is the wood source. If waste wood from wood-processing industries is being used as a feedstock, then the size and shape of the wood pieces are largely predetermined. However, if the wood comes from plantations grown specifically to supply the lulns, there is greater scope for optimising size and shape characteristics. Desirable features include:
(1) A diameter that is not too large - a maximum of 200 mm has been suggested [lo]. If diameters are too large, the wood interior will not be fully carbonised - unless the heating period is prolonged (and overall kiln productivity decreased). (2) A length that is suited to the kiln design - 450 to 600 mm has been recommended for transportable metal kilns [ 101 but much greater lengths are usual in beehive kilns. Transport and handling considerations will also affect the choice of wood length. (3) Straightness and a lack of protrusions - these features make it easier to create a low voidage, closely packed wood stack and to maximise the mass of charcoal made in each batch. In practice, compromises have to be made. Where a wood is very hard, for example, it may be more cost effective not to split large wood pieces but to recycle them through the kiln until they are hlly carbonised. Separation of “brands” (wood
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pieces that are not fully carbonised) for reprocessing is in fact a frequent occurrence on kilns; these brands can result from poor distribution of heat within the kiln as well as from the use of oversized wood pieces. Charcoal producers undoubtedly learn a lot from experience about how to adapt kiln operating conditions to suit woods with different properties. Nevertheless, a better understanding of how size and shape properties affect wood decomposition processes should help them do this more efficiently. For instance, research has shown that the larger the wood particle, the higher is the charcoal yield, providing that the particle does not crack [12]. Cracking provides volatiles from the wood interior with a rapid escape route out of the wood. Ths is disadvantageous as it reduces the extent of secondary carbon-forming reactions. Cracking also decreases the average size of the charcoal product. The extent of cracking can be reduced by decreasing the rate of heating. However, if taken too far, the gains in the area of charcoal yield and quality will be offset by the fall in annual output. In the case where particles are non-isometric, the orientation of the longest wood dimension relative to the wood grain can also have a small effect on charcoal yield [12]. This finding would seem to be of most relevance to those producing charcoal in retorts, where the wood is often cut into uniformly sized pieces before being carbonised. If kiln operators tried to exploit this finding, more extensive preprocessing of the wood charge would be needed; given that the likely gains in yield are small, the cost effectiveness of doing this is questionable. It should be evident from the above discussion that, in theory at least, for any given wood charge there is a set of operating procedures that will maximise the benefits to the charcoal producer. Altering procedures to suit each wood batch may be possible on small-scale owner-operated kilns but it seems unlikely that larger kiln batteries could show the same flexibility. Here there is a need to coordinate the activities of employees responsible for loading, running and unloading a group of kilns. Making the best use of their time requires close adherence to a preset schedule, and scope for modifjmg operating procedures is limited. In such cases, it seems that theory has relevance only insofar as setting the original cycle times is concerned or in helping producers to know how best to blend wood stocks from different sources or to work out what extent of further preprocessing of wood stocks is worthwhile. MACROSCALE PROCESSES
During a kiln’s carbonisation stage, part of the wood charge is burnt, and the heat released provides the driving force for pyrolytic decomposition of the remaining wood. This involves a variety of heat transport, mass transport and reaction processes. These are dealt with individually below. Heat transport
A key tactic in maximising kiln profitability is to minimise the fraction of the wood charge that is burnt. This means makmg the best possible use of the heat generated during combustion of t h s (sacrificial) material. One obvious action is to limit heat losses to the surroundings. Various approaches would appear feasible, for example decreasing the thermal conductivity of the kiln walls or surrounding the luln with insulation. However, reducing heat losses from the kiln exterior will also lengthen the kiln cooling stage, with fewer charcoal batches being produced annually. This loss of
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productivity will soon offset any savings made by reducing heat losses during carbonisation. It also helps explain why measures to reduce external heat losses are missing from many luln designs. One way round this problem is to find another means of cooling the charcoal. Circulating cool gases through a kiln is feasible in theory, but use of air would cause the charcoal to catch fire, and inert gases are most unlikely to be available at kiln sites. Quenchmg with water has been used [l] but the thermal shock when the cold water strikes the charcoal surface can cause additional fracturing and an unwanted reduction in average particle size. A promising alternative, observed under test at a luln site in Brazil, is to introduce a fine spray of water into the top of the kiln. By the start of the cooling stage, the shrinkage that occurs during carbonisation has created a sizeable space in the top half of the kiln. The injected water droplets evaporate in this space before reaching the charcoal surface. This cools the gas and accelerates the cooling of the kiln contents without affecting charcoal quality. There is another approach that may not reduce the amount of wood burnt but that could still improve profitability. This is to stack the wood in such a way that those pieces least likely to yield good quality charcoal are the pieces most likely to be burnt. A logical extension of this approach is to design kilns so that the heat needed for carbonisation is obtained by burning wood unsuitable for charcoal production. This would be easiest to achieve if the wood combustion zone were separated from the carbonisation chamber. It was reported fairly recently that the Brazilian company, Acesita, had built a test kiln with such an external combustion chamber. This was done to see if better use could be made of forest residues unsuitable for conversion to charcoal [ 11. It was reportedly found that the cost of harvesting these residues had made the proposed scheme uneconomic and that the external chamber was now being used only as an ignition point and to control air flows into the main chamber [ 11. The most energy efficient way of providing the heat needs of a kiln would be to burn the vapours given off as the wood inside the kiln decomposes. If the hot gases formed in thls combustion step could be diverted back into the kiln, the fraction of the wood charge sacrificed to provide heat could be substantially reduced. (msis the approach adopted in continuously operated charcoal retorts, which lie outside the scope of this paper.) What makes this idea impractical for a single kiln is that much of the heat needed has to be available before evolution of vapours from the wood commences. However, it is tempting to ask whether better use couldn’t be made of this heat at kiln sites where a number of kilns operate simultaneously; if kilns were operated on a staggered schedule, heat recovered from the vapours of one kiln could be used in the next kiln, and so on. Even if it is impractical to direct the energy content of these vapours back into the kiln, capturing some of thls energy nevertheless remains an attractive idea for kiln owners. What has been done at some luln sites in Brazil is to pass kiln off-gases through condensers [l]. These serve mainly to alleviate air pollution problems but they also recover the heavier volatile components in the kiln off-gases. The recovered condensate has been used as a supplementary fuel in nearby blast furnaces. Whilst this may not change the energy efficiency of the charcoal production process itself, it does mean that more effective use is being made of the energy content of the original wood. Within the kiln, the most important heat transport problem is ensuring that the heat generated during partial combustion of the wood charge is distributed quickly to where it is needed. In my opinion, this is the aspect of kiln operation where fbndamentals can contribute most to efficient charcoal makmg.
161 1
In an ideal kiln,the hot gases would distribute themselves through the wood stack in such a way that all wood pieces would be exposed to the same pattern of surface temperature changes. In practice this is impossible to achieve. What normally happens is that some readily combustible waste or fuel is ignited at one or more points within the kiln. Enough air is admitted to sustain limited combustion at each of these points and heat is carried away into the rest of the woodpile by the expanding volume of hot combustion gases. As these gases pass over the surfaces of wood pieces surrounding the ignition point they lose heat progressively to these wood pieces, which in turn heat up and start to decompose. Initially, when the surrounding wood is cold, the gases lose their heat within a comparatively short distance of the ignition point. However, as the surrounding wood becomes hotter, heat is carried to wood pieces further from the point of ignition. In effect a heat front moves outwards through the woodpile from each of the ignition points. Behind this heat front follows a so-called carbonisation front; this defines the boundary of a zone within which the wood has been exposed, for a sufficient length of time, to temperatures high enough for charcoal to form. Within the kiln,carbonisation can only be said to be complete when the carbonisation fronts from the various ignition points have converged and extended throughout the woodpile. In effect, the time taken for the carbonisation front to reach the extremities of the woodpile defines the length of the carbonisation stage in the overall charcoal production cycle. Factors affecting the movement of carbonisation fronts should therefore be of great interest to charcoal producers, since any shortening of the carbonisation stage should improve kiln productivity. Some experimental information on carbonisation fronts is available. For example, Briane & Doat [6] have illustrated how the carbonisation front develops and moves over time in an earthmound kiln with its ignition point situated centrally and near the top of the woodpile inside. Also Shah et al. have monitored the temperature changes occurring during charcoal production in a 2 m3 pilot scale Magnien kiln fitted with six thermocouples [26].However, I am unaware of any comprehensive study of factors influencing the rate of movement of carbonisation fronts. Various strategies are available to producers wishing to minimise the length of the carbonisation stage. One approach is to segregate the wood on the basis of size so that the larger wood pieces are closer to the point of ignition and the smaller pieces further away. Such an arrangement, for a modified Casamanca kiln, is illustrated in Briane & Doat [ 6 ] . This approach takes account of the fact that the time taken for heat to penetrate to the centre of a log of wood decreases as the log’s diameter decreases. The time needed for the wood to carbonise fully decreases correspondingly. The delays before the carbonisation front reaches the kiln extremities are therefore less serious if the wood at these points is of a smaller size. Another strategy is to increase the number and spread of ignition points. With a properly designed arrangement of ignition points it should, in theory, be possible to reduce the time needed for the carbonisation front to envelop the entire woodpile. However, to do this most effectively calls for carell coordination of the rates of advance of the temperature fronts moving outwards from each ignition point. Coordination of these fronts on an existing kiln would necessitate measuring temperatures continuously at a large number of points through the kiln.Not only is this practically inconvenient for the charcoal producer but it is made very difficult by the marked shrinkage of the woodpile that accompanies carbonisation. It would also add significantly to operating costs.
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Using measuring instruments for control purposes may well be practically unrealistic during normal luln operations. However, specific research investigations using a range of instruments can be rewarding. This is well illustrated by an investigation carried out in South Africa on some Armco-Robson hlns [27]. These kilns were made from corrugated iron sheets bent and joined to form a roughly semi-circular tunnel 12 metres long, 2.5 metres high and 3.8 metres across the base. Each end of the tunnel was closed off by a masonry wall. The intention was to use these kilns to produce a consistent high quality charcoal product with a fixed carbon content in excess of 90%. The original luln design had a limited number of air inlet points distributed along the base of the kiln. With thls design, unacceptably large variations in charcoal quality occurred across the kiln. It was felt that this could be due to the air mlet points being too widely spaced. This was investigated using oxygen probes to map oxygen concentrations around each mlet (igmtion) point. In this way, the sphere of influence of each inlet point was delineated. From this mformation it was determined how many more air inlet points would be needed to obtain more uniform conditions in the kiln and hence a more consistent product [27]. It would be even more helphl to kiln operators and designers if, for a given set of conditions, the progress of temperature and carbonisation fronts through the woodpile in a luln could be predicted theoretically. To do this would require a comprehensive model that describes reaction, heat transfer and mass transfer processes in the kiln for wood pieces of a known size and moisture content stacked in a particular way. Development of such a model is probably still a long way off but some information is being gained on the fimdamental aspects of this challenging problem. One crucial factor in determining how fast heat moves outwards from the ignitiodcombustion points is the permeability of the woodpile. Th~shas been shown by work carried out on a 3 todday pilot-scale wood residue pyrolysis plant in Papua New Guinea [28]. Air was admitted to the base of this unit through a rotating ‘airgitator’. An experiment undertaken using a fine sawdustlshavings fraction of less than 4 mm in size showed that the heat produced in the partial combustion zone penetrated only a short distance upwards, creating a narrow intensely hot zone immediately above the combustion zone. Temperature gradients at the top edge of this intensely hot zone were extremely high, with wood particles only a short distance above this zone remaining at temperatures close to ambient. This pattern of behaviour can be explained in terms both of the resistance to flow posed by the packed bed of small particles together with the large surface to volume ratios of the fine particles. Together these created a situation where both the heat transfer coefficient and the heat transfer surface were large, and hence the combustion gases lost their heat content very rapidly. A markedly different pattern of behaviour was observed when a coarse fraction of wood shavings greater than 4 mm in size was used. In this case the hot combustion gases could pass so easily through the packed bed of particles that they were still quite hot when they reached the top of the bed of particles. Temperature gradients through the unit were gradual, and outside the combustion zone temperatures were low and well below those needed to initiate pyrolysis [28]. Not only was there no pyrolysis but some of the heat produced was escaping unused from the unit - a waste of wood resources. It is evident that what was needed in the pyrolysis unit was a pattern of behaviour somewhere in between the above. This was shown to be achievable when the fine and coarse fractions of the feed were combined, though maintenance of stable operating conditions was difficult [28].
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Given the comparatively high permeability of a woodpile composed of logs up to 200 mm in diameter, and containing very few small wood pieces, temperature profiles in the kiln might be expected to be gentle, as in the above-mentioned pyrolysis unit when using a feed of coarse shavings. Partial confirmation of this comes fiom a recent study of temperature profiles in a packed bed of comparatively large wood particles (80 mm long and 60 mm in diameter) heated by hot inert gases [29].Temperature gradients were gradual across the upper part of this bed but a lot steeper close to the gas inlet. The latter finding was attributed to the fact that the wood used was green and that a high rate of moisture removal was occurring in the inlet region. It was also noted that temperatures at different levels in the bed were not particularly uniform, suggesting (not unexpectedly) that some channelling of the gases was occurring. What is not clear from the above discussion is why the pyrolysis unit using a coarse feed of wood shavings was unable to initiate pyrolysis whereas in kilns, where the permeability of the wood particles is even greater, pyrolysis is readily achieved. This is attributed to the fact that the combustion zones in kilns are localised at a few inlet points whereas in the pyrolysis unit the combustion zone covered the unit's entire base. The effect of limiting the number of air inlet points is to increase both the volumetric heat release rate and the gas flow rate. This will increase both the temperature difference between the gas and the wood particles surrounding the inlet point and also the velocity and turbulence levels in the gas flows (and hence the heat transfer coefficient). So a greater rate of heat transfer and a correspondingly steeper temperature gradient can be expected in the region around each air inlet point. One implication of the above discussions is that for any given kiln there will exist an optimum number and arrangement of air inlet points. If there are too few points, it could be expected that the carbonisation fionts will take longer than is desirable to reach the kiln extremities. If there are too many points, it may be necessary to increase rates of air addition in order to achieve desired temperature levels, and this may bring about a reduction in charcoal yield.
Mass transport A number of the mass transpodmass flow problems of relevance to kiln operation have been alluded'to above. It is evident that modelling of kiln behaviour will require a better understanding than we have at present of the gas flow paths within wood stacks of various patterns. Also needed is a knowledge of how the channels through such wood stacks alter as wood pieces carbonise and shrink - in a non-uniform way - and the woodcharcoal pile partly collapses. A technique that could be helpful in characterising gas flows in such piles has been developed [30]. It involves measuring the air permeability of woodchips and appears readily extendable to heaps of charcoal and other woodcharcoal mixtures where the permeability is not too high. Another mass transport process about which more needs to be known is the migration of condensable vapours that occurs ahead of the carbonisation fiont. Both water and other volatile substances are known to vapourise as temperatures rise and to be carried by gas flows into cooler regions where they condense. As temperatures at their new position increase these compounds may again evaporate and move on to cooler regions. This cycle of vapourisation and condensation is important since, as is widely appreciated in briquette-making, it can enhance heat transfer rates quite significantly.
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Reaction processes
Many of the reaction processes of relevance in a charcoal kiln are microscale processes in the sense that they are associated with the decomposition processes in individual particles. Reactions not falling into hs category are largely those associated with the combustion of wood pieces close to the air inlet points. Given the global lmportance of wood combustion, it was initially assumed when preparing this paper that there is s not to be already sufficient information available on wood combustion reactions for h a limiting factor in developing a charcoal luln model. However, in a recent paper it was stated that ‘no generally applicable model exists to describe the thermal conversion of a packed bed of solid fie1 particles’ [3 11. This paper goes some way towards remedying this lack but only for a wood combustor operating with an adequate air supply. Further work is needed to determine how this model needs to be modified before it can be applied to the more oxygen-limited combustion environments in kilns.
CONCLUDING COMMENTS It is evident that the contribution fundamental research can make to improving charcoal luln efficiency is constrained by practical and economic considerations. Further research into the processes occurring within individual wood pieces could in theory help in the fine tuning of kiln operating procedures. However, especially on sites where a number of kilns are being operated simultaneously, kiln schedules are strongly influenced by the availability of manpower and this can severely limit operating flexibility. It is nonetheless likely that adding a good understanding of microscale processes to wisdom gained through practical experience could still be of help to producers when wood characteristics alter and operating procedures have to be modified. A better understanding of the hdamentals of macroscale processes seems to have far more potential to assist charcoal producers improve kiln designs and operating methods. However, there is much still to! be learnt about these processes and the development of a comprehensive and accurhte model of kiln behaviour appears a long way off. The benefits of developing such a model would nevertheless be considerable on the heat transport, mass transport and certainly sufficient to warrant fkther and reaction processes taking place in
ACKNOWLEDGEMENTS Ms Melanie Viljoen’s help in obtaining much of the information used in preparing this paper is much appreciated.
REFERENCES 1.
2.
Rosillo-Calle F., de Rezende M. A. A., Furtado P. & Hall D. 0. (1996) The charcoal dilemma. Intermehate Technology Publications, London. D’Apote S. L. (1998) IEA biomass energy analysis and projections. In: Biomass Energy.-.Data,Analysis and Trends, pp. 151-79. IENOECD, Paris.
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Ackerman F. & de Almeida P. E. F. (1990) The industrial fuelwood crisis in Minas Gerais. Energy Policy, 18 (7), 661-8. Foley G. (1986) Charcoal making in developing countries. Technical report No. 5 , Earthscan - International Institute for Environment and Development, London. Emrich W. (1985) Handbook of charcoal making. D. Reidel Publishing, Dordrecht. Briane D. & Doat J. (1985) Guide technique de la carbonisation. EDISUD, Aixen-Provence. Chidumayo E. N. (1991) Woody biomass structure and utilisation for charcoal production in a Zambian rniombo woodland. Bioresource Technology,37,43-52. Khristova P. & Khalifa A. W. (1993) Carbonization of some fast-growing species in Sudan. Applied Energy, 45,347-54. Mok W. S-L., Antal M. J., Szabo P., Varhegyi G. & Zelei B. (1992) Formation of charcoal from biomass in a sealed reactor. Ind. Eng. Chem. Res., 31, 1162-1166. Paddon A. R. & Harker A. P. (1980) Charcoal production using a transportable metal luln. Rural Technology Guide 12, Tropical Products Institute, London. Wartluft J. L. & White S. (1984) Cornparing simple charcoal production technologies for the Caribbean. Volunteers in Technical Assistance, Arlington, USA. Connor M. A. & Salazar C. M. (1988) Factors affecting the decomposition processes in wood particles during low temperature pyrolysis. In: Research in ThermochemicalBiomass Conversion (Ed. by A. V. Bridgwater & J. L. Kuester), pp. 164-78. Elsevier, London. Thomas R J (1977) Wood structure and cornposition. In: Wood Technology: Chemical Aspects (Ed. by I. S. Goldstein), pp. 1-23. American Chemical Society, Washington, DC. Brito J. 0. & Barrichelo L. E. G. (1977) Correlapjes entre caracteristicas fisicas e quimicas da madeira e a produgBo de carvPo vegetal: I. Densidade e teor de lignina da madeira de eucalypto. IPEF, Piracicaba, 14,9-20. Connor M. A., Viljoen, M. H. & Ilic, J. (1996) Relationships between wood density, wood permeability and charcoal yield. In: Developments in Thermochemical Biomass Conversion (Ed. by A. V. Bridgwater & D. G. B. Boocock), pp. 82-96. Blackie, London. Browne F. L. & Tang W. K.(1962) Thermogravimetric and differential thermal analysis of wood and of wood treated with inorganic salts during pyrolysis. Fire Research Abstracts and Review, 476-9 1. Hirata T., Kawamoto S. & Nishunoto T. (1991) Thermogravimetry of wood treated with water-insoluble retardants and a proposal for development of fueretardant wood materials. Fire and Materials, 15,27-36. Jenkins B. M., Bakker R. R., Baxter L. L., Gilmer J. H. & Wei J. B. (1997) Combustion characteristics of leached biomass. In: Developments in Thermochemical Biomass Conversion (Ed. by A. V. Bridgwater & D. G. B. Boocock), pp. 1316-30. Blackie, London. Connor M. A., Kisler J. P., Alesich N. I., Kane M. G., Watkins R. N. & Shallcross D. C. (1997) The pyrolytic decomposition and subsequent combustion of rice husks. In: Developments in Thermochemical Biomass Conversion (Ed. by A. V. Bridgwater & D. G. B. Boocock), pp. 67-81. Blackie, London. Paeniu B. (1988) Charcoal Kiln Project. An Innovative Rural Technology of the Energy Office of the Commonwealth of the Northern Marianas. Case Studies of
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Rural Development Experiences in the Pa@: Case Study No. 2, South Pacific Commission, Noumea, New Caledonia. Kilzer R. J. & Broido A. (1965) Speculations on the nature of ceilulose pyrolysis, Pyrodynamics, 2, 151-63. Antal M.J., Croiset E., Dai X., DeAlrneida C., Mok W. S-L., Norberg N., Richard J-R. & A1 Majthoub M. (1996) High yleld biomass charcoal. Energy and Fuels, 10 (3), 652-8. Cutter B. E. & McGinnes E. A. (1981) A note on density change patterns in charred wood. Wood and Fiber, 13 (l), 39-44. Smith D. N. & Lee E. (1958) The longitudinal permeability of some hardwoods and softwoods. Department of Scientific and Industrial Research, Forest products Research: Special report No. 13, H.M.S.O., London. Comstock G. L. (1965) L o n g i t u h l permeability of green Eastern Hemlock. Forest Products Journal, 15,441-9. Shah N., Girard P., Mezerette C. & Vergnet A. M. (1992) Wood-to-charcoal conversion in a partial-combustion luln: an experimental study to understand and upgrade the process. Fuel, 71,955-962. Johnstone D. A. & Gore W. T. (1985) The manufacture of industrial grade charcoal in Armco-Robson kilns. Paper presented at a Conference: Forest Products Research International - Achievements and the Future, Pretoria, 19pp. Connor M. A. (1983) Heat and mass transfer considerations in fuel production from wood wastes by pyrolysis. Regional Journal of Energy, Heat and Mass Transfer,5, 179-94. Aganda A. A., Murray P. W. & Kionga-Kamau S. (1997) Temperature profiles in a wood packed bed heated by hot inert gases. Trans IChemE, 75, Pt A, 677-84. Ernstson M-L. & Rasmuson A. (1 992) Field and laboratory measurements of the air permeability of chipped forest fuel materials. Fuel, 71,963-70. Bruch C., Peters B. & Nussbaumer T. (2000) A general model for the investigation of packed bed combustion with respect to wood. Paper presented at a Conference: Progress in Thermochemical Biomass Conversion, Tyrol, Austria, 1722 Sept., 13pp.
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Effect of Four Physical Characteristics of Wood on Mass and Energy Flows from Slow Pyrolysis in Retorts Y.Schenkel CRA,Agricultural Engineering Department Chausse'e de Namur, 146 - B 5030 Gembloux Belgium
ABSTRACT : The optimisation of charcoal production in a retort kiln calls for control of the carbonization mass and energy flows. These depend essentially on three types of factors : the physico-chemical characteristics of the raw material, the operational parameters and the reactor parameters. Carbonization experiments have been conducted to assess and to model the effect of four physical characteristics : moisture content (two levels : 0 %, 37 % dry basis), density (two levels : beech wood (650 - 740 kg/m' anhydrous basis) and poplar wood (398 - 426 kg/m' anhydrous basis)), ~ cm, length dimension and shape (two levels : cubes of 4 cm side and blocks of 4 x 4 I6 parallel to the fibres orientation). The carbonization final temperature was 500" C, the residence time at this temperature was I00 min. The heating rates were 2 and 20" C/min. These experiments show the considerable importance of moisture content regarding mass and energy flows ; the other characteristics are of secondary importance. The results lead to the development of a phenomenological modelling of mass and energy flows of the total volatile matter produced through wood carbonization in retort kilns, based on the logistic symmetric function. INTRODUCTION
The physical characteristics of wood can exert a considerable effect over the carbonization process and products. Water present in wood solids influences heat transfer : the evaporation of the water requires a significant part of the heat and limits the increase of temperature of the solid material. This results in a longer process of carbonization [ 1-31, lower mass and energy yields [2, 4, 81, and the quality of charcoal is modified [ I , 41. Heat transfer inside the wood so!id depend OR its thermal conductivity. This property is anisotropic and increases with the density, the moisture content and the temperature [ 10- 131. The thermal conductivity is higher in the longitudinal direction than in the radial and tangential directions : from 1.8 [ 121 to 2.5 [ 14, 151. The specific heat capacity of wood depends on its temperature and moisture content but not on its density or on the wood species [12, 16, 171. Several authors have proposed a formula to determine the specific heat capacity of wood [ 13, 16, 18, 191.
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The permeability of wood is a characteristic of the relative resistance it poses to internal material (water, steam, gases) transfers when heated. Permeability is also an anisotropic property. It is higher in the longitudinal direction (parallel to the wood fibres) in comparison with the transverse direction [20 - 221. Permeability varies considerably from species to species and increases with the temperature of the wood solid matrix [22 - 231. Because of the anisotropy of wood, the dimension and the shape of wood solids are determining factors of the process of thermal decomposition. Heat transfer is more rapid in the longitudinal direction, the limiting factors are related to the transversal direction. Consequently the thickness of the wood solid and the ratio transversal surface/tangential surface are critical [7, 24, 291. An increasing dimension of wood particles results in a longer carbonization time, an increased production of char and non condensable gases to the detriment of pyrolysis oil, although it is very difficult to predict the effect of the dimension and shape of wood particles on the composition and the quantity of each product of pyrolysis. The dimension determines also the type of thermal decomposition of the wood particle. If the particles are small (“thermally thin”), pyrolysis is controlled by the chemical kinetics ; if the particles are thick (“thermally thick”), pyrolysis is controlled by the heat transfers inside the wood solid. The limit in thickness is 5 mm [30], 0.2 mm [3 I ] or 1 mm [32]. But the dimension and the shape of wood particles also determine several characteristics of the bed of particles, such as the porosity, the bulk volume and the ratio between the surface of reaction and the volume of the particles. These parameters have an effect on the pyrolysis process, especially in fixed or moving beds. But many quantitative relations are still largely unknown. The density and the porosity of wood seem also to play a role in the pyrolysis process. But the effects observed are rather qualitative than quantitative. Wood with a high density and a low porosity tends to produce more charcoal [23, 29, 331. The density influences also the productivity of pyrolysis reactors through the bulk volume : the same volume of a reactor will produce more charcoal if a high density wood is pyrolysed compared to a light wood. The effect of the physical characteristics of wood on slow pyrolysis or carbonization has thus been quite extensively studied. However, these studies focus mainly on the qualitative effects of the physical parameters of the material. The optimization of the carbonization process as operated by charcoal producers around the world requires quantitative relations, particularly regarding mass and energy flows as determining operational parameters of a dynamic process such as carbonization. The carbonization process aims at the production of charcoal, preferably in large pieces (> 2 cm). Consequently, the carbonization process is mainly the thermal decomposition of thermally thick particles. A considerable amount of work has been done on the characterization and modelling of the pyrolysis of thermally thick particles of wood, [ 13, 17, 25, 28, 29, 32, 34 - 471. These experiments have been conducted on single particles of wood (cylinder, sphere, parallelepiped). However, few experiments have been carried out on a bed of thermally thick particles of wood [48-511, and we have not identified in the literature any work on the modelling of the pyrolysis of a bed of thermally thick particles. We have therefore conducted an experimental work aiming at : 1” Identifying the physical characteristics of wood that have a significant effect on the
dynamics of carbonization of beds of thermally thick particles.
1619
2" Building a model linking the mass and energy flows of this carbonization process to the physical characteristics of wood.
MATERIALS AND METHODS The carbonization tests were conducted in a 27 1 thermobalance kiln. A short description of the installation is given below. A complete description is presented in
~ 1 . (a) The reactor : cube-shaped, it has a capacity of 27 litres. The inside walls are made of refractory bricks. The feedstock is put in a metallic basket and the reactor is closed tightly with a cordon of clay which is replaced at each experiment. Sealed up, the reactor is totally airtight. (b) The heating control system : the electrical resistances are driven by a numeric regulator which sets the temperature profile and gradient as well as the final temperature of pyrolysis. The amplitude variation of the real temperature is more or less 20°C in comparison with the temperature settings. (c) The electronic balance : the reactor is set on an electronic balance to follow continuously the loss of weight (measurement precision : 0.5 YOat 10 kg). (d) The gas conditioning and analysis system : at the exhaust of the reactor, the pyrolysis gases are first condensed, filtered and continuously analysed (non condensable fraction). The liquid fraction (condensable gases) is collected and weighed. Thermocouples give the gases' temperature at different levels of the plant : in the reactor and along the conditioning line. The non condensable gases are analysed by means of a NDIR spectrometer (determination of carbon monoxide (CO), carbon dioxide (CO,) and methane (CH4) content), a thermal conductivity analyser (hydrogen (H,) content) and a content). magnetomechanics analyser (oxygen (02) The pyrolysis oil is collected at the bottom of the condensation columns, weighed to obtain the wet mass before determining its moisture content (Karl Fisher method). The main experimental plan is based on 4 factors and 6 fixed parameters :
4.
Species : Fagus sylvatica (beech), Populus x canadensis (poplar). Dimensionhhape : cubes 4 cm side, blocks 4 x 4 x 16 cm (length parallel to fibres). Moisture content : anhydrous (HO) and wet (moisture content ranging between 23 and 55 YOd.b., average 37 YOd.b. - H37). Heating rate : 2 and 20 "C/min.
1. 2. 3. 4. 5. 6.
Final temperature : 500 "C. Initial temperature : 20 "C (at time 0). Residence time at final temperature : 100 min. Mass of samples : 4.5 kg (anhydrous), bulk arrangement. Replications : 3. Cooling : natural in the kiln, at least 24 hours.
1.
2. 1
J.
The scheme is a full factorial plan. Other factors have been tested in orientation experimentation : cubes of 2 and 8 cm side, medium moisture content (1 1 and 21 YO
1620
d.b.) and wood blocks arrangements (peripheral and central piles, alternate rows). The two chosen wood species present very different physical characteristics : density, thermal conductivity, permeability. Table 1 gives the mean values of the major physical characteristics ; the detailed values are available in [52].
Table I Mean values of moisture content, density, thermal conductance, specific heat capacity, void volume, ratio surface of reaction/volume of the particles of the wood samples (Fagus sylvatica - beech, Populus x canadensis - poplar). Cubes 4 Moisture content (% d.b.) HO H37 Density (kg/m3 d.b.)
Beech Blocks 16 0.00 36.10
0.00 38.25
650 to 740
Thermal conductance (W/K at 293 K) HO 17.434 H3 7 24.299 Ifibres HO 7.748 H3 7 10.799
4.153 5.698 7.382 10.130
// fibres
Cubes 4 0.00 36.48
Poplar Blocks 16 0.00 38.88
398 to415
11.028 14.843 4.902 6.597
2.767 3.774 4.9 I9 6.709
Specific heat capacity (kJ/kg K at 293 K) HO I .235 2.038 H3 7 2.034 0.56 0.64 Void volume (YO) 0.56 0.64 I50 1 12.5 Surface/volume 140 112.5 The mass and energy flows of carbonization have been analysed through a phenomenological approach, as regularly developed in chemical engineering. To determine the parameters of the mass and energy flows, we have considered the production of non condensable gases and pyrolytic liquids as well as the volatilization of the solid, as a growth function : the masses and energy produced by the carbonization process, measured according to the time of carbonization, are cumulated. In order to analyse and compare the results obtained, a mathematical function has been searched and the symmetric logistic function identified. The equation of this function is : =M
,(I *-K-+I)
where y = mass of product (g). x = time of carbonization (min). M = maximal mass corresponding to a infinite time (8). a = time corresponding to 50 YOof the produced mass (abciss of the inflexion point (min). b = indicator of the reaction speed at the point of 50 % of produced mass. The symmetric logistic function is shown in figure 1 . For the solid volatilization function, a transformation of variables has been applied to convert the decreasing function (loss of weight) into a growth function (quantity of volatile matter produced). 1621
After adjustment, a statistical analysis of variance is applied to the three parameters of the symmetric logistic function, M, a, b.
M 2000
1000
0 0
100
a
200
300
Temps min-
Figure I The symmetric logistic function and its parameters, M, a, b. The energy flows of carbonization have been determined for the pyrolytic liquids and the non condensable gases. In the case of pyrolytic liquids, the energy content is determined in two steps. Firstly, the actual net calorific value of pyrolytic liquids is measured on a wet basis. Multiplying the net calorific value by the corresponding mass of pyrolytic liquids, we obtain an energy flow. In a second step, we must take into account the fact that, in an industrial process, the two fractions of gaseous effluents condensable and non condensable phases - are generally burned in the gaseous phase. This means that the heat required by the vaporisation of the two fractions of condensable gases (water and pyrolytic liquids) is not necesgary. Knowing the proportion of each phase for each sample of condensable gases, we have added to the actual net calorific value of the wet condensed matter, the product of the mass of each phase and its latent heat of vaporisation (water : 2500 kJ/kg, pyrolytic liquid : 450 kJkg). The calorific value of non condensable gases is calculated on the basis of their chemical composition in terms of CO, H2, CH4, C2H4. Then, the calculation is identical to the one used for the condensable fraction : multiplication of the net calorific value by the corresponding mass of non condensable gases and obtention of an energy flow per unit of time. A total energy flow is obtained by adding the energy flows of the condensable and non condensable fractions of the gaseous effluents of carbonization Like the mass flows, the energy flows have been transformed into a growth function, to which the symmetric logistic function is adjusted. The parameters of this function, M, a, b, are analysed by means of an analysis of variance. These parameters do not correspond to a production of mass but to a production of energy (kJ).
1622
RESULTS
In this paper, the results are analysed for the two total functions, the total volatile matter production and the total energy production. The results related to each fraction of the gaseous effluents, condensable and non condensable phases, are presented in detail in [52]. MASS FLOWS Moisture content A very highly significant effect of the moisture content of wood on the symmetric logistic function is observed as would be expected (table 2 and figure 2).
Table 2 Effect of the moisture content on the parameters of the symmetric logistic function adjusted to the total volatile matter mass flows.
M
a 101 159
3 200 5 162 97
HO H3 7 Standard error
b 8.13 27.76 1.21
1.4
total volatile ~~~~
matter (9)
5000 4500 4000 3500 3000
HO HR2/
r/
,2500 2000 1500 1000 500 ‘
0
100
200
~
...~
-H 0% HR 2”C/minp .....- .H 37% HR 2”C/rnin
300
400
500
Time of carbonization(min)
~
I --
Figure 2 Evolution of the total volatile matter mass flows according to the moisture content of beech wood samples (anhydrous - HO, wet - H37) (heating rate : 2”C/min).
As expected, the factor M (maximal mass of matter volatilized) is significantly influenced by the moisture content of wood. The difference between anhydrous and wet wood samples is naturally explained by the quantity of water which has to be evaporated, contributing to an increase of the value-of M for the symmetric logistic functions for wet wood samples (H37).
1623
The two other factors a and b are also influenced by the moisture content. The mass flows are delayed and slowed down by an increasing moisture content of the wood, as shown in figure 2. The effect of moisture content is so overwhelming that it could hide a possible effect of the other parameters. Consequently, a further analysis of the results has been conducted separately on the results for anhydrous and wet wood samples.
Anhydrous wood samples The mean value of the parameters of the symmetric logistic function are given in table 3. Table 3 Mean values and standard error (italic) of the parameters of the symmetric logistic function adjusted to the total volatile matter production from anhydrous wood samples (HO) carbonization.
Beech cubes 4 heating rate 2 "C/min 20 "C/min Beech blocks 16 heating rate 2" C/min 20 "C/min Poplar cubes 4 heating rate 2 "C/min 20 "C/inin Poplar blocks 16 heating rate 2"C/min 20 "C/min
M
a
b
3 171 (126) 3 170 (29)
157 ( I ) 47 (1)
9.3 (0.6) 7.1 (0.6)
3 267 (127) 3 255 (49)
154 (2) 48 (2)
7.7 (0.7)
3 188 (100) 3 190 (51)
149 (2) 46 (1)
9.1. (0.4) 6.2. (0.3)
3 175 (30) 3 186 (45)
156 ( I ) 55 (4)
9.4 (0.6) 10.2 (2.1)
6.1 (0.3)
The factor M (maximal mass of total volatile matter) is not significantly influenced by the shape, the species or the heating rate. However, the factors a and b are significantly influenced. The heating rate has of course a considerable effect on the factor a (time corresponding to 50 % of volatile matter production). The time is delayed by a factor of 3 for the heating rate 2 "Clmin, compared to the heating rate 20 "C/min. The factor b (indicator of the reaction speed at the point of 50 % production) is also influenced, but significantly less. The reaction is more rapid in the case of heating rate 20 "C/min (b = 7.4) than of a 2 " C h i n rate (b = 8.9). We observe a significant effect of the shape of the wood pieces on factor a. This time is slightly delayed for the wood blocks 16 (a = 103) in comparison with cubes 4 ( a = 100). In the case of factor b, it is the species that has a significant effect. The carbonization reaction is more rapid for beech (b = 7.5) then for poplar (b = 8.7). For both factors a and b, we observe a significant interaction between the species and the shape. Blocks 16 of poplar are significantly different from the other wood samples : factors a and b are higher. The result is that the carbonization process of blocks 16 of poplar is delayed and slowed down.
1624
Wet wood samples The mean values of the parameters of the symmetric logistic function are given in table 4. Table 4 Mean values and standard error (italic) of the parameters of the symmetric logistic function adjusted to the total volatile matter production from wet wood samples (H37).
Beech cubes 4 heating rate 2 "Cimin 20 "C/min Beech blocks 16 heating rate 2" C/min 20 Wmin Poplar cubes 4 heating rate 2 "C/min 20 "C/min Poplar blocks 16 heating rate 2"C/min 20 "C/min
M
a
b
5 636 (72) 5 121 (186)
236 (I) 75 (2)
37.6 (0.3) 16.3 (0.8)
4 834 (81) 4 876 (100)
228 (2) 89 (3)
27.6 (2.2) 22.7 (1.2)
5 035 (340) 5 071 (120)
224 (6) 78 (2)
29.0 (5. I) 19.7 (0.5)
5 969 (915) 4 749 (80)
253 (9) 86 (3)
44.4 (10.8) 24.8 (2.2)
As for anhydrous wood samples, the three factors - shape, species, heating rate - have no significant effect on the factor M. Also, the heating rate has of course a significant effect on the factors a and b : the production of volatile matter is delayed and slowed
down for the rate 2 W r n i n . The shape of wood pieces significantly influences also the factor a. This one is equal in average to 153 min for cubes 4 and to 164 rnin for blocks 16. A significant interaction between the shape and the species is also observed for the factor a. As for anhydrous wood samples, blocks 16 of poplar differ significantly from other wood samples by a delayed volatilization of the material. ENERGY FLOWS
Moisture content As for mass flows, the analysis of variance shows a very highly significant effect of the
moisture content of wood samples on energy flows (table 5). Table 5 Effect of the moisture content on the parameters of the symmetric logistic function adjusted to the total energy flows of carbonization. HO H37 Standard error
M 24 502 29 123 745
a 1 04
I75 1.8
1625
b 7.3 20.4 0.7
The maximal energy production is higher for wet wood samples than for anhydrous wood. The difference lies in the energy production by the non condensable fraction of gaseous effluents, while the energy production by the pyrolytic liquids is similar for anhydrous and wet wood samples [52]. Furthermore, energy flows are significantly delayed and slowed down by the water present in the wood. Again, as for mass flow, the effect of moisture content can hide other significant effects of the other factors of variation. Consequently, a further analysis of variance has been conducted separately on anhydrous and wet wood samples. However, this supplementary analysis of variance shows two common observations for anhydrous and wet wood samples. I " The factor M (maximal quantity of energy) is not significantly influenced by the shape, the species or the heating rate. Thus, only the moisture content has a significant effect on the total production of energy from the gaseous effluents. 2" Whatever the moisture content, the heating rate has a significant effect on the factor a (time corresponding to 50 % production) but not on the factor b (indicator of the reaction speed at the point of 50 % production). There is thus a delaying effect on the energy production when the heating rate is decreasing.
Anhydrous wood samples The mean values of the parameters of the symmetric logistic function are given in table 6. We have identified significant effects of the shape and the species on factor a and b of the total energy production, as well as a significant interaction between shape and species. Blocks 16 of poplar are characterized by a total energy flow delayed and slowed down compared to other wood samples.
Table 6 Mean values and standard error (italic) of the parameters of the symmetric logistic function adjusted to the total energy production from anhydrous wood samples (HO) carbonization.
Beech cubes 4 heating rate 2 "C/min 20 "C/min Beech blocks 16 heating rate 2" C/min 20 "C/min Poplar cubes 4 heating rate 2 "C/min 20 "C/min Poplar blocks 16 heating rate 2"C/min 20 "C/min
M
a
b
27 120 (2402) 21 488(1067)
159 (3) 47 (0)
8.4 (0.6) 6.6 (0.2)
24 147 (1012) 25 490(571)
155 (3) 49 (4
6.6 (0.4) 5.5 (0.1)
23 264(1429) 24 273(722)
153 (1) 48 (1)
6.8 (0.1) 6.3 (0.5)
24 382(1274) 25 491 (1881)
160 (2) 58 (3)
7.5 (1.1) 10.7 (1.2)
1626
Wet wood samples The mean values of the parameters of the symmetric logistic function are given in table 7.
Table 7 Mean values and standard error (italic) of the parameters of the symmetric logistic function adjusted to the total energy production from wet wood samples (H37) carbonization.
Beech cubes 4 heating rate 2 "C/min 20 "C/min Beech blocks 16 heating rate 2" C/min 20 "C/min Poplar cubes 4 heating rate 2 "C/min 20 "C/min Poplar blocks 16 heating rate 2"C/min 20 "C/min
M
a
b
34 991 (244) 30 855 (2089)
248 (I) 80 (1)
22.9 (0.9) 12.6 (1.2)
27 190 ( I 766) 30 361 (1352)
241 (6) 105 (5)
19.2 (2.8) 19.9 (1.8)
27 486 (2229) 27 945 (31 77)
239 (7) 100 (3)
19.5 (4.5) 19.8(1.7)
27 971 (4940) 29 033 (3115)
270 (7) 110 (7)
26.8 (3.2) 23.3 (3.5)
The factor a is significantly influenced by the species and the shape of wood samples. Total energy flows are delayed for poplar in comparison with beech and for blocks 16 in comparison with cubes 4. However, there is no significant effect on factor b nor significant interaction. DISCUSSION
The major fact resulting from the analysis of mass flow from wood carbonization is the general and considerable effect of moisture content on the process. The effect on the factor M of the symmetric logistic function (maximal mass of volatile matter produced at an infinite time) is expected : the water present in wood is evaporated and constitutes a part of the volatile matter. It is thus natural to observe a factor M higher for wet wood samples (H37) than for anhydrous samples (HO). The effect of moisture content on the two other factors of the symmetric logistic function (a - time corresponding to 50 % production, b - indicator of the speed of reaction) is also noteworthy. Confirming the results of orientation experiments (moisture content * dimension) [52] these results demonstrate the delaying and slowing down effects of the water presents in wood on the carbonization process. We expected also a natural effect of the heating rate on the factors a and b of the symmetric logistic function. This effect is observed : the production of volatile matter is delayed and slowed down for a heating rate of 2 "C/min in comparison with a 20 "C/min rate. Compared with the effect of the moisture content, the effect of the species and the shape of wood samples is considerably lower. In fact, the interaction shape*species shows that blocks 16 of poplar are characterized by total volatile mass flows delayed and slowed down whatever the moisture content of the wood. It is not an effect of the
1627
shape or of the species : cubes 4 of poplar and blocks 16 of beech have factors a and b of the symmetric logistic function almost equal to those of cubes 4 of beech. It is definitely the blocks 16 poplar which show different results. The cause of this difference could be a possible combined effect of the thermal conductance and the gas permeability. The thermal conductance is the lowest for blocks 16 poplar (see table 1) and the permeability to gases is the lowest for poplar. These two factors combined could contribute to delay and slow down the production of volatile matter for blocks 16 of poplar. However, we must stress the fact that the difference between these blocks 16 poplar and the other samples is not that big, although significant ; in the case of factor a, the difference is 5 to 6 min for anhydrous wood samples, 20 to 30 min for wet wood samples, when moisture content generates differences between anhydrous and wet samples of the order of 60 min. The adjustment of the symmetric function to the energy flows from wood carbonization is an original and dynamic.(and no more static) approach. The analysis of the symmetric logistic function demonstrates again the dramatic effect of water. As for mass flows, energy flows are delayed and slowed down for wet wood samples (H37). Water intervenes through the large quantities of heat it requires to be evaporated and eliminated from the solid matrix. Regarding the effect of the shape and the species, the analysis must be conducted separately for anhydrous (HO) and wet (H37) samples. For anhydrous wood pieces (HO), we observe the same effects as for mass flows ; a delay of the energy flows for blocks 16 of poplar. There is a direct link between energy and mass flows. However, for wet wood samples (H37), we observe a significant effect of the shape and of the species but without significant interaction between both factors. Energy flows are delayed for blocks 16 compared with cubes 4 and for poplar compared with beech. The effect of the lengthening of the wood blocks in the direction of the fibres as well as the effect of the species, act on the physical properties of wood samples in carbonization (decrease of the thermal conductance and of the permeability to gases). These effects could be enhanced by the moisture content of wood. Regarding the factor M of the symmetric logistic function (maximal quantity of energy produced), we observe a very high variability of the value taken by the factor M. This high variability of experimental results does not allow us to demonstrate a determining effect of one of the factors of variation. CONCLUSIONS
Moisture content of wood appears to be the most important physical parameter to take into account in wood carbonization. An increasing moisture content decreases the production of pyrolytic liquids and increases the production of non condensable gases by enhancing secondary reactions of pyrolysis inside the solid matrix. This effect is increased by the shape of wood samples ; wet blocks of 4 *4 * 16 cm produce less pyrolytic liquids and more non condensable gases than other wood samples. However; moisture content does not influence the chemical composition of carbonization products. This confirms the fact that water present in wood acts physically and not chemically in the carbonization of wood at low temperature and heating rate. Moisture content delays and slows down mass and energy flows of carbonization. But for energy flows, it is difficult to demonstrate a determining effect of water on the maximal quantity of energy produced ; the values we obtain are very variable and do not lead to a reliable quantitative relation between the moisture content and the factor M of the symmetric logistic function.
1628
Compared with the moisture content effect, the other parameters tested have a considerably lower effect on the carbonization process. If the heating rate also determines the mass and energy flows, as expected, it does not influence the carbonization yields nor the chemical composition of the products. This confirms the results of the literature : heating rates lower than 100 "Chin do not influence mass balances and products composition from the carbonization of wood. The shape and the species of wood samples (cubes of 4 cm side and blocks of 4 * 4 * 16 cm // fibres) do not exert a significant and general effect on carbonization cycles. We simply observe a different behaviour of blocks 16 of poplar that show delayed and slowed down mass and energy flows. The cause is perhaps a combined effect of a low thermal conductance (lengthening in the direction of fibres) and of a low permeability to gases (poplar species, lengthening in the direction of fibres). Besides, the shape and the species of wood samples slightly influence energy flows from carbonization. Energy production cycles from gaseous effluents of carbonization are delayed for blocks 16 in comparison with cubes 4, for poplar in comparison with beech. This new approach of mass and energy flows of carbonization based on the symmetric logistic function can be used to define a phenomenological modelling of wood carbonization.
REFERENCES Gore W.T. (1 982). Charcoal Production and Properties. CSIR Special Report Hour 262, Pretoria, South Africa, 1 17 p. Carre J, Htbert J., Lacrosse L. (1984). Analyse critique de la valorisation des matieres ligneuses par voie seche. Rapport CE, DC VIII, 1 13 p. Shah N., Girard P., Capart R. ( 1 989) - Carbonization of Wood : Product Analysis and Energy Assessment. Applied Energy, 34,223-241. Paddon A.R., Harker A.P. ( I 979). The Production of Charcoal in a Portable Metal Kiln. Tropical Products Institute - Report GI 19,28 p. Girard P. (1 989). Fiche technique de carbonisation. Document CTFT. Deglise X., Richard C., Rolin A,, Franqois H. (1980). Fast Pyrolysis/Gasification of Lignocellulosic Materials at Short Residence Time. In Proc. (( Energyfrom Biomass )>, I"' EC Conference, PALZ, CHARTIER and HALL ed., Applied Science Pub., 548-553. Beaumont O., Schwob Y. (1984). lnfluence of Physical and Chemical Parameters on Wood Pyrolysis. Ind. Eng. Chem. Process. Des. Dev., 23,637-641. Antal M.J., Croiset E., Dai X., De Almeida C., Mok W.S.L., Norberg N., Richard J.R., Majthoub M. (1996). High-Yield Biomass Charcoal. Energy Fuels, Vol. 10, 3,652-658. Humphreys F.R., lronside G.E. (1980). Charcoal from New South Wales Species of Timber. In : GORE (1982). [ 101 Stamm A.J. (1964). Wood and Cellulose Science. The Ronald Press Company, New York. USA, 549 p. [ I 1 J Graboski M. BAIN R. (198 1). Properties of Biomass Relevant to Gasification. In ReedTB, 1981,41-71. [12] Ragland K.W., Aerts D.J., Baker A.J. (1991). Properties of Wood for Combustion Analysis. Bioresource Technology,37 (1991), 161-168. [ 131 Gronli M. (1996). A Theoretical and Experimental Study of the Thermal Degradation of Biomass. Ph.D Thesis, Institutt of Termisk Energi og Vannkraji, Trondheim. Norway, 258 p + 2 annexes. 1629
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of Decomposition of Wood During Pyrolysis. PhD Thesis, University of Melbourne, Australia, 285 p. [ 1 81 Koch P. ( 1 969). Specific Heat of Ovendry Spruce Pine Wood and Bark. Wood Sci., 1,4,203-214. [ 191 Ten Wolde A., Mcnatt J.D., Krahn L. (1988). Thermal Properties of Wood and Wood Panel Products for Use in Buildings. DOEIUSDA - 21697, Oak Ridge National Laboratory, USA.In :Ragland et al., 1991. [20] Smith D.N., Lee E. (1958). Forest Products Res. Spec. Report 13. Her Majesty’s Stationery Office, London. [21] Comstock G.L. (1970). Directional Permeability of Softwoods. Wood and Fiber, 1,283-289. [22] Roberts A.F. (1970). A Review of Kinetics Data for the Pyrolysis of Wood and Related Substances. Combustion and Flame, 14,261-272. [23] Connor M.A. Viljoen M.H., Ilic J. (1997). Relationships between Wood Density, Wood Permeability and Charcoal Yield. In Proc. “Developments in Thermochemical Biomass Conversion”, BRIDGWATERA. V . ? BOOCOCK D.G.B. ed, Blackie Academic & Professional, Glasgow, UK, 82-96. [24] Connor M.A. (1983). Heat and Mass Transfer Considerations in Fuel Production from Wood Wastes by Pyrolysis. Reg. J. Energy Heat Mass Transjkr, 5 (3), 179194. [25] Martin G. ( 1 984). Pyrolyse gazeification du bois - Aspects physico-chimiques. ThBse de doctorat d’ingknieur, lnstitut National Polytechnique de Lorraine, France, 2 vol., 240 p. et 148 p. [26] Chrysostome G., Lemasle J.M. (1986) - Fluidised Bed Oxygen Gasification of Wood. In Proc. : (( Advanced Gasrfication )J, E.C. Series E, vol8, Energy from Biomass, A.A.C.M.Beenackers et W. Van Swaay ed, D. Reidel Publishing G, 2871. [27] Connor M.A., Salazar C.M. ( 1 988). Factors Influencing the Decomposition Processes in Wood Particles During Low Temperature Pyrolysis. In Proc. (( Research in Thermochemical Biomass Conversion JJ, BRIDGWATER A. V., KUESTERJ. L. ed., Elsevier Applied Science pub., Londres. UK, 164- 178. [28] Bilbao R., Millera A., Murillo M.B. (1994a). Heat Transfer and Weight Loss in the Thermal Decomposition of Large Wood Particles. / n Proc. (( Advahces in Thermochemical Biomass Conversion )J, BRIDGWATERA. ed, Blackie Academic (e Professional, Glasgow, UK, 833-858. [29] Zaror C.A. ( 1 982). Studies of the Pyrolysis of Wood at Low Temperatures. PhD Thesis, Imperial College, Londres, UK, 356 p. [30] Chan W.R., Kelbon M., Krieger-Brockett B. (1988). Single - Particle Biomass Pyrolysis : Correlation of Reaction Products with Process Conditions. Ind. Eng. Chem. Res., 27,226 1-2275. [3 11 Simmons G.M., Gentry M. (1986). Particle Size Limitations Due to Heat Transfer in Determining Pyrolysis Kinetics of Biomass. Journal OfAnaIytical and Applied Pyrolysis, 10, 1 17- 127. [32] Lucchesi A.. Maschio G.,Rizzo C., Stoppato G.( 1 988). A Pilot Plant for the Study of the Production of Hydrogen - Rich Syngas by Gasification of Biomass.
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In Proc. (( Research in ThermochemicalBiomass Conversion N, BRIDGWATER A. V., KUESTERJ.L. e d , Elsevier Applied Science pub., London, UK, 642-654. [33] Kumar M., Gupta R.C. (1997). Influence of Carbonization Conditions on the Pyrolytic Carbon Deposition in Acacia and Eucalyptus Wood Chars. Enera sources, 19,295-300. [34] Bamford C.H., Crank J., Malan D.H. (1946). The Combustion of Wood. Partl. Proc. of the Cambridge Philosophical Society. 42, I 16- 1 82. [35] Roberts A.F., Clough G. (1963). Thermal Decomposition of Wood in an Inert Atmosphere. 9IhIntl. Symposiumon Combustion. The CombustionInstitute, 158166. [36] Lee C.K., Chaiken R.F., Singer J.M.(1976). Charring Pyrolysis of Wood in Fires by Laser Simulation. I6lhIntern. Symp. on Combustion, 1459-1470. [37] Pyle D.L., Zaror C.A. (1984 b). Models for the Low Temperature Pyrolysis of Wood Particles. In (( ThermochemicalProcessing of Biomass j), BRlDGWATER A. V. ed., Butterworths R Co pub., Londres, UK, 20 1-2 16. [38] Chan W.C.R. (1983). Analysis of Chemical and Physical Process during the Pyrolysis of Large Biomass Pellets. PhD Thesis, Universityof Washington, 197 p [39] Desrosiers R.E., Lin R.J. ( I 984). A Moving Boundary Model of Biomass Pyrolysis. Solar Energy, 33,2, 1 87- 196. [40] Saastamoinen J., Aho M. ( 1 984). The Simultaneous Drying and Pyrolysis of Single Wood Particles and Pellets Made of Peat. Intl. Symposiumon Alternative Fuels and Hazardous Wastes, American Flame Research Committee, 1-29. [41] Capart R., Fagbemi L., Gelus M. (1985). Wood Pyrolysis : A Model Including Thermal Effect of the Reaction. In Proc. 3rdEC Conference on (( Energy.fiom Biomass )), Elsevier pub., Londres UK, 842-846. [42] Huff E.R. (1 985). Effects of Several Parameters on Burning Times of Wood Pieces. In Proc. (( Fundamentals of ThermochemicalBiomass Conversion Y, OVERENDR.P., MILNE TA., MUDGE L.K. e d , Elsevier Applied Sciencepub., Londres, UK, 76 1-775. [43] Simmons W.W., Ragland K.W. (1985). Single Particle Combustion Analysis o f Wood. In Proc. (( Fundamentals of ThermochemicalBiomass Conversion M, OVERENDR.P. MILNE TA., MUDGE L.K. ed., Elsevier AppliedScience pub., Londres, UK, 777-792. [44] Krieger-Brockett B., Glaister D.S. (1988). Wood Devolatization - Sensitivity to Feed Properties and Process Variables. In Proc. M Research in Thermochemical Biomass Conversion )), BRIDGWATERA. V., KUESTERJ. ed., Elsevier Applied Sciences pub., London, UK, 127- 142. [45] Bilbao R., Murillo M.B., Millera A., Mastral J.F. (1991). Thermal Decomposition of Lignocellulosic Materials : Comparison of the Results obtained in Different Experimental Systems. ThermochimicaActa, 190, 163- 173. [46] Koufopanos C.A., Papayannakos N., MASCHIOG G., Lucchesi A. (1991). Modelling of the Pyrolysis of Biomass Particles. Studies on Kinetics, Thermal and Heat Transfer Effects. Canadian Journal of Chemical Engineering, 69, 907915. [47] Connor M.A.. Daria V., Ward J. (1994). Changes in Wood Structure During The Course of Carbonization. In Proc. (( Advances in ThermochemicalBiomass Conversion )), BRIDGWATERA. V . ed, Blackie Academic h Professional, Glasgow, UK, 846-858. [48] Briane D., Haberman A. (1984). Essais comparatifs de six systemes de carbonisation artisanale. Association Bois de Feu. Paris, France, 188 p.
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[49] Shah N. ( 1 990). Carbonisation discontinue du bois en fours a combustion partielle. Thbe de doctorat, Universitk de Compisgne, France, 177 p. [50] Fossum M., Hustad J.E. (1 994). Biomass Gasification for Industrial Production of Tar and Charcoal. in Proc. (< Advances in Thermochemical Biomass Conversion )), BRIDGWATERA. V. ed., Blackie Academic & Professional, Glasgow, UK, 1242- 1256. [5 1J Buy Tien (1996) - Gasification of Wood : A Multi-Stage Approach. PhD Thesis, Asian Institute of Technology, Bangkok, Thailand, 1 13 p. [52] Y. Schenkel (1999). Modelization of mass and energy flows from wood carbonization in retort kilns (PhD thesis in French). Gembloux (Belgium), FacultC Universitaire des Sciences Agronomiques, 328 p.
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Influence of Temperature, Residence Time and Heating Rate on Pyrolytic Carbon Deposition in Beech Wood Chars Y. Schenkel CRA,Agricultural Engineering Department Chausse'e de Namur, 146 - B 5030 Gembloux Belgium
ABSTRACT : The concept of Pyrolytic Carbon Deposition is used to explain the difference measured between theoretical and observed mass yields of coal coke and wood char production. This concept has been applied on beech wood carbonization through a factorial experiment conducted on beech blocks of 2 cm side for 12 final temperatures (from 200 to 800" C), two residence times at final temperature (15 and 45 min) and three heating rates (2 - 5 - 10" C/min). Final temperature and residence time show significant effects. The Pyrolytic Carbon Deposit (PCD) increases with the final temperature of carbonization up to a peak corresponding to a temperature of 350" C, then decreases with increasing temperature. The PCD increase is related to an intensive volatilization phase of the wood material hence to a high level of recombination of the volatilized matter with the carbon structure (secondary reactions). The PCD decrease corresponds to a further thermal decomposition of the deposit itself. Residence time has also a significant effect but at low temperatures, below 400" C. A higher deposit is observed for the residence time at final temperature of 45 min, in accordance with a more developed volatilisation of the solid.
INTRODUCTION For carbonization, the maximum theoretical mass yield corresponds to the entire conversion of the carbon content of lignocellulosic materials into char. In other words, no single atom of carbon should be volatilized, the total amount of carbon atoms is converted into a structure of pure carbon. The maximum theoretical mass yield of charcoal would be expected to range between 45 and 50 YO,depending on the carbon content of the lignocellulosic material. This theoretical approach is unrealistic : large numbers of carbon atoms are volatilized during the carbonization process. Consequently, determining the maximum mass yield requires very accurate determination of the carbon content of both gaseous effluents and charcoal, according to the carbonization temperature. In order to determine this maximum mass yield of carbonization, we have conducted an. experiment on beech wood carbonization. The results of this experiment
1633
were then used in two different approaches to evaluate the maximum mass yield, the carbonization maximisation function and the pyrolytic carbon deposition. MATERIALS AND METHODS
The carbonization experiments were conducted on anhydrous beech cubes of 2 cm side in a laboratory electrical kiln. The wood blocks were placed in crucibles closed by a cap. Samples were introduced into the kiln at room temperature. The factors are : (1) The final temperature of carbonization : 200, 250, 275, 300, 325, 350, 375, 400, 450,500,600 and 800 "C. (2) The residence time at final temperature : 15 and 45 min. (3) The heating rate of the kiln : 2 - 5 - 10 "C/min.
Four replications were conducted for each set of experimental conditions. Several characteristics are determined on the charcoal : (1) Mass yield. (2) Proximate and ultimate composition. (3) Calorific value.
RESULTS
Table I summarises the average data observed for two major carbonization parameters, the mass yield and the fixed carbon content of charcoal. A highly significant effect of temperature was observed, as well as a significant effect of the residence time and a significant interaction between these two factors. The effect of heating rate was not significant. MAXIMISATION FUNCTION OF CARBONIZATION
The percentage of carbon remaining in the charcoal and the percentage of hydrogen and oxygen eliminated from the wood, were determined as a function of the final carbonization temperature. This is the actual objective of carbonization : to maximise the carbon content of the residual solid char while maximising the elimination of hydrogen and oxygen from this solid. The result is two inverse functions - captured carbon and volatilized hydrogen and oxygen - that are shown in figure 1. The product function of these two functions is also determined. This function (C * (H,O))can be defmed as the maximisation function of the carbonization process (see figure 1).
1634
Table 1 Mean values and standard error of the mass yield and fixed carbon content of charcoal, as a fbnction of the carbonization temperature and the residence time at this temperature.
Temperature ("C) 200 250 375 400 475 800
Mean 99.9 1 98.73 40.70 35.32 31.15 23.26
Temperature 200 250 375 400 47 5 800
0
Mean 15.17 15.51 56.03 65.48 74.43 93.98
200
Mass yield (% d.b.) Residence time (min) 15 Standard error Mean 0.01 99.59 0.23 94.37 1.17 36.76 0.38 34.0 1 0.31 29.98 0.16 23.08 Fixed carbon (% d.b.) Residence time (min) 15 Standard error Mean 0.23 15.81 0.18 17.04 2.00 64.61 0.76 70.28 0.59 77.19 0.22 95.25
600
400
45 Standard error 0.04 0.82 0.33 0.34 0.28 0.17
45 Standard error 0.13 0.30 0.43 0.75 0.25
0.24
800
Temperature of carbonization ("C)
Figure 1 Evolution of the carbon content of charcoal, of the hydrogen and oxygen content of the volatile matter, and of the product of these two functions (C * (H, 0)),as a function of the final carbonization temperature, for a residence time at final temperature of 45 min and a heating rate of 2 OC/min. 1635
Figure 1 shows the rapid increase of the maximisation function of carbonization (C * (H, 0))up to a temperature of 400 OC, followed by a slight hrther increase up to 800 "C. Above 400 "C, hydrogen and oxygen continue to be eliminated while proportionally less carbon is lost. However, the marginal gain in carbon is low and is obtained by means of a high energy consumption. Figures 2 and 3 show the maximisation function (C * (H, 0)),for the three rates and the two residence times at final temperature, 45 and 15 minutes.
%
45 50
F------"'---2"Clmin
- -cs - Heating
rate
5"Clrnin
0
200
400
600
,
800
Tern perature o f carbonization ( " C ) .-
~~~~
Figure 2 Maximisation function of carbonization as a function of final carbonization temperature, heating rate (2 - 5 - 10 "Clmin) and for a residence time at final temperature of 45 minutes. Again, a rapid increase of the maximisation h c t i o n of carbonization can be observed up to 400 "C, followed by a slight increase at higher temperatures. It is also remarkable to note the tendency (although not significant) observed for the heating rates : the higher the heating rate, the lower the maximisation function, especially at temperatures higher than 400 "C. Analysing the data, we notice that this difference is mainly due to the carbon content of charcoal. This carbon content is higher for the 2 "C/min heating rate than for the 5 "C/min rate, which itself is higher than the 10 "C/min rate, at all temperature levels. The hydrogen and oxygen content of the volatile matter at the same temperature levels does not differ from one heating rate to another. These results confirm some results of the literature [ 2 ] - [6] : a low heating rate tends to increase the production of char to the detriment of pyrolysis oil and gases.
1636
r
45
T-..~...~ .... .............................................................. "_
....
........ . ........,................
"
-
7-
I
400
200
600
Heating rate 2'Clmir
- 6 - Heating rate 5"Clmir
800
Temperature of carbonization ("C)
Figure 3 Maximisation function of carbonization according to final carbonization temperature, heating rate (2 - 5 - 10 Wmin) and for a residence time at final temperature of 15 min.
PYROLYTIC CARBON DEPOSITION Another approach has been developed in the field of mineral coal research in coke production. This approach is based on the volatile contents of the non-thermally treated material and of the char. As a result of the process of carbonization, the whole volatile matter of wood must be eliminated, except the residual volatile matter contained in the char. Consequently, the theoretical mass yield of carbonization is the ratio between the non volatile matter of the wood on one hand and the non volatile matter of the char on the other hand. The formula is [7] :
=
100 - V M , 100 - VM,
* 100
where M y , =, theoretical
mass yield of carbonisation, %.
VM, =volatile content of anhydrous wood, % db. VM,, = volatile content of anhydrous charcoal, % db. Classically the mass yield of charcoal is calculated as the ratio between the mass of charcoal produced and the mass of initial wood : My,, =-Mac
*loo
Maw
1637
where
Mac= mass yield of charcoal on anhydrous basis, %. M , = mass of anhydrous charcoal, kg.
M,
=mass of anhydrous wood, kg.
In practice, mass yields of carbonization (low heating rate) are always higher than the theoretical mass yields. The theory of pyrolytic carbon deposition (PCD) explains this phenomenon [6]-[12]. The level of PCD can be estimated by the means of a formula which calculates the PCD as the difference between the mass yield of carbonization ( m o c ) a n d the theoretical mass yield (Mc1,,). Table 2 summarises the value of PCD at 12 carbonization temperatures and two residence times. The volatile content of beech wood samples is 84.38 % d.b., determined on 9 samples. The evolution of the pyrolytic carbon deposit is shown in figure 4. Table 2 Determination of the pyrolytic carbon deposition as a function of the final carbonization temperature and of two residence times at final temperature (beech cubes 2 cm side - mean values for the three heating rates 2 5 - 10 OC/min).
-
Temperature ("C)
200 250 275. 300 325 350 375 400 450 500 600 800
VMac RT15 RT45 84.50 83.87 84.18 82.63 82.64 80.33 79.17 75.56 69.99 65.76 56.08 52.05 43.04 34.44 33.64 28.78 24.44 26.80 19.03 10.62 3.57 4.60 '
RT15 99.91 98.73 89.99 74.97 52.04 35.56 27.42 23.54 20.67 19.29 17.48 16.37
MY, PCD RT15 RT45 RT15 RT45 99.91 99.59 0.00 2.75 4.44 98.73 94.37 0.01 93.47 87.40 3.49 8.01 84.79 75.46 9.82 11.54 65.53 58.95 13.49 13.33 5 1.98 46.81 16.41 14.23 40.70 36.80 13.28 12.97 35.32 34.01 11.78 12.08 31.15 29.99 10.48 10.02 28.55 9.26 25.39 7.92 16.20 23.26 23.08 6.89 7.88
RT45 96.84 89.93 79.40 63.93 45.63 32.58 23.83 2 1.93 19.97
VM,, = volatile content of the charcoal, % d.b. My,,= theoretical mass yield of carbonization, %.
ma,= mass yield of carbonization, %. PCD = pyrolytic carbon deposition, %.
RT15 = residence time at final temperature of 15 min. RT45 = residence time at final temperature of 45 min.
1638
Pyrolytic Carbon &position (%)
14
12 10
8
l6 I
4
2
io
j m I
300
400
500
600
700
800
Temperature of carbonization ("c)
Figure 4 Evolution of the pyrolytic carbon deposition (PCD) as a function of final carbonization temperature and two residence times at final temperature (RT 15 and 45 min). The pyrolytic carbon deposition is low at low temperatures (200 and 250 "C), then increases to reach a maximum around 350 "C and decreases at higher temperatures. The low values noticed at low temperatures correspond to a phase of pyrolysis during which the volatilization of the solid matrix is just beginning. The recombination of atoms of carbon with the carbonized structure is thus strongly reduced. The maximum level of PCD at a temperature of 350 "C corresponds to an intensive phase of volatilization and hence to a high potential of recombination of the atoms of carbon with the solid matrix in pyrolysis. The decrease of the PCD with an increasing temperature above 350 "C is related to a thermal decomposition of the deposit, a secondary volatilization. The data we observe are in contradiction with the data published by [7]. They observe an increase of the pyrolytic carbon deposition up to 800 "C followed by a decrease up to 1000 "C.But the authors [7] publish result only for three temperatures (600, 800 and 1000 "C)and give only one value for each test without any information on the experimental plan (number of replications for example). We also observe a difference between PCD levels for the two residence time at final temperature. These differences appear at low temperatures, below 350 "C. The deposit is higher for the residence time 45 min, although the maximum peak is higher for the residence time 15 min. These differences are due to a more developed volatilization of the solid matrix for the residence time 45 min [I]. The results observed for the residence time 15 min show also a higher variability.
1639
CONCLUSION
Two approaches have been developed to assess the theoretical or maximum mass yield of carbonization. The maximisation function of carbonization underscores the critical phase of low temperature (below 400 "C) in the carbonization process. This function also confirms the importance of a low heating rate to maximise the production of charcoal. The evolution of the pyrolytic carbon deposition is mainly influenced by the temperature of carbonization. This deposition reaches a maximum peak at 350 "C, around 14 to 16 YO.Then, the pyrolytic carbon deposition decreases regularly down to 7 "C at 800°C. The residence time of the solid matter at final temperature is also a determining factor. A high residence time (45 min) leads to a higher pyrolytic carbon deposition, although the peak at 350 "C is high for the short residence time (15 min). The effect of the residence time appears mainly at low temperatures, below 450 "C.
REFERENCES 1. Schenkel Y, (1 999) Modelization of mass and energyflowsfrom wood
carbonization in retort kilns. (PhD Thesis in French). Gembloux, Belgium, Faculte Universitaire des Sciences agronomiques, 328 p. 2. Doat J., Deglise X. (1982). Gaztification par pyrolyse Cclair de quelques bois tropicaux. Bois et For& des Tropiques, 198, 59-74. 3. Mackay D.M., Roberts P.V. (1982). The Influence of Pyrolysis Conditions on Yield and Microporosity of Lignocellulosic Chars. Carbon, 20,95- 104
4. Beaumont O., Schwob Y. (1984). Influence of Physical and Chemical Parameters on Wood Pyrolysis. Ind. Eng. Chem. Process. Des. Dev., 23,637-64 1.
5. Kumar M., Gupta R.C. (1993). Influence of Carbonization Conditions on Physical Properties of Acacia and Eucalyptus Wood Chars. Transactions of the Indian Institute of Metals, 46 (6), 345-352. 6. Kumar M., Gupta R.C. (1997). Influence of Carbonization Conditions on the Pyrolytic Carbon Deposition in Acacia and Eucalyptus Wood Chars. Energy sources, 19,295-300. 7. Chiu Y.F., Hong M.T. (1983) Influence of Volatile Matter and Deposited Carbon on Coke Yield from Coals. Fuel, 62, 1150-1 152. 8. Anthony D.B., Howard J.B., Hottel H.C., Meissner H.P. (1 976). Rapid Devolatilisation and Hydrogasification of Bituminous Coal. Fuel, 55, 12 1-128. 9. Koufopanos C.A., Papayannakos N., Maschiog G., Lucchesi A. (1991). Modelling of the Pyrolysis of Biomass Particles. Studies on Kinetics, Thermal and Heat Transfer Effects. Canadian Journal of Chemical Engineering, 69,907-915.
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10. Martin G. (1984). Pyrolyse - gazeification du bois - Aspects physico-chimiques. Th&e de doctorat d 'inge'nieur,Institut National Polytechnique de Lorraine, France, 2 vol., 240 p. et 148 p. 11. Di Blasi C., Russo G. (1994). Modeling of Transport Phenomena and Kinetics of Biomass Pyrolysis. In Advances in Thermochernical Biomass Conversion (Ed. BRIDGWATER A.V), Blackie Academic and Professional., Glasgow, UK, 906921.
12. Antal M.J., Croiset E., Dai X., De Almeida C., Mok W.S.L., Norberg N., Richard J.R., Majthoub M. (1996). High-Yield Biomass Charcoal. Energy Fuels, Vol. 10, 3 , 652-658.
1641
Catalysed carbonisation of fine woodworking industry residues JZandersons, A. Zhurinsh, A. Tardenaka, B. Spince Latvian State Institute of Wbod Chemistry, 27 Dzerbenes str., LV 1006, Riga, Latvia
ABSTRACT: Sawdust, shavings,veneer shorts and other fine residues of the woodworking industry are promising raw materials for production of charcoal. Studies performed at the Latvian State Institute of Wood Chemistry have demonstrated that engme oil refinery acid tar (EORA tar) is a good catalyst of wood carbonisation. Charcoal yield can be increased by 20 to 54 % in comparison with that in the case of non-catalysed pyrolysis charcoal. The destruction and dehydration of hemicelluloses and cellulose starts already at 120 to 130°C. At 200°C. 70 % of acids and more than 80% of furfural are evolved. The formation of wood tar is decreased by an acidic catalyst promoting the process of cellulose dehydration.The dehydration of cellulose in the temperame range up to 280°C causes the formation of precursors of charcoal stmdures and provides a high charcoal yield. The maximum of tar evaporationis observed at higher temperatures than usually. In the temperature range 350 to 450°C the orgamc matter of the EORA tar or the decomposition products of the sulphonic acids are distilled off.The lower heating value of this mixture of tars is 1.4 times higher than that of an ordinary wood tar. Therefore, the combined heating value of volatile products is maintained at a high level and will ensure the sufficiency of heat energy for the carbonisation process even if a damp wood should be pyrolysed. The process utilises harmful indmtrial residues to produce a charcoal, which can be processed into activated carbon, carbonaceous construction materials and fuel briquettes. INTRODUCTION
Fine residues of the woodworkingindustry such as sawdust, shavings and veneer shorts are in low demand and are used as a fuel at best. The amount of these residues is quite impressive. For example, veneer shorts comprise 15 to 20% of the amount of veneer blocks while the amount of sawdust is up to 12% of the volume of sawlogs. Nowadays the carbonisation of these wood residues is an industrial practice. However, the charcoal yield is only 15 to 20% on the o.d wood basis if modem industrial carbonisation units such as a Herreshoff furnace (roaster) is used (2). One of the main
1642
reasons for the low charcoal yield of fine wood particles in industrial apparatus is the fast pyrolysis nature of the process, namely a high rate of heating at particle temperatures rangmg from 200 to 280°C. If the heating rate exceeds 20-25°C min-',the depolymerisation reactions of cellulose prevail and a lot of soluble tar is formed whereas the yield of charcoal is low (1). The solution to this problem is in the promotion of the dehydration reactions of cellulose using catalysts during the preexothermic period of the heated wood particles. Catalysts of wood pyrolysis are used to increase the yield of charcoal and to obtain a charcoal with modified properties or to produce definite valuable volatile products. The substances promoting condensation reactions are employed if high yields of charcoal are desired (4, 5, 7). Some catalysts increase the apparent density of charcoal, the pore dimensions and their adsorbency nature. Of all silvichemicals charcoal is currently the most in demand Commercial interest is focused on deciduous wood charcoal or charcoal briquettes. The Lewis and Brmsted acids lower the starting temperature of wood decomposition and charcoal formation and along with an increased char formation, diminish the yield of some volatile products of pyrolysis. The latter fact is an undesirable phenomenon if an energetic self-sufficiency in the carbonisationof wood is striven for (10). Therefore, the objective of the present investigation was to elucidate not only the effect of different available catalysts upon the yield of charcoal, but also that on the yield of volatile products. EXPERIMENTAL SECTION
The grey alder (Alnus incuna) wood was chosen as a test species in the present study for laboratory experiments, because it was also intended to elucidate the effect of catalysts on the mechanical properties of charcoal. Grey alder wood specimens were prepared in the shape of blocks measuring 2 x 2 ~ 2cm. The chosen amount of catalyst solution was soaked into wood and the specimens dried at ambient temperature to the moisture content 7 to 8%. An electrically heated thermoreactor was used. The heating rate was 3 to 4"C/min, the maximum temperature 500 to 520°C. To pyrolyse the sawdust and chips under isothermal conditions at the thermoreactor wall temperature 550"C, a pilot scale thermoreactor equipped with a two paddle rotating stirrer was used (11). Analyses of the volatile condensable wood thermodestruction products were done by using spectrophotometric methods developed at our laboratory (8). The settled tar was extracted with chloroform from condensate @yroligneous acid). An aliquot amount of extract was dried and the content of settled tar was determined gravimetrically. The dissolved tar was determined gravimetrically after settled tar extraction, by drylng an aliquot amount of water phase solution. For analysis of acids, esters, alcohols, ketones, aldehydes and furfural the test solution was prepared by dilution of a weighed sample of clear fraction of the condensate 10 g to 100 ml with distillate water. The content of acids and esters was determined by potentiometric titration, using potassium hydroxide and hydrochloric acid standard volumetric solutions, c(K0I-I) = 0.1 moV1; c(HC1) = 0.1 mom. An automatic potentiometric titration assembly RTS - 822 "Radiometer" (Denmark)was used. Alcohols were determined using ammonium cerium (IV) nitrate reagent, which reacts with alcohols to form a stable colour complex. Colour intensity was
-
1643
measured spectrophotometricallyat 490 nm. Content of alcohols was calculated against the standard solutions of methanol. Ketones were determined spectrophotometrically by reaction with salycilaldehyde in basic medium. Colour intensity of the formed complex was measured at 490 nm and content of ketones calculated against the standard solutions of acetone. Aldehydes forms yellow-orange complex with antron in 75 % H2SO4, but furfud - blue one. Intensity of colours was measured spectrophotometricallyat two wawelenghts - 490 nm (I,- of aldehydes complex) and 590 nm (I- of furfud complex). Content of aldehydes was calculated against the formaldehyde standard solutions, but amount of furfural - against furfud standard solutions. Sulphuric acid, diammonium phosphate, zinc chloride and EORA tar from the Inchukalns deposit site were tested as carbonisationcatalysts. The highly harmful waste acid tar is characterised by a good solubility in water: therefore, in an open air depository pond, it is found as a 50 % solution.inwater, containing 4.3 % sulphuric acid and 70 % sulphonic acids (R-SQH) both on the dry matter basis with an average sulphonic acid molecular mass of 610 (3,6). RESULTS AND DISCUSSION
Under the effect of carbonisation catalysts, the thermal decomposition of wood components sets in at a temperature approximately by 100°C lower than in the case of non-catalysed carbonisation of wood. Table 1 lists the yield of the wood carbonisation solid residue recalculated on fixed carbon to exclude the effect of the fluctuations of volatiles and the ash content upon the charcoal yield. The fixed carbon content of charcoal in the experiments represented in Table 1 was 78% to 81%. The increased ash content of charcoal was caused by the catalysts, ( W ) 2 m 4 and ZnC12. The yield and overall heating value of the volatile products are given in Table 1. This shows that the charcoal yield under the influence of the above-mentioned catalysts increases from 5 to 39%, but the decrease of tar yield is approximately proportional to the increase of the charcoal yield. The most effective catalysts (H2S04 and ~ ) 2 H p o 4cause ) a dramatic decrease in the dissolved tar yield (a decrease of 72.6 and 66.1%, respectively, on the basis of non-catalysed pyrolysis yield). The decrease of the settled tar yield is also sigmficant., but less remarkable in the case of H2S04(37.8 and 71.1% respectively). The decrease of the tar yield was the principal reason for the serious diminishing of the heating value of volatile products if sulphuric acid or (NH&HP04 were used as catalysts. The same interrelationship can be observed between the decrease of the gas yield and its heat energy. It has been demonstrated that engine oil refinery acid tar is an excellent catalyst of wood carbonisation (9). The increase in charcoal yields ranges from 20 to 54% on the basis of non-catalysed pyrolysis charcoal yield. The structure of grey alder charcoal shown on the electron scanning microscope images (see Fig. 1) differs a lot from that of the charcoal obtained using the EORA tar as catalyst (see Fig. 2a). Fig. 2a shows that some vessels are filled with a glassy material, possibly a coke from the EORA tar. Fig. 2b demonstrates that the glassy material fills the vessel. However such coke inclusions are comparatively few, and we suppose that these by no means are the main cause of the remarkable increase in the charcoal yield
1644
Ul
2
c.
7.19 100.0
24.6 6.2 4.5 7.2 1.9 2.3 0.9 1.8 1.4 17.1 5.04 0.02 1.04
Without catalyst
*concentration of catalysts- wt % on the 0.d. wood basis
YO
MJkg
Combined heating value;
CHS
H2
Aldehydes Ketones Furfural Noncondensable gases Including combustible: CO
Esters
Charcoal (as fixed carbon) Dissolved tar Settled tar Acids Alcohols
Products of wood pyrolysis
4.1
4.91 68.0
0.02 0.67
3.80 52.8
0.8 0.4 1.3 2.4 11.9 2.9 0.02 0.60
1.5
34.1 2.1 1.3 6 .O
5.0%
3.0%
30.0 1,7 2.8 7.6 1.5 0.7 1.O 1.o 3.8 13.7
(m4)2m4
H2so4
6.06 84.3
25.8 5.1 2.4 7.3 3.3 0.9 1.0 1.1 3.7 17.8 4.8 0.01 0.90
ZnClz 1.4%
7.16 99.6
7.00 97.4
873 6.5 3.0 1.2 1.4 0.8 2.9 13.0 3.65 0.03 0.73
8,2 6.3 3.2 1.3 1.6 1.4 3.3 13.5 3.39 0.01 0.43
31.3
7.22 100.4
14.1 4.33 0.06 1.23
4.0
7.6 1.2 1.1 0.6 0.4
8,2
36.3
Engine oil refinery acid tar 9% 17%
29.5
5%
Table I . Yield (wt% on the 0.d. wood basis) of grey alder (Alnus incana) wood catalytic pyrolysis products and the heating value of their volatile components.
Fig. 1. Cross-section (a) and longitudinal section (b) of grey alder charcoal It has been earlier demonstrated that sulphonic acids increase the charcoal yield. The amount of sulphuric acid in wood soaked with the EORA tar catalyst is no more than 0.7to 0.8% on the o.d wood basis. However the combined action of these cellulose dehydration promoters causes an increase in the charcoal yield
Fig. 2. Cross-section (a) and longitudinal section (b) of grey alder charcoal prepared using the EORA tar catalyst.
1646
In this connection. the information about the formation of the main volatile condensable products during the pyrolysis process is of interest. Tlus mfornlation is presented in Figures 3 , 4 and 5 .
0 350 Temperature, ' C
200
280
450
Fig.3. Formation of acids during catalytic* pyrolysis * in Fig.3,4 and 5 the mginr oil refmery acidtar catalyst cmvxntration m wood is
expressed in wt 00' on
the 0. d. woad basis.
It has been shown by our experiments, that the destruction of hemicellulose and cellulose starts at approximately 120 to 130°C. and up to 200°C the overwhelming amount of acids and fi,ufural evolved (if the EORA tar concentration in wood is 17% on the 0.d. wood basis at the temperature 200°C up to 70% of acids and 82.5% of furfural are evolved. see Fig. 3,4.5).
200
350
180 T Q m p e r l t U r e , OC
Fig. 4. Formation of krfural during catalytic pyrolysis
1647
450
9
8
m
7
'
6
1,
d
d g 4
g
3
e
2 1 0
ZOO
280
350
450
500
Temperature. 'C
Fig. 5. Formation of tar during catalytic pyrolysis
At the same time at the temperature 200"C, the formation and evaporation of tar is observed. Although the amount of tar is only 0.8% on the 0.d wood basis (see Fig. 5). it is 5 times as high as that without the use of the catalyst but approximately the same as with sulphuric acid. The maximum of tar formation is at higher teniperatures than if non-catalysed pyrolysis is employed. It means that at temperatures 350°C to 450°C the organic matter soaked into wood as the EORA tar catalyst or its decompositionproducts is evaporated. The lower heat value of this tar is 1.4 times hgher than that of the orchary wood pyrolysis tar (41.5 M J k g and 29.6 MJikg, respectively). Owing to the organic matter introduced into wood with the sulphonic acids, the combined heating value of the volatile products is kept at a non-catalysed pyrolysis level (see Table 1). The charcoal obtained has apparent density as well as longitudinal crushing strength and modulus of elasticity to be 244 g/dm3, 153.9 kg/cm2and 21.1; 277 g/dm3, 158.6 kg/cm' and 23.0; 326g/dm3. 182.7 kg/cm' and 23.2; 341 g/dm3, 201 g/cm' and 31.7 for non-treated alder wood charcoal and that soaked with 5%. 9% and 17% EORA tar specimens. respectively. These properties of charcoal are important not only for technical-grade lump charcoal, but also for production of hgh-quality charcoal briquettes and carbon materials from fine charcoal particles. Sawdust and veneer shorts are especially suitable raw materials for catalytic pyrolysis. In the present study the sawdust was treated with an appropriate amount of an EORA tar aqueous solution in a paddle mixer. the mixture was kept in a closed vessel for 48 hours, and the catalyst-soaked wet sawdust (the moisture content 45 to 50% on the wet basis) was pyrolysed in a pilot-scale thennoreactor equiped with a two paddle rotating stirer at a constant reactor wall temperature of 550 to 600°C. The charcoal yield calculated as fixed carbon was increased from 18 to 32.5% when the EORA tar concentration in the catalysed sawdust was 17% on the o.d wood basis. The bulk density of the charcoal was also increased from 188 to 220 g/l. The duration of the
1648
process was decreased moderately, since the most time-consuming operations were drying of wood and heating of charcoal to achieve tlie desired fixed carbon content. EORA tar is a promising catalyst to prepare charcoal for activated carbon production. The charcoal obtained in runs using the EORA tar atalyst was activated in half the time necessaq to activate the charcoal prepared without the use of the catalyst. The adsorption efficiency of activated carbon was 1.3 to 1.6 times higher if the charcoal was produced in a catalysed process, and the pore size of the adsorbent was also uniform. The charcoal prepared in an EORA tar catalysed process contained 0.3 to 0.5% of sulphur. The activated carbon d d not contain sulphur. It stands to reason that the implenientation of this catalyst in practice requires solution of the sulphur &oxide emission problem. The high yield of good-quality charcoal from dspersed wood residues and tlie utilisation of harmfkl wastes are arguments for the use of t h s technology. The level of sulphur dioxide emissions from a unit for catalytic pyrolysis of sawdust is the same as for a powerhouse using coal with low sulphur content. These emission levels should be tolerated to clean up the dangerously contaminated sites causing groundwater pollution.
C0N CLU SI 0N S Sawdust and other fine residues of the woodworlung industry are in low demand and thus create dsposal problems. To orgaruze an economically effective charcoal production tlieir yield should be improved. It has been disclosed that engme oil refinen acid tar can be used as a wood carbonisation catalyst. The charcoal yield of fine dspersed wood is increased up to 1.7 times. making the production practicable. In addtion to that, the process self-sufiency in a heat carrier is ensured by the combined combustion of pyroligneous vapour and the EORA tar carbonisation volatiles. The superfluous residue of woodworlung industry and the environmentally dangerous pollutant are used to ad\antage. The quality of charcoal is improved because the apparent density. crushing strength and modulus of elasticity are increased. Though the charcoal contains 0.3 to 0.5 % sulphur. it is eliminated during the activation process. The activated carbon obtained demonstrates a uniform pore size and its activation time is less by half in coinparison with non-catalJzed charcoal activation durability. The insignificant level of sulphur dioxide emission from a unit for catalytic pyrolysis should be tolerated to clean up the dangerously contaminated sites causing groundwater pollution and to manufacture a value-added product from a low in demand wood waste. REFERENCES
Antal M.J.. Jr. & Warhegyi G. (1995) Cellulose pyrolysis kinetics: The current state of knowledge. bid. Eng. Chem. Rex. 34. 703-717. FA0 Forestry Paper 63. (1985) Industrial Charcoal malung. FA0 UNO. Rome. 134 P. Frolov A.F.. Titova T.S.. Karpova I.V.. Denisova T.L. (1985) On a composition of the engine oil refinery acid tar. Kl?ini.va I technolopa topli\w I tiinsel K h e m i s t n ' and Technology of Fuel and Lubricants) N o 6. 37-38. (ulRussian) Kislitsyn A.N. (1990) Wood Pyrolysis. Chemistry. Kinetics. Products. Novel Processes. Publ. House "Lesnaja proniyshletuiost". Moscow. 3 12 p. (in Russian)
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Kislitsyn AN., Rodionova Z.M., Lebedeva Z.I., Kalugin E.N. (1984) Study of structure formation and modification of charcoal characteristics in the presence of chemical reagents. Khimiya drevesinv (Wood Chemistry), No 1, p. 83-89. (in Russian) 6. Kukurs 0.& Valdmanis J. (1999) I n W n s acid tar repository. I Stratification and physical characteristics of repository. Lahian Chemical Journal, 4, 100-3. (in Latvian) 7. Sek~guchiY. & Shafizadeh F. (1984) The effect of inorganic additives on the formation, composition, and combustion of cellulosic char. J. Appl. Polym. Ski. 29. 5.
1267-1286.
Spince B., Zhurinsh A & Zandersons J. (1998) Chemical analysis of wood pyrolysis liquid products. Latvian Chemical Journal, 3,92-5. (in Latvian) 9. Zandersons J., Zhwinsh A, Spince B. & Tardenaka A. (1996) Method for praducing charcoal using a carbonisation catalyst. Latvian pat. No. 11995. 10. Zandersons J., Zhurinsh-A & Gravitis J. (1999) Carboni&tion of wood residues by utilizing volatile thermolysis products combustion heat. Cell. Chem. Techn., 33. 8.
157-60. 11. Zandersons J.. Gravitis J., Kokorevics A., Zhurinsh A., Tardenaka A. & Spince B. (1999) Studies of the Brazilian sugarcane bagasse carbonisation process and products properties. Biomass and Bzoenergy, 17,209-19.
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Thermal Desorption Technology: Low Temperature Carbonisation of the Biomass for Manufacturing of Activated Carbon G. E. Someus Thermal Desorption Technology Group L.L.C. of North America European Branch: H-I222 Budapest, Szechenyi59., Hungary
ABSTRACT: The Thermal Desorption Technology Recycle-Reduce-Reuse TDT-3RTM apparatus is an indirectly fired, horizontally arranged, patented (US 5,707,592), continuously operating rotary luh design, whereas <600 "Celsius - material core carbonisation temperature, vacuum (0-50 Pascal) in true reductive environment and separate process of steam activation is used. The efficient heat transfer is managed through the mantel. No internal moving parts are used for material throughput. Alternative continental basic materials: sawdust (hard and soft wood), straw, refuse grain and hardcoal. Pre-treatment of the basic material: grinding homogenisation to <5 mm, drying to <18 % moisture content. Main process: phase I. low temperature carbonisation, phase 11. steam activation. Post treatment: screening and packaging. Gasvapour phase stemming from the main process is separately treated, condensed for offsite utilisation of liquid and non-condensable gas is burned out, or all gas-vapours are directly burned off for onsite energy utilisation. Non-destructive Air Pollution Control Device "APCD" multi venturi off-gas scrubber is used. Process water is cleansed and recirculated. The TDT-3RTM rotary reactor and the comprehensive design does not containing any exotic technical constructions andor exotic materials.
INTRODUCTION The method of the low temperature carbonisation of the carboniferous materials is known since long time, whereas different apparatus t e c h c a l solutions have been used. Many of these traditional carbonisation apparatus producing variating quality end products with too high burn off and does not meet the new environmental standards. For the innovative TDT-3RTMapparatus design the following main technical and engineering considerations have been taken for value added conversation of the biomass: (1) As biomass is efficient heat insulator, the efficient indirect heat transfer to the material - through to core - is difficult and takes time.
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(2) Jhgh initial moisture content have significant negative influence on the
carbonisationprocess.
(3) Upon condensation of the gas-vapours for utilisation as bio-oil, the initial moisture content will be enriched in the liquids, which complicates its utilisation, especially over 25 % of moisture content. (4) Flexible sealing between rotary kiln moving and stationary parts requires special solutions. ( 5 ) The thermal engineering of the mechanical construction design requires special solutions. (6) Pre and post treatment of material require comprehensive solutions. (7) Biomass provides lower yield of activated carbon production versus hard coal, therefore special end product applications need to be developed in order to compensate specific production cost efficiency loss at similar investment. Table 1. Fixed carbon content in weight % of raw materials (dry substance) used for the production of activated carbon
Material Soft wood .Hard wood Coconut shell Grain and a g o products Hard coal Anthracite
AmroXimate carbon content (%) 35 40
35 40
60-75 90
Large corporations have advanced industrialised activated carbon manufacturing technologies, however, these large units using mostly hard coal to maintain continuously supplied high yield production. Hard coal - as natural substance may vary in chemical composition, and often including natural add-on inorganics and heavy metals, which might be disturbing elements at applications. Biomass - as natural substance - may also vary some in composition, but heavy metal contamination is less of care, only if man made contamination is the case. The efficient carbonisation of the biomass requires high efficient heat transfer equipment from the heat source to the material, which is one the main key elements of the TDT-3RTM technology equipment design.
-
THE KEY ELEMENTS OF THE TDT3RTM TECHNOLOGY The TDT-3RTMtechnology is an high efficient indirect heat up and indirect cool down process and apparatus of the carboniferous basic material to achieve value added refined end products.
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The strategic components of the TDT-3RTM technology: Indirectly heated, horizontally arranged, rotary kiln for Thermal Desorption and low temperature carbonisation of the 0 - 5 mm input basic material, containing the following technical characteristics: True reductive environment. Indirect heat transfer through the mantel to the ground material with extended surface area resulting efficient heat transfer to the feedstock, often with bad thermal conduction character. Advanced technical solution for the rotary kiln & between the moving and stationary parts for closed-input and output system. Advanced technical solution for the continuous throughput of the feedstock where no moving parts are used. Permanent operation and material throughput. Process under vacuum. Automatic process control of the operation. S b l e construction, containing no exotic technical solutions and materials, are avadable for continuous feed stock, 0.8 Two main full scale reactor m3/h (6,000 m3/year) and 2.5 m3k (18,000 m3/year) with multi reactor installation opportunity. Multi venturi off-gas scrubber for high efficient cleansing of contaminated flue gases stemming from thermal conversion of rehse and waste materials, containing the following advanced and innovative technical solutions: Multi venturi technical solution. Combined off-eas cleansing effect, high efficient counter stream water spray washing, drop out and removal effect. Combined water recirculation system. For variating flue gas volumes the system is externally infinitely variable. External cooling opportunity to avoid boiling up. Modular tower construction. Solid material cooler: Indirect cool down of carbon products from 600 "Cto 40 "C. Avoids oxidation - thus ash creation - during cooling process.
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TDT-3RTM GENERAL CONCEPT DESCRIPTION PRE-TREATMENT Grinding and homogenisation to less than 5 mm of basic material. The ideal moisture content of the basic material > 5 % would be, however in practice 8-18 % is used. If necessary drying of the basic material is executed to decrease moisture content from >35 % to <10 %. Pre screening of the basic material is used, as the sizing of the input material will be reflected into the sizing of the end products. The closed input to the TDT-3R rotary kiln's internal area is managed by a mechanical device, which compress the feedstock to remove air from the inter space between particles and forcing material into the thermal processing area.
THE MAIN PROCESS The main component of the TDT-3RTMis a specially designed, indirectly fired, rotary reactor in which the material in a reductive environment is carbonised - partially vaporised andor gas-out - in low vacuum (0-50 Pascal) between the temperature ranges of 550°C - 600°C. The reactor rotates around its symmetric axle, is horizontally arranged and cylindrically formed with no refractory line installed inside. The high quality steel mantel is heated from outside through the lined combustion chamber. h i d e the reactor body, blades promote the transportation of the material. The reactor is a permanently working vessel; the basic material enters in at the input-end while the carbonised material and the pyrolysis gas-vapour phase are discharged at the output-end, separately from each other. Generally, carbon char and inorganics will remain in the solid phase while volatile organic compounds and water will be concentrated into the pyrolysis gasvapour phase. PYROLYSIS GAS-VAPOUR PHASE from the reductive decomposing process is either condensed for achievement of bio-oil and non-condensable gases for energy utilisation or all directly combusted on site at a minimum 850 OC 2 sec. residence time. Alternative post combustor has been developed for operating temperature of 1,250 OC 2 sec. residence time as well, with fast cooling and heat recovery opportunity. The remaining gas is cleansed in a high capture efficient, indirectly cooled, wet gas multi venturi scrubber prior to discharge. Scrubber process water is cleansed prior to discharge and the neutralised precipitate with from the water treatment is utilised.
CARBON PHASE is discharged for medium temperature steam activation. The end product is indirectly cooled, dried, and selected by our adjustable air selector for final sizing and utilisation.
TDT-3RTM PROCESS REACTOR AND IT'S OPERATION The required amount of energy input is basically supplied from hot flue gases. These are produced in the combustion chamber for direct burn-off of the thermolysis gas vapours and heats the reactor body from outside the mantel. Heat transfer happens indirectly through the mantel to the material.
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THE HEAT TRANSFER CONTAINS THREE PHASES: Primarilv:
Energy transfer from heat under the material in the rotary luln. The material, which is ground down to have extended surface area, is permanently mixed, which mixing process is promoted by the blades inside the reactor body. The heated material is to be permanently replaced by cooler materials again and again. Therefore the thermal conductivity of the biomass material is of less importance and can be within a wider range. This is a very important and vital technical factor, particularly in relation to the biomass refuses, which are usually bad thermal conductors and in practice are of mixed physical and chemical character.
Secondarilv:
Radiation heat transfer from the inner top surface of the reactor body.
Tertiary:
Over the temperature of 275 O C (530 OF) an exothermic reaction starts during the decomposition of the material.
The exothermic process is a slow process, therefore the extended thermolysis gasvapour production will not result in an explosive production of gas vapours. The thermolysis gas-vapour production is promptly removed fiom the reaction space by the master fan. The thermal engineering design of the reactor is related to the throughput capacity of the reactor and the extremely qualitative variations of the input material. No matter if the basic material is of organic, inorganic and/or mixed character, the chemical components will be separated at a certain treatment temperature if the boiling point of the primary target contaminant component(s) are under 600 "C.The pyrolysis reactions are not only a sequenced series of reactions, but parallel series of reactions as well, with different levels of energy (1,2). The surface oxygen structures are quite stable at low temperatures below 200 "C,irrespectively of the temperature at which they are formed. However, upon heating up over 300 O C the surface oxygen structures produce C 0 2 and over 500 "C CO and Hz which is completed about 600 "C in vacuum. To provide a successful Thermal Desorption process, three different factors are needed: (1)
Material, which is thermally decomposable.
(2)
High enough temperature that the waste material pass the complicated decomposition phases.
(3)
Long enough reaction time.
CARBONIZATION PHASE: There are four well distinguished phases concerning the TDT-3RTM thermolysis process inside the reactor: (1)
Warm up phase: up to 150 "C - 160 "C. Characterised by the evacuation of the fiee and most of the bounded water fiom the material.
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(2)
(3)
(4)
Thermal decomposition phase: up to 270 "C - 280 "C. Characterised by heavy discoloration of the material, and the evacuation of the remaining chemically bounded water, with simultaneous development of gasification. Thermal desorption phase: up to 380 "C - 400 "C. Characterised by self carbonisation with exothermic chemical reactions, heavy gasification process and escape of volatile compounds from the material. Thermolysis gas-vapour is continuouslyremoved. Stabilisation phase: up to 600 "C. Remaining volatile compounds are completelyremoved from solid phase.
During,carbonisation phase the basic material is gas out, e.g. devolatized. This process opens the pores where the distribution of the micro pores, meso pores and the macro pores are approximately reflections of the pore characteristics of the original basic material. This is most important that the carbonisation phase perfectly devolatize the basic material, through to core structure, as imperfect devolatization may result internal micro condensation of the tar, which may significantly decrease the expanded pore developments.
Figure 1.: TDT-3Rm rotary kiln
-
ACTIVATION PHASE:by steam - water vapour activation for production of active sites and wider pores. In OUT testing program for biomass carbonisation we have experienced that medium temperature water vapour activation rather promotes meso and macro pore development,than micropores. The production of meso and macro pore 1656
structured activated carbon is our prime goal, as many of the ground water decontamination projects are related to hydrocarbon contamination’s, where the contamination is consisting of large organic molecules to be removed. Chemical activated Field Demonstration Plant “FDP” tests has also been executed for the production of activated carbon from sawdust, involving mixing an inorganic chemical compound with the carbonaceous raw material such as activating agents phosphoric acid and zinc chloride at lower temperature. However, use of zinc chloride poses the danger of zinc traces in the end-product, therefore - although technically is available - steam activation is used. The true reductive environment avoids ash creation, whtch ash may decrease the end product quality. There are some chemically bounded oxygen in the basic material and between the pressed feedstock as well, which will be directly burned off during the carbonisation process. In our experience the total ash content at the end product will remain at a range of approx. 3-6 %, depending also on the basic material natural inorganic composition characteristics as well.
Table 2 Biomass Raw Material Based Activated Carbon End Product Minimum Quality Characteristics (taken from industrial production] Iodine Adsorption BET N2 Surface Area Volatile Content Methylene blue adsorption Moisture Ash content
mfg In
k
% g/ 1oog % %
800 750 >I 10 2 3-6
POST TREATMENT, PROCESS RESIDUAL MATERIAL MANAGEMENT Carbon end product is selected by an TDT-3RTM designed adjustable air selector in any size ranges less than 5 mm, whereas the lowest limit starts at 0-63 micron sizing range. The characteristics of the TDT-3RTM emissions, whereas the process avoids: (1)
Creation of flux of soot and particles into gas-vapour phase, and
(2)
Creation of high off-gas volumes, which are costly to treat, and
(3)
Non-perfect burn out of organic components in the post combustion phase by reaching at least 850 OC for 2 sec. true residence time, and
(4)
Flux of soot and particles into gas-vapour phase, and
(5)
Creation of NOx, SOX, CO and C02.
The TDT-3RTMtechnology meets the US. RCRA Miscellaneous Units 40 CFR 264 Subpart X for THERMAL DESORBERS and compatible E.U. Norms for THERMOLYSIS with the following main characteristics: (a)
Thermal Desomtion Chamber Indirect-fired heat source used for primary desorption chamber.
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(b)
Relatively low operating temperature. Air Pollution Control Devices (APCDs) Non-destructive APCDs used.
(c)
Waste Residual Management Treatment of residuals is separate from the thermal desorber.
(alprimary desorption chamber, @)condensation or burning of pyrolysis The TDT-3RTM gas vapours and (c)non destructive APCD off gas scrubber are separate devices, whereas ( 1)treated solids, (2)condensate residuals, (3)APCD residuals, (4)organic air emission, (5)metal air emission, (6)acid gas emission treatment are according to all the relevant comprehensive U.S. and E.U. regulatory requirements for Operational Control, Residuals and Air Emission Parameters.
Figure 2.: TDT-3RTM Field Demonstration Plant
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Table 3 Off-gas Emissions:
Units
(273 K, 101,3 Way11 % 02) IPR 5/3 U.K. 1996 17. TDT-3RTM EU Dir. 89/369/EEC BlmSchVUGer Max. Values many
Dust
30
10
2
THC (VOCl
20
10
HCl
30
10
1 1
HF
2
1
091
SOXas SO2
300
5
NOx as NO2
350
co HE.Cd
100
50 100 50
0.1
As. Cr. Cu, Pb. Ni, PCDDRCDF
50
0,05
0.01
1
0,05
0.001
1
091 non
0.001
0.1
Target
100
detectable
non detectable
THE STATUS OF THE TDT-3RTM TECHNOLOGY The status of the TDT-3RTMtechnology today is at the end of innovative phase and early commercialisation phase. Full scale applications are prepared on detailed execution manufacturing level. FDP has been successfully tested with 0,5 m3/h continuous throughput capacity. Based on the FDP experiences, up until the end of the go's, 0,8 m3/h and 2,5 m3/h throughput capacity plant has been completely and detailed engineered and designed, for which project applications are under progress.
ECONOMICAL ASPECTS: (1)
At least 30 % decreased production cost even at smaller capacity installations.
(2)
Provides high energy recovery profile.
(3)
Low investment and 0 & M cost provides short amortisation time even at smaller capacity installations (approx. 3 years).
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CONCLUSfONS The TDT-3RTMis utilizing rehse biomass for value added production of activated carbon. Indirectly heated horizontally arranged rotary kiln equipment is used for low temperature carbonization. Steam activation is used for activation. Gas-vapour phase either condensed, where bio-oil is utilized and noncondensable gases burned out, or directly burned for onsite energy recovery. The design is prepared to meet the new US.and E.U. technical requirements, environmental norms and standards. Complete solution, no need for offgas, process water, solid residue post processing. The process is characterized by flexible operation and simple maintenance. The process is characterized by modular design. The TDT-3RTM does not containing exotic technical solutions and construction materials.
is patented, original solution. The TDT-3RTM
REFERENCES 1. 2.
Bansal R. C., Donnet J., StoecMi F., (1988) Active Carbon McEnancy B., Dovaston N., (1975) Carbon
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Workshop Reports This paper is a collation of the reports from the eight workshops held during the conference: Analysis and characterisation of pyrolysis liquids 1. 2. CHP systems 3. Combustion 4. Economics of bio-energy systems 5, Fundamentals of pyrolysis 6. Gas clean-up 7. Gasification tar management 8. Research needs
ANALYSIS AND CHARACTERISATION OF PYROLYSIS LIQUIDS A. Oasmaa' and D. Meierb
' VTTEnergy,P.O.Box 1601, FIN-02044 YTT, Finland IWC, Leuschnerstrasse 91, 0-21031 Hamburg, Germany INTRODUCTION The first IEA (International Energy Agency) thermochemical Round Robin was organised in 1988 as part of the IEA Voluntary Standards Activity led by BC Research (1). The main conclusions were: the precision for carbon was excellent, while hydrogen, oxygen by difference and water were more variable, and oxygen by direct determination was poor. It was recommended to use a wider variety of samples in the future studies. Since then considerable progress has been made both in the field of oil production and oil analysis. Therefore, two separate Round Robins were initiated in 1997: one withm EU PyNE (Pyrolysis Network) and the other w i h IEA PYRA (Pyrolysis Activity). The objective of the EU PyNe Round Robin was to compare existing analytical methods without any restrictions. Two pine pyrolysis oils were analysed by eight laboratories for viscosity, water, heating value, elemental analysis, pH, solids, and density. The accuracy for hydrogen, water by Karl-Fischer, and density were good. The xylene-distillation method was stated to yield erroneous results. High variations were obtained for nitrogen, viscosity, pH, and solids. Ethanol was concluded to be more suitable for solids determination than acetone (2). The main objective of the IEA PYRA Round Robin was to determine the interlaboratory precision and methods applied for elemental composition, water, pyrolytic lignin and main compounds. Two poplar oils were analysed by the IEA PYRA participants. It was concluded that the precision of carbon and hydrogen was very good, liquid sample handling plays a very important role in the C, H analysis, water by Karl-Fischer titration was acceptable, but should be checked carefully, and the method for the determination of pyrolytic lignin should be improved (3).
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The latest Round Robin was carried out by 12 laboratories during January - March 2000. Four pyrolysis oils were provided: BTG softwood mixture (spruce and fi), Dynamotive pine (85% pine, 15% spruce), Ensyn hardwood mixture, and Pyrovac softwood bark (113 fir and 2/3 spruce with traces of hardwood bark). Based on the feedback from previous Round Robins it was decided to make a suggestion for the instructions on handling and analyses. Summary One oil was inhomogeneous due to a high water content causing phase-separation.
This may lead to erroneous results. The oil producer pointed out some problems during production, which now have been at least partly solved. In general, the accuracy of all physical analyses was good. Some results were systematically on high side, which most probably is due to poor calibration of the equipment. Water Water addition method (4) was suggested for calibrating the Karl-Fischer titration method for pyrolysis oils. Solids. It was pointed out that the solids content does not indicate the absolute amount of solids in the oil, because the submicron particles of char present after filtration are difficult to measure. However, it was concluded that this analysis is accurate enough for its present purpose. The microscopic analysis of the oils showed a high amount of small particles below 1 pm. The pore size of filter paper may be reduced down to 1 pm.It was also stated that cohrpounds originated from bark or needles do not dissolve well in alcohols. A more powerll solvent, like methanol-methylene chloride (1:l) can be used. Carbon, hydrogen. nitrogen. It was suggested to compare results with another standard, for example, for fossil fuel, in order to find out the normal deviations of the method. The standard deviations for carbon and hydrogen fall well into these limits (4), and it was concluded that the accuracy of these elements was good. Variation in the nitrogen content of white wood oils was obvious due to the similarity of nitrogen content and detection limit for nitrogen (4). Viscosity. Pyrolysis oils are Newtonian liquids ( 5 ) and hence, kinematic viscosity is applicable. The viscosity results at 20 and 40 OC were very consistent. The smaller standard deviation at 40 "Cis logical because of the high temperature dependency of pyrolysis oils. Consequently, it was also suggested for the convenience of the end-users to determine the viscosity at two temperatures, i.e. 20 and 40 OC. It was pointed out that inhomogeneity of the oil may lead to phase-separation in the capillary tube and hence, to erroneous results. Measuring the viscosity of inhomogeneous oil in a rotaviscotester is suggested.
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Stabilitv index. The stability results varied a lot. Because the stability is measured as a change in viscosity the error in viscosity results yields an error in the stability index. Due to this it was discussed to exclude the stability results of one inhomogeneous oil. A comment was presented that an increase in the water content of the oil improves its stability. Another comment was made on the correlation of viscosity and the stability index. More testing and further development of the stability test are needed. The stability index should not be the only measure of the storage stability of oil. Another simple test method may be necessary. The authors wish to bring out that, in some cases, the instructions had not been delivered to technicians, whch has lead to erroneous results. Hence, the stability results will be re-checked. It was commented that soot stabilises emulsions. On the other hand, soot in the pyrolysis oil most probably yields high particulates emissions. Chemical characterisation. The results of chemical characterisation were not very consistent. It was discussed to prepare standard solutions with known amounts of compounds for quantitative analyses. A question was also raised if the functional groups in pyrolysis oil should be analysed analogously with petroleum residues instead of individual compounds, as by quantitative "C-NMR. The amount of PAH (polyaromatic hydrocarbons) was extremely high for one pyrolysis oil, and it was discussed that more attention should be paid to the analysis of toxic compounds in the oils. The oil producer commented later that the high PAH m a y be due to contamination of other fuel and this will be checked. Pvrolysis oil specification. Fuel oil specifications (6) were only discussed very briefly because oil end-users were not present. pH was suggested to be added in order to help in material choices. It was mentioned that the volatility index or boiling point range distribution would be valuable. However, it has been proved that the boiling point cannot be determined for pyrolysis oil because of the thermal instability of the oil (7, 4). More feedback from end-users is needed for assessing the maximum allowable variation for each property. CONCLUSIONS In general, the accuracy of all physical analyses was good. Additionally, it was concluded that inhomogeneous oils may give erroneous results, at least considering lunematic viscosity and stability index. Good laboratory practice, like proper calibration of equipment, prevents systematic errors. Further, a high standard deviation may be due to the fact that the value is too close to the detection limit of the equipment. This is typical of the case with nitrogen. The main conclusions of the Round Robin were: Karl-Fischer titration is recommended for analysing water in pyrolysis oils. Solids content as ethanol insolubles is accurate for white wood oils but a more powerful solvent, like a mixture of methanol and methylene chloride (1 :1) is needed for extractive-rich oils. For the elemental analysis at least triplicates are recommended. Kinematic viscosity is an accurate method for pyrolysis oils. Stability index needs more clarification and testing: Results of chemical characterisation were not very consistent. It may be necessary to prepare standard solutions with known amounts of compounds for 1663
quantitative analyses. The final results of the Round Robin will be published in PyNe final report in April 200 1. The main conclusions were as follows: Karl-Fischer titration can be recommended for analysing water in pyrolysis oils. The method of water addition is suggested to be used for method calibration. The solids content using ethanol as a solvent is accurate for white wood oils. However, if the feedstock contains extractives, for example, from bark andor needles, a mixture (1:l) of different solvents, like methanol and methylene chloride is recommended. For elemental analysis, at least triplicates are recommended. Nitrogen is not very accurate for white wood oils due to its low amount and the equipment limitations. Kinematic viscosity is applicable to pyrolysis oils because of its accuracy and Newtonian behaviour of pyrolysis oils. However, method calibration should be taken care of. Stability index needs more testing and its correlation with the water content and original viscosity should be studied. Another simple test method for stability may be needed. The results will be re-calculated because the determination was in some cases erroneous. In case of inhomogeneous oils, some analyses like kinematic viscosity and stability index cannot be applied due to possible phase-separation of oil during viscosity determination. The results of chemical characterisation were not very consistent. It may be necessary to prepare standard solutions with known amounts of compounds for quantitative analyses.
REFERENCES
Mckinley, J. W., Overend, R. P. & Elliott, D. C. 1994. “The ultimate analysis of biomass liquefaction products: The results of the IEA Round Robin #1”. In: Proc. Biomass pyrolysis oil properties and combustion meeting, 26 - 28 September 1994, Estes Park, CO. Golden, CO: NREL. Pp. 34 53.(NREL-CP430-72 15.) Meier, D. Technical PyNE Group Report. Characterisation and analysis. In: Minutes of 5th PyNe Meeting 28th February to 3rd March 1998, Salzburg, Austria. Bridgwater, A,, Czernik, S., Diebold, J., Meier, D., Oasmaa, A,, Peacocke, G., Piskorz, J. & Radlein, D. “Fast pyrolysis of biomass: A handbook”. Newbury: CPL Press., 1999.188 p. Oasmaa, A., Leppilmiiki, E., Koponen, P., Levander, J. & Tapola, E. “Physical chiuacterisation of biomass-based pyrolysis liquids. Application of standard fuel oil analyses”. Espoo: VTT Energy, 1997.46 p. + app. 30 p. (VTT Publications 306.) Leroy, J., Choplin, L.& Kallaguine, S. “Rheolological characterizationof pyrolytic wood derived oils: Existence of a compensation effect”. Chem Eng. C o r n , 1988, VOI.71, pp. 157 176.
-
-
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6
7
Diebold, J. P., Milne, T. A., Czermk, S., Oasmaa, A., Bridgwater, A. V., Cuevas, A., Gust, S., H u m D. & Piskorz, J. “Proposed specifications for various grades of pyrolysis oils”. In: Bridgwater, A. V. & Boocock,D. G. B. (eds.). Developments in Thermochemical Biomass Conversion, B d , 20 - 24 May 1996. Vol. 1. London: Blackie Academic & Professional, 1997.Pp. 433 - 447. Meier, D., Oasmaa, A. & Peacocke, G. V. C. “Properties of fast pyrolysis liquids: status of test methods. Characterisation of fast pyrolysis liquids”. In: Bridgwater, A. V. & Boocock, D. G. B. (eds.). Developments in Thennochemical Biomass Conversion, Banff, 20 - 24 May 1996. Vol. 1. London: Blackie Academic & Professional, 1997. Pp. 391 - 408.
CHP SYSTEMS B Jenkins a and J Brammer a
Dept. of Biological and Agricultural Engineering, One Shields Avenue, Davis, California, CA 95616-5294, USA Aston Universig, Aston Triangle, Birmingham 8 4 7ET, UK
DISCUSSION: This workshop was held to discuss problems and opportunities in the development and application of biomass fuelled combined heat and power systems (CHP), and to make recommendations for research, demonstration and implementation of CHP systems. Although the development of cohesive and coordinated policies relating to CHP was recognized as being important, the discussion focused on technical issues and recommendations. Main findings of the discussion: The need for comarative studies: Studies have been done of the general feasibility of biomass CHP. However, there was a general consensus that more comprehensive comparisons were needed of the technical and economic performance of biomass fuelled CHP in competition with other energy systems, such as natural gas distribution for residential heating. The need for ODtimisiner CHP: Optimisation was seen as being important to CHP implementation to address issues of scale, heat and power partitioning, technology selection, fuel selection, fuel flexibility, quality of heat and power, valuation of heat and power, and the potential use of biomass in existing CHP systems using other fuels. The optimisation should be performed on the basis of economic performance for the various technology and fuel combinations, including the range of scales from the very small (e.g., farmer operated or other systems running Stirling engines, microturbines, organic Ranlune cycles, or reciprocating engines fuelled perhaps with pyrolysis oil), to larger systems including co-feeding existing systems. Optimisation studies are needed to assist in decision-making relating to the specific configuration of any CHP system.
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The need for reliable and economic small scale svstems: Although included under the general category of optimisation, there was a general consensus of the need for more effort on small-scale or micro-scale CHP systems. In part this stems from the perception that many small-capacity heat loads exist, and hence there are opportunities available if appropriate technologies can be identified. Current technologies were considered to be too expensive in general. Greater research and development is needed for the small-scale systems (< 500 kW,, < 2 MW,) to reduce cost, even if that means accepting a lower efficiency. The need for fuel standards: Standards defining the quality of biomass fuels and fuel blends are lacking, especially in relation to application in CHP systems. Technical and economic performance is improved when biomass fuels can meet specifications of the design, and designers need good information on fuel properties. Standards should address biomass quality, as well as processing and upgrading of biomass (e.g., pyrolysis oil, pellets, sewage sludge dewateringlconversion). The absence of environmental concerns mecific to CHP: CHP systems were not perceived to have environmental concerns different from other power and heat systems.
RECOMMENDATIONS OF THE WORh3HOP: 1. Criteria should be established to aid decision making on the selection of a CHP 2.
3.
4.
5.
6. 7. 8. 9.
system as opposed to separate generation of heat and power. Studies should be made of the economics of building biomass-fuelled district heating networks for new residentiallcommercial communities, and in particular of the comparative costs of these and natural gas distribution networks for the supply of fuel for heating. Protocols or methods should be developed for optimising biomass CHP systems, with particular reference to: 0 costs for electricity and heat 0 whether or not to upgrade fuels 0 potential economies of scale in capital and operating costs 0 the use of fuel blends Studies should be continued into the properties of biomass fuels and blends, with reference to mitigating the impacts of fouling, corrosion, and other detrimental effects on plant efficiency, operation, and maintenance (particularly in co-fired systems). Increased research should be encouraged on small-scale biomass CHP systems (< 500 kW,), especially low-maintenance,simple, fuel-flexible systems. Biomass CHP systems should be demonstrated at the commercial scale, including the potential production of chemicals as co-products of the heat and power system. Partnerships should be encouraged between developed and developing countries for the transfer of biomass CHP technology. A state-of-the-art review of biomass CHP systems should be prepared to aid in information dissemination and new program development. The summary of recommendations from all the conference workshops should be made available as a stand-alone document targeted at policy makers and government program managers.
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COMBUSTION L Baxter and B Jenkins a
Sandia National Laboratories, 7011 East Avenue, MS 9052, Livermore, CA 94550, USA Dept. of Biological and Agricultural Engineering, One Shields Avenue, Davis, California, CA 95616-5294, USA
SUMMARY OF DISCUSSION: A brief discussion was held of issues relating to biomass combustion. Main topics
included Principal emissions of concern are particulate matter and NOx, and work should continue on developing means to reduce these pollutants from combustion systems. Co-combustion of biomass with natural gas or gas from biomass gasification offers potential efficiency improvements relative to separate combustion of biomass, as well as potential emission reductions through reburning. Combustion of waste wood presents special problems from a regulatory standpoint, and research should continue on reducing pollutant emissions. There is some potential for blending with clean wood without substantially altering emission control systems in existing units. Better fuel characterizationis needed for waste materials. Feedstock costs remain high in many cases, and logistics of fuel delivery are seen as a critical area for improvement. Improvements are needed in combustion technologies to unprove combustion stability and to reduce the impact of transients, especially on emissions. Greater effort is needed to take advantage of new developments such as flameless oxidation and other advances. Fuel pretreatment is viewed as a way to improve combustion behavior and reduce a number of operating problem. Greater effort is needed on separation of undesirable constituents and components from biomass fuels. Ash disposal remains a concern and better uses for ash or means to mitigate special or hazardous characteristics of ash are needed. Methods to optimize blending of different feedstocks are needed, including predictions of inorganic behavior in fuel blends during combustion. Good models are needed for the design and operation of combustion system. Expert systems, might, for example, contribute to diagnosis of operating problems or aid in the design of more complex systems. RECOMMENDATIONS 1.
2.
Fuel properties and environmental impacts: a. Conduct research into the environmental impacts of biomass and waste fuels including the impacts of hazardous constituents and components. b. Continue to conduct research into the general properties of biomass with special emphasis on predicting the properties and behavior of fuel blends in combustion system. Fuel logistics: 1667
Improve logistical support systems for using biomass and waste materials, including handling, transportation, blending, and feeding systems. b. Develop test methods, standards, or other protocols for providing quality assurance in biomass fuels delivered to commercial users. Modelling and design: a. Generate comprehensive databases of information from commercial biomass plant operations appropriate for validating combustion models. b. Develop more detailed comprehensive models for better understanding of biomass combustion phenomena. c. Develop engineering models for improved design and operation of biomass combustion systems. a.
3.
ECONOMICS OF BIO-ENERGY SYSTEMS Yrjo Solantausta ',David Beckman ', Ian Burdon ' a
VTT Energy, Biologinhja 5, PO Box 1601, Espoo FIN-02044 kTT,Finland ZETON Inc., 5325 Harvester Road, Burlington, Ontario L7L 5K4,Canada PB Power Ltd (Men and McLellan), Amber Court, William Armstrong Drive, Newcastle upon Tyne NE4 7YQ,UK
SCOPE OF THE WORKSHOP Technoeconomic assessments are an integral part of any process development effort. Assessments vary in scope and size at different stages of process development work. However, it is recognised that uncertainties in input data may often be considerable, which may make use of results difficult. TOPICS DISCUSSED
Comparisons of technologies on consistent basis When new bioenergy technologies are compared to each other, it is of prime importance to carry out all stages of work on a consistent basis. In practise this means of carrying out rigorous performance analysis, sizing, and costing for all alternatives in detail. It was also pointed out that perhaps cost analysis should not be used as a primary criterion, when processes at a very early development stage are assessed. In this case a rigorous performance analysis (and the respective efficiency) may be a better criterion in assessing a fbture potential of a given technology. Objectives of studies and communication to politicians Participants of the workshop expressed that it would be important to convey the messages from bioenergy feasibility studies to politicians and other decision-makers. However, it was not entirely clear in what format and from which platform such
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messages should be addressed. The IEA Bioenergy Executive Committee was mentioned as one possibility. Uncertainties in the estimates When analysing several new processes, typically not all processes are at the same level of development. This makes comparison difficult. In principle different contingencies for capital investment should be used. The problem is that reliable contingency factors are not usually available. One suggestion made at the workshop was to require considerable reduction in cost (for example, half the cost compared to existing alternatives) without any contingencies for new systems. Economic estimates of new bioenergy processes tend to be positive. Reasons for this are that the economic feasibility studies are based on process data that is too uncertain, and results fiom too small scale equipment. Studies assessing the cost of new technologies have shown that the cost of the first demonstration plant and products are on average four times the cost of the original estimate 1.
The IEA Bioenergy techno-economic assessments One platform that the IEA countries have used to help development of bioenergy applications has been the techno-economic assessments carried out during the past 20 years. An expert group has been assembled, which has developed standard procedures for the assessments. Lately the work has been carried out with industry, whch typically has supplied cost data for the assessments. CONCLUSIONS AND RECOMMENDATIONS
A consensus of the workshop was that an engineering analysis should be included early on in the development process. It was recognised that carrying out process evaluations tiom the start of the work may identify unknown aspects in a new process. In fact evaluations should be carried out during the whole time of process development. 1. Studies where the data is uncertain should first concentrate on process performance analysis first, such as efficiencies, and not on economic analysis. 2. Economic feasibility studies should not be done on processes where the basis of the data is too uncertain. General process comparisons can be misleading. It is better to analyze a technology on a specific site and set of local conditions.
REFERENCES
I Menow, E., Chapel, S., Worthing, C., “A Review of Cost Estimation in New Technologies: Implications for Energy Process Plants”. Prepared by RAND Corporation for the U.S.Department of Energy. Santa Monica, CA. RAND-R2481-DOE. July 1979.
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FUNDAMENTALS OF PYROLYSIS B. Krieger-Brockett and J. Piskorz a
‘
Universityof Washington,Benson Hall BF-10, Seattle, WA98195-1 7S0, USA Resource Transforms International Ltd., 110 BaBn PI., #5, Waterloo, ON N2V 127. Canada
INTRODUCTION This is a summary of the PITBC workshop in which we discussed and reviewed selected aspects of hdamental research in the area of thennochemical biomass conversion. The workshop, attended by over 15 researchers from Europe, Americas and Asia, concentrated on needs, trends and outlook for biomass pyrolysis fundamental research. The remarks fell into these categories: 0 the importance and generalities of true fundamental science, as well as the need for fundamental research in the field of biomass thennolysis/pyrolysis; 0 the recent advances in light of a long tradition of biomass utilization studies conducted.during the past century; 0 the importance of communication, in a meaningful and systematic way, of significant accomplishments in the field to decision makers in the forest products and renewable energy industry and to technology developers. “Theformulation of a problem is often more essential than a solution, which may be merely a matter of mathematical or experimental skills. To raise new questions, new possibilities, to regard old questions$-om a new angle, requires creative imagination and marks real advances”. Albert Einstein
DEFINITION Mankind’s fascination with pyrolysis started with the discovery of fire. Pyrolysis is usually understood to be thermal decomposition of organic matter occurring in oxygen-depleted or oxygen-free atmosphere, although there are other less general terms that are used to describe this same process (i.e., devolatilization, destructive distillation, carbonization, liquifaction, gasification). The pyrolysis step is of critical importance for all solid fuels and feedstocks since this step precedes any subsequent steps in any gas-phase combustion and biomass conversion schemes. The pyrolysis step, particularly in so-called “fast pyrolysis” happens in a few seconds or less depending on biomass particle size and heating rate. On such short time scales, chemical reaction kinetics, mass transfer processes, phase transitions and heat transfer phenomena play important roles and can influence the ultimate conversion outcomes. Also, due to the same short time intervals, fundamental aspects of those transformations are not easily elucidated. An important objective is to obtain a comprehensive understanding of significant factors determining biopolymer thermolysis, and therefore to predispose or quantitatively predict certain conversion rates znd products. There is no doubt that benefits to society from such work will come from technological developments and innovations leading to improved renewable fuel utilization, new high value chemicals, a sustainable highquality environment and improved well being of all.
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DISCUSSION
A thread throughout the discussion was the extreme importance of emphasizing the general applicability offundamental experiments and models based on first principles. While the meaning of words such as “basic” and “fundamental” to diverse audiences was debated, there was general agreement that the results we label fundamental are those that have general applicability, that is, they are NOT system-specific. In this endeavour we agreed that vigorous attention must be paid to reporting chemical as well as physical details of our experiments and models in order for other researchers to be able to evaluate and use OUT results. The details that must be measured are numerous. Indeed, they are more numerous for biomass than other solid fuels owing to the high reactivity and heterogeneity of biomass (growth, storage, and reaction conditions make a difference in the conversion product slate). In addition to reaction conditions, we agreed that at a minimum the following reactant chemical specifics and physical parameters of the biomass will be useful for fundamental evaluations: 0
0
chemical specifics: species names, compositions, and preparatiodcollection method physical parameters: particle sizes, particle heat transfer properties (thermal diffusivity), heats of reactions, particle mass transfer properties (effective diffisivity)
It was remarked that our research area has a long tradition based in the forest products industry and pulping chemistry. In those fields, somewhat slower and less thermally severe reactions were employed than we now study in biomass thermal conversion to gases, liquids, and energy. In many contemporary thermal conversion processes, the biomass particles and the fluids in reactors are NOT necessarily at the same conditions of temperature and composition and frequently have temperature and composition gradients in both the particles andfluids. Thus, in fundamental studies, we agreed that it will be usefbl to quantitatively report or estimate all fluid (gas or liquid) and solid compositions, gradients, and temperature-time histories, as well as flow characteristics that determine energy and material transport between the biomass particles and the fluids in the reactor. Applauding the utility of fundamental pyrolysis data and predictions, Drs. Lede and Suuberg both commented that we should periodically critically synthesize our results. In particular, the results should be reported and evaluated in light of systematic approaches r e c o p e d by industry. Of necessity, these approaches incorporate both chemical and physical data and experiments at increasing length scales. It was suggested that reactor theory or the reaction engineering approach is a valuable tool in this regard. This approach systematically synthesizes the knowledge, often using dimensionless groups to categorize behaviour, fiom the molecular length scales (chemical composition), to the microscale, (the small homogeneous porous particles in which no gradients of temperature or composition exist), to the mesoscale (thermally thick, but practical feedstock particles in which temperature and composition gradients dominate), all the way to the industrial reactor length scale that must operate economically. It was further emphasized that economical biomass reactor technology must process a heterogeneous collection of mesoscale particles that are influenced by local fluid mechanics, temperature, and fluid composition. The importance of collaboration with industry was underscored in order that we may incorporate realistic constraints of the technology into the range of experimental
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conditions and simulation predictions we study. For example, we discussed the lack of fundamental pyrolysis predictions suitable for use in practical large fixed or moving bed reactors, large fluidized beds, etc. employing real feedstocks. It was remarked by the industrial representatives that particular attention must be paid in pyrolysis studies to the often very different residence times and states of mixing for the fluids in contrast to the solids. The reaction engineering approach was suggested because it has previously been highly successful in finding novel ways of incorporating validated hdamental information into industrial processes, notably in the field of catalytic reactions, catalyst synthesis, and reactor design. There was considerable discussion on the wealth of historical findings reported in the literature regarding biomass pyrolysis and the devolatilization preceeding biomass combustion. J. Piskorz listed relevant qualitative observations dating from 1914 regarding pyrolysis products from biomass. With such a wealth of literature to draw from, a discussion revealed that we now must emphasize that we can be more quantitative than in the past. In being more quantitative, we must also be more critical and complete in our research in order to put our new findings, and models derived from them, in the proper and meaningful perspective. New instruments allow us to be increasingly quantitative and accurate, but it is our obligation to communicate how this often overwhelming amount of data can be meaningfully used. For example, it was once reported that there were of the order of 20 substances in wood tar when the most advanced (at that time) detectors and columns, perhaps FID and packed columns, were used. Now, with GCMS or HPLCMS systems available, it has been shown that hundreds of compounds are routinely found in tar. However, many of these compounds appear in concert with others or are “surrogates” for other compounds. Thus it is useful to “lump” or simplify the compounds into groups, which appear at the same reaction time, or under the same pyrolysis conditions, or from the same starting sample. These co-varymg compounds also must be distinguished from those that are unique to particular experimental conditions or samples under study. It was suggested that computer-based data analysis techniques (often involving multivariate statistical methods) can aid in this classification or simplification, as has been so profitable in other thermochemical conversion endeavours, for example, as applied to coal and petroleum. Again, it was emphasized that there is a need for a critical synthesis of the wealth of experimental data into regimes of behaviour, and simpler predictive equations or simulations, that are useful to the technologists in industry who are designing industrial scale reactors. Regarding the fundamental, general applicability of what we measure, the workshop attendees agreed that it was important to communicate the importance of our findings, and that perhaps we need better marketing skills. Biomass is in a unique position. It is the only renewable source of carbon-based hels and biopolymers, but it is in direct competition with fossil fuels and other renewable energy (solar energy, wind power, etc). Because the variety of options for renewable resource utilization is so large, the case for bio-based energy or materials has to be made more forcefully and concretely.
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CONCLUSIONS AND RECOMMENDATIONS
General 0
0
0
0
0
Renewable carbon-based biopolymers (biomass) are produced at a large annual rate in the biosphere. Conversion of this biomass by rapid thermal means, although presently in an incubation stage, appears to have a tremendous potential. Emerging renewable energy sources like wind, photovoltaic devices, and hydropower cannot displace the global utilization of biomass. Efforts should be directed to facilitate good interaction between scientific policy formulators, research organizations, and organizations wanting to commercialize pyrolysis conversion processes. Only concerted and cooperative efforts can accelerate the progress.
Specific There is a very considerable literature, knowledge base, and prior art concerning thermal biomass conversion. Some historical examples include: 0
0
0 0
R.C. Palmer (1914) reported increased (up to 42.5 d!) charcoal yields when carbonizing wood under pressure of 150 psig. Klason (1914) accomplished a vacuum distillation of wood (cathode-light vacuum) and a short time. His results - tar yield 43.5wt%, char 19.4 wt%. Pictet, Sarasin discovered levoglucosan in 1918 Bobrow mentioned hydroxyacetaldehydein 1934.
Yet the substantial quantitative research progress of recent years could be easily illustrated by: 0
0
0
0
Identification in pyrolysis oils of literally hundreds of chemicals (from thousands possible); Creation of national and international biomass utilization and conversion databases; Construction of a variety of pilot plants employing biomass pyrolysis; Increasing accuracy and predictive power of pyrolysis mathematical modeling.
Despite h s substantial progress, biomass conversion activities suffer from lack of support by a mature industrial sector (there is none with the exception of the forest products industry, which often has conflicting goals). The oft-mentioned advances in “coal science” were likely due in part to the support of a mature coal-utilization industry. While biomass researchers can profit from models, tools and achievements of clean coal technologies, nevertheless, the government and industrial support is crucial to implementing the research advances in industrial practice. The European Union and Scandinavian countries are fulfilling this obligation to a significant degree already. PITBC and similar conferences have an invaluable role to play in encouraging active discussion, in reporting developments of basic science, and in transferring this know-how to potential industry developers of the hture.
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GAS CLEAN-UP JPA Neeft and N Abatzoglou
ECN - Netherlands Energy Research Foundation, PO Box I, Petten, I755 ZG, Netherlands Kemestrie inc.. 4245 rue Garlock, Sherbrooke, Quebec, JIL 2C8, Canada INTROD UCTION' Contaminants in the product or flue gases from thermochemical conversion of biomass are particles, tars, acid gases (HCl, COS, N,O,, SO,, HCN), basic gases (NH3),and metals (both alkali metals and heavy metals). Several cleaning techniques exist for the removal of these contaminants (I, 2). These techniques can be subdivided into the three classes hot gas cleaning, dry scrubbinglremoval techniques and wet scrubbinghemoval techniques. Table 1 was used as a structure for the discussion in this workshop. On the basis of this table, the following discussion topics were chosen: 1. Catalytic converters for tar 2. Cost of gas cleaning 3. Catalytic removal of NH3 4. Scrubbers / tar removal 5. Particle removal (all techniques)
Two questions were discussed for each of these five topics: a. What are the results with this gas clean-up technology (removal efficiencies, experiences); b. What are the needs for fundamental and applied research; Table 1 Suitability of different gas cleaning methods for different contaminants
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FINDINGS Catalytic convertors The discussion focused on Nickel catalysts. Commercial availability is very good as in Europe alone at least 4 manufacturers provide these catalysts. The pellet catalysts are less expensive than monolith supported catalysts. One of the large uncertainties of these catalysts is their deactivation rate under practical conditions. Deactivation occurs due to reactions with (heavy) metal-, sulfur- and chlorine compounds. Due to the lack of extensive data the costs for applying Ni-catalysed for tar removal in commercial biomass CHP systems are still unknown. The intended application of these catalysts is in the medium to large scale. For small scales, heat recovery has to be applied such as in a reversed flow reactor, otherwise, catalysts can probably not be applied because of h g h costs.
costs The gas-cleanup depends on the biomass fuel, therefore, the cost of the cleanup section is case specific. However, these costs are related to the fuel flexibility: a simple and cheap gas-clean-up system will limit the fuel flexibility and might cause higher overall costs than a more advanced gas clean-up system in which also more contaminated so less expensive fiels can be used.
Catalytic removal of NH3 This topic is covered by a large 5FP EU project ( M E 5 99 00923). This topic did not raise a discussion but is seen as an important field of R&D in particular for the development of hot gas cleaning in biomass-IGCC.
Scrubbers /tar removal Scrubbers for tar removal have a large disadvantage, which is the tar condensate and scrubber water. Several institutes have experiences with scrubbers (mostly wash towers and Venturi scrubbers), however, hardly any data can be found in open literature. Wash towers only remove part of the tars (typically 40%) whereas Venturi scrubbers can have much higher efficiencies of 80% or hgher. Techniques to clean the scrubber water have been researched. Proposed techniques for water cleaning are destruction by wet oxidation and absorption by activated carbon. Reference was made to a 1998 Verenum report (3). Scrubber solvents other than water have occasionally been used to remove tars fiom producer gases, the T e c h c a l University of Vienna has recent experiences (see also the presentation by Hermann Hofbauer in these proceedings). The requirements for further development are the treatment or disposal of (tar containing) waste water, and the efficiency of scrubbers when scrubber solvents other than water are used.
Particle removal (all techniques) Many techmques are used to remove particles fiom biomass fuel gases (see table), but only two techniques can possibly remove submicron particles: electrostatic
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precipitators (ESP) and rotating particle separators (RPS). There is hardly any experience with RPS for particle removal from biomass fuel gases. Particle removal at high temperatures is possible with ceramic or sintered metal filters as has been demonstrated in the Vamamo biomass gasification plant. Filter blinding (reaction of tars in the filter pores) has occasionally been found to be a problem when testing high temperature filters. This phenomenon is, however, not reproducible nor understood; Tars can also be removed as particulates (aerosols) which has only been shown for pyrolysis or updraft gasifier tardoils using ESP's. Experience with the use of ESP for removal of tar aerosols from downdraft or CFB gasifiers is non-existing or is not reported. NEEDS FOR FUNDAMENTAL AND APPLIED RESEARCH 0 0
0
0
0
0
0
Assess the deactivation rate of nickel catalysts under practical conditions; Compare costs of gas cleaning in biomass gasification systems and assess the influence on fuel flexibility on these costs; Develop the catalytic removal of N H 3 which is important particularly in hot gas cleaning for IGCC application of fuel gases (activities are being performed in 5FP EU project); Further develop treatment or disposal techmques of tar containing waste water; Evaluate the efliciency of scrubbers for tar removal using water and other scrubber liquids; Assess the efficiency of techniques for removal of submicron particles from biomass fuel and flue gases, and of removal of submicron tar aerosols from biomass fuel gases; Learn to understand and predict blinding of high temperature ceramichintered metal filters;
REFERENCES 1
2 3
T. Milne, N. Abatzoglou and R.J. Evans: Biomass gasifier "tars": their nature, formation and conversion, Golden, CO (USA), NREL, NREL/TP-570-25357, 68 p. (1998). J.P.A. Neefl, H.E.M. Knoef and P. Ojani:Behaviour of tar in biomass gasification system. Tar related problems and their solutions, BTGECN, Novem-EWAB no.9919,74 p. (1999). P.Hasler, P.Morf, R.Biihler and T.Nussbaumer: Gas cleaning and waste water treatment for small scale biomass gasifiers, Ziirich (Switzerland), Verenum, 102 p. (1998).
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GASIFICATION TAR MANAGEMENT Nicolas Abatzoglou a and Krister Sjostrom
' Kemestrie inc.. 4245 rue Garlock, Sherbrooke. Quebec, JI L 2C8, Canada Royal Institute of Technology, Dept Chemical Technology, Teknikringen 42, StockholmS-10044, Sweden INTROD WCTION
The Workshop focused on the utilization of catalysts at high temperature for tar transformation to combustible gases. The following points have been retained DOLOMITES
Dolomites: This naturally occurring catalysts are relatively efficient in converting tar to non-condensable combustible gases (CO, Hz) and CH4, which is not reformed significantly. The eficiency vanes from 50-90% depending on both tar and dolomites composition. Dolomites age quite rapidly and replacement is required if efficiency is to be kept at the initial level. A fluidized bed catalytic reformer based on the use of dolomites has been commercialized by TPS. The main problems associated with the use of the dolomites are: 0 0
0 0
0
Aging due to coke deposition Fast erosion leading to sue decrease and, in the case of the fluidized bed, entrainment out of the bed Decomposition of carbonates due to high C02partial pressure Poisoning due to the deposition of C1 and S , two hetero-atoms responsible for the loss of the catalytic activity of the Ca and Mg oxides High C1 retention in the dolomite (i.e. when gasifying high C1 content straw) leads to the formation of MgClz and CaC12; these salts are melting at relatively low temperatures and can also form eutectic mixtures with other inorganic material (i.e. silica sands). Thus sintering can occur at relatively low temperatures.
METALLIC CATALYSTS
A wide variety of metallic and bi-metallic catalysts are available in the market. The most prominent are nickel catalysts supported in various matrices, namely acidic AlZO3. These catalysts are very efficient in reforming even heavy tar to CO and H2. Simultaneous partial reforming of the synthetic gas cannot be avoided but it is not usually detrimental to the gas quality. These catalysts are used in pellets, spheres or even deposited in monoliths (i.e. honeycomb matrices). In all cases the catalytic reactor configuration is a fixed bed running as a perfect Plug Flow Reactor (PFR). Fluidized bed configurations, as those in FCC systems utilized by the petrochemical industry for the production of reformulated gasoline, are not yet known for this application; the reason being that such a reactor involves an entrained bed configuration for continuously regenerating the catalyst and catalyst erosion becomes a considerable issue.
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The fixed-bed are used in twin-beds configurations, where one of the vessels is operated as reformer while the other is in catalyst regeneration mode. Regeneration is achieved through mild conditions combustion of the deposited coke and S, or through steam reforming of the deposited coke. A new avenue under study is the use of honeycomb monoliths and granular catalytic media both as reformers and particulate filters. Experimental work is underway in Sweden and Canada respectively. Tar reforming has not yet reached the commercial application level; mainly because of the high cost and some peculiar technical problems associated with the use of the existing catalysts. These techcal problems are reported below: b
0
0
The high temperature required for an efficient reforming (typically higher than 800°C) is an important aging factor for the catalyst; the supported metallic catalysts exposed at high temperature fields for long periods of time in presence of organic compounds are subjects to structural changes. The latter are responsible for the crystal lattices collapsing and the loss of the catalyst particles cohesion. Thus the catalyst tablets, spheres or the monolith are transformed to granules or powders, are subsequently entrained by the gas and lost. Coke formation, even though minimized by appropriate catalyst formulations during the last decade, cannot be avoided completely. Typically after 2-3 days on stream the catalysts require coke removal to regain their initial catalytic activity. Such regeneration cycles are introduced using twin-beds configurations as already explained. After each regeneration cycle the catalyst shows a slightly lower efficiency. The number of the possible regeneration cycles before the catalyst being considered useless determine the lifecycle of the catalyst and consequently the cost of the reforming process Heteroatomic contamindnts present in the gas and the tar, like C1 and S, react with the metals at the surface of the catalyst. The process is known as catalyst poisoning because the salts formed are not showing catalytic activity towards tar reforming. Some of this poisoning, basically the one associated with the S can be reversible because high regeneration temperatures can transform the sulfides to sulfates and subsequently decompose the sulfates to S 0 2 / S 0 3 and the corresponding metal oxides. This is not the case with C1. Volatile metals, mainly alkalis, can also deposit on the surface of the catalyst and cause sintering or simply perturb the necessary for the catalytic action electric charges of the surface of the catalyst. This leads to a second poisoning due to chemical poisoning.
It is obvious that the presence of the C1 and S, as well as of volatile metals, is detrimental to any tar reforming catalyhc system and that the success of this process pass through a successful removal of these contaminants prior of entering any catalytic reactor. C1, S and metals getters have been already developed, tested and applied by the coal gasification industry. Nevertheless the biomass gasification field need to address this issue the faster possible in order to allow hot gas conditioning options combined with the use of gas turbines (i.e. in IGCC) to have success in the future. We must remember that the hot gas conditioning option is very promising for the following two reasons: 0
Gas turbines efficiency is higher with high temperature gases
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Hot gas conditioning avoids the use of water scrubbing thus it does not create any contaminated liquid effluent which would need additional treatment prior to its final disposal.
0
CATALYSTS FOR METHANOL SYNTHESIS
Commercial catalyst are available for the production of methanol and other liquid fuels fiom synthetic gases. The main problem is the catalyst deactivation due to chemical poisoning from chlorine and sulphur. The following conclusions were reached: Need for R&D in the field of C1, S and metals removal om hot synthetic gas stream Need to test various existing catalysts at, at least, demonstration scale, in order to establish their lifecycles and consequently successfully address the technical barriers and economics of their utilization Need for R&D in terms of developing catalysts that are robust enough to resist thermal restructuring (i.e. exceptional spinels) Need for R&D in terms of catalysts regeneration (steam or controlled combustion). Appropriate profiles must be established as function of the catalyst composition and contamination level.
1. 2.
3. 4.
Table 1 gives a summary of the existing catalytic systems for tar removal as presented in 1998 by Abatzoglou in a meeting of the IEA Gasification Task Team. Table 1 Summary of existing catalytic systems for tar removal (Abatzoglou) PAH 0
0
0
0
High temperature complete oxidation Thermal cracking (temperature well above 1100 "C); soot production; unsatisfactory conversion Catalytic cracking (enhanced by high temperature) 0 Zeolites (not efficient above 850 "C) Dolomites 0 0 Nickel catalysts 0 Lewis acids Catalytic steam reforming 0 Basic matrices (Mg0-CaO)/Dolomites 0 Alumina supported Ni and NiO catalysts
PCDD/F High temperature complete oxidation 0 Low temperature (300 400 "C) catalytic oxidation 0 Catalytic cracking (enhanced by high temperature) 0 Zeolites 0 Lewis acids 0 Catalytic steam reforming 0 Basic matrices (MgO-CaO)/Dolomites 0 Alumina supported Ni and NiO catalysts 0
-
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RESEARCH NEEDS IN THE AREA OF BIOMASS A Segerborg-Fick and S-Y Yokoyama a
European Commission, DG Research, Rue de la Loi 200, Brussels B-1049, Belgium National Institute for Resources and Environment, Tsukuba, Ibaraki 305, Japan
RESEARCH STRA TEGIES AND POLICIES Thls workshop was attended by eight people, which shows that the main interest of this Conference audience was not about research strategies and policies. This was an interesting observation as most of the delegates were researchers themselves. Half of the group was from outside Europe, mainly from Asia (Thailand and Japan). The discussion began by considering the European Research Area (ERA) that is a view on how to organise research on a European level in which creating centres of excellence is the main topic. The concerns on a European level is that there are now many different research programmes instead of emphasising collaboration with the resources and knowledge base available. It was interesting to discuss the similarities with, for example, MITI in Japan and US research programme. USA has had experience with centres of excellence for a long time in biomass research and the outcome is not always positive. One of the problems is that a huge amount of money is centralised at one place. This means a large increase in administration. There is also a risk that the flexibility to change research directions and content gets more difficult in large research groups. Generic research and unknown research groups could lose out in this type of research organisation. Some European countries were worried that within this ERA discussion, they would lose their research position in the biomass area because they would need to share money and knowledge instead of concentrating on being the best in their field. RESEARCH AREAS IN THE BIOMASS FIELD There was also discussion on what research areas should be emphasised. Feedstock security is a main issue in the biomass area. Therefore pre-normative research is very important. Gasification for electricity production on both small-medium and large scale is interesting when co-generation is used. Gas cleaning is an important area to use research resources. Generic research is very important for understanding the chemical reactions in the gasification processes. The knowledge can be used for improving efficiency and the overall emissions. Emissions reduction for all kind of biomass incineration is important to satisfy current and future legislation. Ash cleaning and recycling to replace the nutrients removed from the soil is an important area for research. In particular, the ash has to be clean and without heavy metals and other contaminants.
CONCLUSIONS One of the conclusions from this workshop were that not all research should be done on a centralised level like centres of excellence. There is an impending risk that innovative research will be lost. One idea would be to give a small amount of money
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to a lot of researchers and let them prove their slalls. The ones that succeeded should then continue to get more money and build up their promising research. There was also criticism in joining well-developed research with less developed. Countries in the biomass frontline were worried that their research would lose momentum and not develop in a positive way. Research money should be put on the following subjects: Gasification including co-generation, Gas cleaning with generic research including chemical reaction, Feed stock security, Pre-normative research, Emission reduction for all types of thermo-chemical conversion of biomass, Ash recycling and cleaning. RESULTS The main result of the workshop was that a European company found a partner in Thailand to help the to solve their bagasse problem. This shows that Europe already has a lead in RTD and that technology transfer can be achieved.
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Author Index
Aarna I, 1246 Abatzoglou N, 365, 1661 Abeln J, 109 Adomi M,499 Aksenov DG, 1207 Andersons B, 1550 Andries J, 473,1158 Anselmo E, 693 Antal MJ Jr, 1179 Arauzo J, 252,346 Arvelakis S,564 Assink D, 298 Axell M, 743
Brown RC, 379 Bruch C, 585 Bufinga GJ, 162 Burdon I, 998, I66 I Cabanas A, 929 Cabanillas A, 137,209 Cannon M, 524 Carrasco J, 1268 CenK, 1091 Cerrella GE,I 1 16 Chaala A, 1349 Chen G, 1158 Chen Y, 465 Chiaramonti D, 1525 Chirkova J, 1550 Choi C-S, 358 Chomet E, 365, 1577 Cillero E, 137 Clemens A, 630,758 Coelho ST,843 Coll R, 1540 Connor MA, 1603 Corella J, 333, 887 Cukierman AL, 11 16 Czemik S,977,1577
Baeyens J, 766 Baglioni P, 1525 Bai XS,908 Baklanova ON, 1509 Bandi A, 1459 Bamo M, 32,47,426 Baumgart F, 1459 Baxter, L, 1661 Beck RW, 1197 Beclanan D, 831,867,1661 Beckmann M, 564 Bellais M, 1129 Beran R, 499 Beregovtsova NG, 1388 Berg DA, 8 12 Berg M, 322,875 Berge N,875 Berger R, 656 Bermudez RAP, 1 143 Bilbao R, 346 Biollaz S,388,573 Blanchette D, 1296 Boenigter H, 488 Bonelli PR, 1 116 Bonini M,1525 Boocock DGB, 1517 Boukis I, 1259 Boukis N, 109 Boutin 0, 1034 Brage C, 162 Brammer J, 298,307,1661 Branca C, 1 143 Brandin J, 875 Bridgwater AV, 298,307,977,1281, 1482
Daavitsainen JHA, 705,779 Daey Ouwens C, 488 Dai X, 1 179 Dam-Johansen K, 1061 Davidsson KO, 1 129 de Caumia B, 1296 de Jong W, 473 De Ruyck J, 599 de Vries R, 799 Della Rocca PA, I 1 16 den Uil H, 488 Di Blasi C, I 143 Di Felice R, 188 Dinjus E, 109,221 Dizhbite T, 1 171 Dobele G, 1500,117 1 Dorrington M,162 Domnsoro JL, 929 Dote Y,956 Drozdov VA, 1509 Dudouit C, 1364
1683
Duplyakin VK, 1509
Henriksen U,32,92 Herdin G,499 Hernandez EG,1143 Himmelblau A, 1197 Hofbauer H, 199,64 1 Hofmans H, 272 Honsbein D,766 Hooper RJ, 630,758 Hoppesteyn P, 473 Houbak N, 92 Huang H, 465 Hugener M, 150 Huijnen H, 272 Hustad J E, 32,47,61,426
El Asri R, 599 Elliott DC,1 186 Escalada R, 929 Espenb, B-G,322 Faix 0, 1171, 1405,1500, 1550 Fang Z, 396 FeikC, 1577 Fossurn M, 426 French R, 1577 Fujino J, 964 Gaegauf CK,614,656,896 Gagnon M,365 Ganesh A, 1025 Gansekoele E, 1586 Garcia L, 346 Garcia M, 1349 Garcia-IbaAez P,209 Garcia-Ybarra PL,209
Imarnura Y, 1396 Ino& S, 1219,1326 Irbe I, 1550
Jackson G, 44I Jacoby WA,1540 Jenkins BM, 713,1661 Jensen A, 1061
Gea G, 252 Gehrmann H, 564 Generalis SC,128 1 Gerdes Ch, 1374 Gerhauser H, 1281 Gifford J, 758 Glaser G, I517 Gebel B,32,92 Goldschmidt B,524 Gornez F, 929 Gong D, 630,758 Gonzalez A, 929 Goudriaan F, 1312 Grammelis P,789 Greil C, 162 Griselin N, 908 Grenli M, 1179 GubynskyyM, 1213 GustafssonTE, 122 Gustavsson L, 743 Gutikrrez M, 887 Gyftopoulou ME,1259
Kajimoto T, 1396 Kakaras E, 789 Kaminsky W, 1374 Kastelein R,272 Kato A, 237 Kehlenbeck R, 188 Kelsall 0, 524 Kenten SRA, 452 &el JHA,272 Kikuchi H, 1396 Kim S-B,358 Kircher K, 693
Kluth M, 109 Knoef H, 162 Konar SK, 1517 Konings AJA, 799 Konishi R, 1338 K o m v AA, 599 Koukios EG, 564 Krieger-BrockettB, 1011, 1661 Kruse A, 109 Kunstner H, 1452 Kurkela E, 122 Kuznetsov BN, 1207,1388 Kyt6 M, 1468
Hague RA, 1281 Hajaligol M R, 1226 Hansen U,1452 Hansson S, 536 Hargitai T, 875 Hasler P, 150, 162 Hata T, 1396 Heginuz E, 322 Hein KRG,473,656,896, 1433 Helsen L, 1417 Hem'ch E, 22 I
Laatikainen-LuntamaJ, 122 Laitinen RS,671,705,779 Larfeldt J, 1046 LauerMJ,851 Leal MRLV, SO9
1684
Leckner B, 743,1046 Ud6 J, 1034 Lee K-W,358 Lee S-W, 358 Leung Y C, 1158 Li J, 630,758 Liliedahl T, I129 Lindman E-K,867 LoMer G, 64 I Lyberg M,824
Papamichael I, 1259 Peacocke C, 44 I, 1482 Peters B, 585 Peitersson JBC, 1I29 PfaffD,713 Piskon J, 977,166 1 Plaksin GV,1509 Pogomtz M,85 1 Prins R, 73 Prins W, 452 Pudas M, 67 1 h&tolasR, 252
Macquat Y, 614 Maniatis K, 1 MarinN, 1388 Martinez JM, 929 Matsumura Y,237 MeierD, 1171, 1374, 13%,1405,1500, 1550, 1661 Meijer R, 799 Melaaen M Chr, 1046 Milosavljvic I, 1246 Minowa T, 396,12 19 Mironova N, I I71 Moilanen A, 122 Moonen RHW,452 Moreira JR, 843 MorfPh, 150 Moms M, 509 Moss HDT,1586 Mozaffarian M, 405 Murillo M B, 252 Munvanashyaka JN, 1564
Qvale B, 92 Radtke S, 1500 Ragland KW,8 12 Rauch R, 199 Raveendran K, 1025 Reed TB,693 Rieckmann Th, 1076 Risnes H, 32,6 1,730 Rodriguez I, 101I Rodriguez JJ, 929 Romey I, 499 Rbnnbiick M, 743 Rosh Ch, 499 Rossinskaja G, I500 Roy C, 1296,1349,1564 Rozendaal CM, 799 Ruijgrok WJA, 799 Ruiz E, 137 Rumpel S, 221
Naber J E, 1312 N m S-S, 358 Neeft JPA, 162, 1661 Nee@ M,524,536,549 Nieminen J, 549 Nussbaumer Th, 150,573,585,941 Nuutinen LH, 705,779
Saenz D,929 Saez F, 929 Saka S, 1338 Salvador ML,346 SalunannR, 941 Sanati M, 824 S b h e z JM, 137 Santom A, 1143 Sasaki €4,237 Sawayama S, 1219 Schenkel Y,1364, 1618,1633 Schmieder H, 109 Segerborg-Fick, A, 1661 Senehw K,630,758 Sharypov VI, 1388 Shimizu B, 1179 Shishko Y,1213 Simell PA, 162 Sipila K, 1468 SjWr6m K, 162,322,499,1129, 1661 Smeenk J. 379
Oasmaa A, 1468, 1661 Odenbrand I, 524,536 Ogi T, 956, 1219,1326 Oja V, 1226 Ollesch T,1405 Ollila HJ, 705 Onder CH, 573 Ostlie LD, 812 Ostman A, 867 Otero J, 137 Padban N, 524,536 Padinger R, 9 18 Pakdel H, 1564
1685
Smolders K, 766 Seerensen LH, 32 Solantausta Y, 831,867, 1661 Soldaini I, 1525 Someus GE, 165I Senju OK, 730 Serensen LH, 61, 122 Sotirchos S, 1586 Spince B, 1642 Stahl K, 221,524,536,549 Startsev AN, 1207 Stenseng M, 1061 Storm C, 1433 Strenziok R, 1268, 1452 Stromberg B, 1234 Struis RPWJ, 73 Stucki S, 73,388 Sturzenegger M, 388 Suomalainen M, 162 Suuberg EM, 1246
van Doom J, 265 van Swain WARM, 452 Venderbosch RH, 1268, 1586 Virtanen ME, 67 1,705 Visser HJM, 272 Vblker S, 1076 von Scala C, 73 Vourliotis P, 789 Vriesman P, 322 Vvedenskaya T, 1213
Tahara K,420 Tam MS, 1179 Taralas G, 176 Tardenaka A, 1642 Telysheva G, 1171, 1500 Thunman H, 743 Tiainen MS, 671,705,779 Toledo JM, 333,887 Tondi G,1525 Tranvik AE, 824 Turn SQ, 713
Yamaji K, 964 Yamamoto H, 964 Yang J, 1296 Yates JG, 188 Yazaki Y, 1326 Yin X, 465 Yokoyama S-Y, 420,956, 1661 Yoshida T, 237 Yjola J, 678 YuC, 1091, 1107
Wagenaar BM, 1268,1586 Waldheim L, 509 Walker M, 44 1 Wartmann J, 499 Weber J V, 1388 Wieer U, 896 Winter F, 64 1 Wistrom C, 379 Wu C, 465
Zandersons J, 1642 Zeevalkink JA, 13 12 Zethraeus B, 824 Zhang W, I09 1 Zhang W, 1 107 Zheng S, 465 Zhurinsh A, 1642 Zielke U, 162 Zintl F, 1234 Zwart R W R, 405 Zylbersztajn D, 843
Udas S, 1540 Uil H, 287 OnaI 6,473 Unterberger S, 656,896, 1433 Usenko A, 1213 van de Beld B, 298, 1312 Van den Bulck E, 1417 van der Aa BJ, 1268 van der Drift A, 265 van der Wal S, 1312
1686
Subject Index The index is based on keywords for each paper. The page numbers shown after each index entry refer to the first page of the paper that considers that topic Activated alumina, 15 17 Activated carbon, 1642, 1651 Activation, I509 Adhesives, I 197 Ageing, 1 171 Agglomeration, 122, 272,671, 705,779, 824 Agricultural residues, 1 179 Agricultural waste, 22 1 Agro-industrial residue, 209 Air staging, 94 1 Alkali, 358, 705, 713 Allothermal pyrolysis, 22 I Almond shell, 713 Alternative feedstocks, 1011 Alumina activated, 15 17 Ammonia, 322,524 Amorphous material, 779 Analysis, 137, 1374, 1564 gas, 162 principal components, 101 1 SEM,564 tar, 150 thermoeconomic, 843 Anhydrodigosaccharides, I034 Applications assessment, 85 I Aqueous phase processing, 1186 Arsenic, 1396 Ash, 122,272,824 chamcterisation, 564 deposition, 758, 789 interaction, 272 sintering, 122 Assessment, 867, 1213 applications, 85 1 life-cycte, 420 parameter, I2 13 Atomisation, 1459 Availability, 998
cost, 964 future, 964 potential, 22 1 supply, 964 utilisation technology, 964 Biofuels, 11 16, 1234 wet, 678 woody, 758 Bio-oil, 867, 1171, 1207, 1259, 1268, 1374, 1396, 1452, 1468, 1482, 1550
applications, 977, 1268 atomisation, 1459 characteristics, 977 combustion, 1452, 1459, 1586 emulsion, 1525 properties, 977 upgrading, 977 Black box, 998 Black liquor, 252 Boiler, 779, 1468 residential, 875 Brazil, 509 Burner settings, 1468 Cake filtration, 730 Calcium, 1586 Calorimeter, 614 Cane trash, 509 Carbon, 1509 activated, 1642, 1657 total organic, 1219 Carbon dioxide, 47, 6 1,420 emissions, 964 gasification, 346 reduction, 956 Carbon monoxide, 47,875,908 emissions, 573, 1459 Carbonisation, 1651, 1179, 1509, 1618, 1633 catalysts, 1642 low temperature, 1651 Cascade impactors, 929 Catalysis, 346, 1186 metal, 73 Catalyst, 396, 1207 carbonisation, I642 deactivation, 875
Bagasse, 509,83 1, 1349 Bark, 678, 1564 Beech wood, 1076, 1633 BIGCC, 488 Biocarbons, 1 179 Biocmde, 1312 Bio-emulsion, 1525 Bioenergy, 8 12
1687
8 12,824,875,896,9 18,956, 1468 bio-oil, 1586 chamber, 573 correlation, 630 fluidised bed, 272,671,705, 713,779,929 gas, 473,524 modelling, 585 packed bed, 585 particle, 908 pyrolysis liquid, 1452 temperature, 630 Combustor swirling, 599 Commercialisation,465, 1 197 Competitiveness,85 I Composition, 1564 flue gas, 630 waste, 1219 Concentrated radiation, 1034 Conceptual design, 388 Conditioning, hot gas, 365 Contaminants, chloro-organic, 1405 Control, 9 18 Cooking biomass, 693 COP3,964 Co-pyrolysis, 1349, 1388 Corn cob, 564 Corrosion, 109 Cost, 85 1 bioenergy, 867,964 Crushing strength, 1642 Cyclohexane, 176
iron, 1388 net-based, 875 nickel, 358 vanadium, 887 Catalytic deoxygenation, 15 17 gasification, 358 oxidation, 887 pyrolysis, 1500,1517 steam reforming, 1577 CCA, see copper-chromium-arsenic Cellulose, 1034, 1076, 1091, 1186 pyrolysis, 1500 Ceramic filter, 473 CFD, 365, 1281 Char, 32,47 characterisation, 1 116 gasification, 61,92 separation, 1281 Characterisation, 209 biomass, 209 char, 1116 Charcoal, 73, 1046, 1179, 1246, 1349, 1364,1603, 1633, 1642 fragmentation, 73 reactivity, 73 wood, 1396 Chemical conversion, 1338 Chemical kinetic modelling, 64 1 Chemical products, 1 186 Chemical quenching, 22 1 Chemical storage, 405 Chemicals, 977, 1197, 1338 Chips, 678,918 Chloride, 73 Chlorine, 887, 1234 Chloro-organiccontaminants, 1405 CHP, 307,499 decentralized, 499 Chromatography, 137 gel-permeation, 150 Chromium-copper-arsenic,1396, 1417 Circulating fluidisedbed, 199,209,265, 333,452,465,766,1259 FICFB, 199 Pyrolysis, 1259 Clean-up, hot gas, 1,379,473 CI-VOCS, 887 CO, see Carbon monoxide CO,, see Carbon dioxide Coal, 61 boiler, 1433 Co-combustion, 789,799, 1433 Co-fire, 758 Co-firing bio-oil, 1586 Co-generation, 509,843 Co-hydropyrolysis, 1388 Combined-cycle, 509 Combustion, 630,641,656,743,758,
Dangemus goods, 1482 Decentralised CHP,499 Decentralised gasification, 499 Defluidisation, 824 Degradation, 1564 Density, 1618 Depolymerization, 1 186 Deposits, 713 Derating, 441 Design, 379,977, 1281 conceptual, 388 statistical experimental, 10 1 1 system, 465 Developing countries, 693 Devolatilization, 1011 Diesel emulsion, 1525 fuel, 1540 Dilution tunnel, 614 Dioxins, 887,1405 Disposal, 1417 District heat, 867 Domestic waste, 1219 Downdraft, 743 gasification, 426
1688
inverted, 693 Drying, 307,678,812,1046 Dynamic modelling, 92 Dynamics, measurement, 573
FLOX burner, 1459 Flue gas, 176,678 composition, 630 Fluidised bed, 122, 188,322,346,
473,824,1091, 1197, 1374, 1396, 1577 circulating, 199,209,265,333, 452,465,766,1259 combustion, 272,671,705,713, 779,929 FICFB, 199 gasification, 272,713 gasifier, 536,599 pyrolysis, 1281 Fluidisation, 365 Forced convection, 693 Fouling, 713
Economics, 465,488,812,831,851,I197 Economy, 13 12 Emciency, 6I4 heat, 630 thermal, 1312 Electric power generation, 465 Electricity, 420,812,831,998 Electron spin resonance spectroscopy,
, 1171 Emissions, 614,641,656,758,789,896,
908,918,1452,1468,1586 CO,573, 1459
Fractional separation, 1564 Fuel bound nitrogen, 473 Fuel cell, 388, I 1 7 9 Fuel, 1338 blending, 758 diesel, 1540 liquid, 977 nitrogen, 322 quality, 743 renewable, 867 size, 743 staging, 941 waste, 524 Functional group, 1234 Fundamentals of pyrolysis, 1603 Fungicide, 1550 Furnaces, 908,9 18,941
(32,964
NOx, 941,1459 pollutant, 908 Emulsion, bieoil-diesel, 1525 Energy, I65I flows, 1618 plantation, 420 renewable, 488 supply system, 405 systems, global, 964 Engine, 307 gag 426,465 operation, 441 tests, 1525 Enthalpy balance, I 3 12 Entrained flow gasification, 221 EnvirOnment, 165I Environmental impact, 509 ESR see Electron spin resonance Evaluation, 730,867, 1025 Exergy, 843 Experimental verification, 1 I58 Extractives, 101 1,1143
Garbage, see Waste Gas, 1651 analysis, 162 chromatography, 137 clean up, 887 cleaning, 441 combustion, 473,524 engine, 426,465 infrastructure, 405 mixing, 452,573,1509 producer, 426 quality, 265 sampling, 162 transportation, 405 turbine, 473, 1452 velocity, 1213 Gasification, 1,47,73,109,122,137,
Fast pyrolysis, 977,1259,1396,1405 liquids, 1577 Fatty acid methyl esters, 1 5 17 Feedstock effects, 1 186 FICFB, 199 Filter ceramic, 473 granular, 379,730 high temperature, 730 moving bed granular, 379 particulate, 379 Filtration, 365 cake, 730 micro, 1171 Finance, 998 Fischer-Tropsch, 488 Fixed bed, 92,743,812,I158 gasification, 150
162,188,209,237,265,287, 307,322,333,365,379,396, 452,509,599,693,743,956, 1234,1433 black liquor, 252 catalytic, 358 char, 61,92
1689
co,, 346
Hydrogen, 32,388,396,405,1577 Hydrogenolysis, 1207 Hydrolysis, 1 186 Hydrolytic lignin, 1509 Hydro-pyrolysis, 1388 Hydrothermal, 396 conversion, 1312 Hydrotreatment, 1540
concept, 221 decentralised, 499 downdraft, 426 entrained flow, 221 fixed bed, 150 fluidised bed, 272,713 low temperature, 358 pressure, 473,524,536,549 small scale, 441 steam, 32, 199,346 two-stage, 92 Gasifier, 209,333,426,465 fluidised bed, 536,599 inverted downdraft, 693 reverse flow,298 stratified downdraft,426 Gas-particle partitioning, 713 Gate fee, 998
IGCC, 524 Ignition front, 743 Image furnace, 1034 Implementation, 851 Inductively coupled plasma, 1396 Industrial, 1468 waste, 1374 Inhibition, 32,47 chloride, 73 Integral systems, 488 Inverted downdraft gasifier, 693 Ironcatalyst, 1388 Iron oxide, 388
GC/MS, 929
Gel-permeationchromatography, 150 Global energy, 964 Global warming, 420 Glucose, 1338 Glycerin, 1577 GPC, see Gel-permeation chromatography Granular bed, 365 Granular filter, 379,730 moving bed, 379
Kieselguhr, 358 Kiln, rotary, 1651 Kinetic model, 1034,1417 Kinetic modelling, 1158 chemical, 641 Kinetic parameters, 61,1129 Kinetics, 32,47,92 150,1061,1076, 1 I29 pyrolysis, 1 1 16
Handling, 1482 Hard woods, 1143 Hazardous components, 1405 Hazardous materials, 1482 Hazards, 1482 Health and safety, 1482 Heat, 1296,1452 district, 867 efficiency, 630 flux densitis, 1034 production, 867 recovery, 678 release, slow, 614 transfer proprties, 1046 transfer, 678,1034 transport., 1076 Heating rate, 1618, 1633 Heavy metals, 1364, 1405 Hemicellulose, 1577 Herbaceous biomass, 221, 1011 High presure, 109 High temperature, 109 filter, 730 Hot gas clean-up, 1,379,473 Hot gas conditioning, 365 Hybrid-poplar, 8I2 Hydraulic resistance, 1213 Hydrogasification,405
Laboratory scale, 1374 Land use model, 964 Layer porosity, 1213 Leaching, 564 Levoglcosan, 1338, 1500 Levoglucosenone, 1500 Levulinic acid, 1 186 Life cycle assessment, 420 Lignin, 1076 hydrolytic, 1509 Lignite, 789 Lignocellulosic wastes, I 1I6 Liquefaction, I3 12, 1326 Liquid, hels, 977 products, 1388 Liquid, pyrolysis, see Bio-oil Liquidization, 1219 Low-temperature, 1417 carbonisation, 1651 gasification, 358 Macro mixing, 573 Macroparticle, I01 1 Mass balance, 1034, 1374 Mass flows. 16I 8
1690
Mathematical modelling, 92, 1158 Measurement dynamics, 573 Measurement in stack, 614 Measurements, 656 Mechanisms, 1034 Metals, 1417 catalysis, 73 heavy, 1364, I405 oxide, 388 release, 1417 Methanol, 420 Micro filtration, 1 171 Micropores, 1509 Mineral materials, 176 Miscanthus, 322 Model, compounds, 1234 kinetic, 1034, 1417 land use, 964 Modelling, 176, 188, 333,405,599, 641, 656,678,693,1046,1061, 1076, 1091, 1107, 1281,1296, 1618 chemical kinetic, 641 combustion, 585 dynamic, 92 kinetic, 1 158 mathematical, 92,307, 1158 NOx, 941 reactivity, 6 1 Moieties, paramagnetic, I171 Moisture, 1129 content, 1618 Moving bed granular filter, 379 Moving stoker, 789 Multivariate regression, 1076
Oxidation catalytic, 887 Oxidation, 109 catalytic, 887 Packaging, 1482 Packed bed combustion, 585 PAH, see Polycyclic aromatic hydrocarbons Paramagnetic moieties, 1 171 Parameter assessment, I2 I3 Particle, 162, 1046, 1 129 combustion, 908 size, 896 size distribution, 929 wood, 1046 single, 1046 thick wood, 1143 Particulate filters, 379 Particulates, 365,441 Partitioning, 7 13 gasparticle, 713 Pellet wood, 867 Percolative disintegration, 73 Petroleum, 1349 acid sludge, 1642 Phenolics, 1338, 1561 Phenols, 1 197 Phosphoric acid, 1500 Physical properties, 1246 Pilot plant, 209,452 Pine, 1326 Plantation, 956 energy, 420 Plasma, 1396 Plywood, 779 Pollutant emissions, 908 Polycyclic aromatic hydrocarbons, 929 Potassium, 122, 1107 Power, 812, 1452 generation, 420 generation, electric, 465 output, 6 14 plant, 1433 Preservative, wood, 1550 Pressure gasification ,473,524,536, 549 Pretreatment, 101 1, 1433 biomass, 564 digestion, 1219 Principal components analysis, 10 1 1 Producer gas, 426 Product slate, 10 1 1 Project finance, 998 Protocol, 162 Pulping, 1186 straw, 252 Pyrolysis, 176,322,346,101 1, 1025, 1046, 1061, 1076, 1091, 1107,
Nanoparticle, 896 Net-based catalyst, 875 Network, 599 reactor, 333 N-heptane, 176 Nickel catalyst, 358 Nitrogen, 524,641 conversion, 94 1 fuel bound, 473 oxides, 9 18 NOx, 524,641, 1586 emissions, 941, 1459 modelling, 941 reduction, 918, 941, 1413 Oil, 1349, 1564 pyrolysis, see Bio-oil quality, 1468 tall, 1540 Olive+il residue, 564 Open core gasification, 426 Optimisation, 656, 9 18 Organic acid, 12 19
1691
1116, 1129, 1143, 1158, 1179, 1197,1234, 1364, 13%, 1417, 1433,1564, 1651 allothennal, 221 black liquor, 252 catalytic, 1500,1517 cellulose, I500 circulating fluid bed, 1259 dynamics, 1143 fast, 977, 1034, 1259, 1374, 1396, 1405 fluid bed, 1281 fundamentals of, 1603 hydro-, 1388 kinetics, 1116 liquid applications, 1268 liquid atomization, 1459 liquid combustion, 1452, 1459 liquid, 867, 1207, 1268, 1481, 1550, 1577 modelling, 1281, 1296 oil, see Bio-oil product yields, 1 143 rotating cone, 1268 tars, 1226 temperature, 1143 time, 1 143 vacuum, 1296,1349, 1564 Py~olyhc,1349 carbon deposition, 1633
oliveail, 564 wood, 1179 Retort, 1618 Reverse flow gasifier, 298 Review, 1,977 Rice straw, 358 Riser reactor, 188 Risk, 998 Rosin acid, 1540 Rotary kiln, 1651 Rotating cone pyrolysis, 1268 Sampling, 137 gas, 162 Sawdust, 346,678,779,12 13, I642 Scale-up, 465,998, 1259, 1296 Scaling relationships, 188 Scanning electron microscope analysis, 564 EDS, 671,705,779 Secondary tar reactions, I50 Selection, 1025 Separation, 1186, 1509 char, 1281 Sewage sludge, 1433 Shape, 16I8 Single particle, 1046 Sintering test, 564 Size distribution, 824 Slagging, 824 Slow heat release, 614 Small scale firing systems, 656 Small-scale gasification, 441 SNG production, 405 Soda black liquor, 252 Softwoods, 1 143 Sorbents, 1509 SOX,see Sulphur dioxide Specification, 1468 Spills, 1482 Spruce, 6 1 Standards, 614 Statistical experimental design, 101I steam, gasification ,32, 199, 346 reforming, 1577 Stirling engine, 1459 Stoker moving, 789 Storage, chemical, 405 Stove, 693 tiled, 6 I4 woodgas,693 Stratified downdraft gasifier, 426 Straw, 122,221,524 pulping, 252 rice, 358 washed wheat, I06 1 wheat, 564,1061 Sugar cane harvesting, 509
Quasitemary diagrams, 67 I Quenching, chemical, 221 Radiation, concentrated, 1034 Reaction enthalpy, 13 I2 potential, 176 Reactions, secondary tar, 150 Reactivity, 32,47,73,92 charcoal, 73 modelling, 61 profile, 61 Reactor, 977, 1296 network, 333 riser, 188 Reburning, 1433 Regression, multivariate, 1076 Release of metal, 14I7 Renewable energy, 488 Renewable fuel, 867 Research needs, 1197 Residence time, 1633 distribution, 573 Residential boilers, 875 Residue, 1349 agricultural, 1179 agro-industrial, 209
1692
Sugar, 1219 Sulphonic acids, 1642 Sulphur dioxide, 1586, 1642 Sunflower husks, 1213 Supercritical water, 109,237, 1338 Swirl flame, 896 Swirling combustor, 599 Synthetic polymers, 1388 System, 307 design, 465
Uncertainty, 998 Upgrading, 1207 bio-oil, 977 Vacuum pyrolysis, 1296, 1349, 1564 Vanadium catalyst, 887 Vapour pressure, 1226 Veneer shorts, 1642 VHC,908 Volatiles, 641, 1651 Volatility of tar, 1226 Vortex, 896
Tall oil, I540 Tar, I, 137, 162,265,333,441,536 analysis, 150 conversion, I50 pyrolysis, 1226 reactions, secondmy, 150 vapour, 176 volatility, 1226 Temperature, 1364, I633 combustion, 630 high, 109 pyrolysis, 1143 Test method, 614 TGA, see Thermogravimetry Thermal desorption, I65 1 Thermal efficiency, 1312 Thermochemical, 176 Thermoeconomic analysis, 843 Thermogravimetry, 32,47,6 1, 1061, I076 TGA-DTA, 564 Thick wood particle, 1143 Tiled stove, 614 Total organic carbon, I2 19 Trace elements, 824 Transfer, 1296 heat, 678, 1034 Transport, 956, 1482 gas, 405 fuels, 488 heat, 1076 Trees, 8 12 Tri-generation, 488 Turnkey projects, 998 Two-stage gasification, 92
Washed wheat straw, 1061 Waste, 887, I2 I9 agricultural, 22 1 biomass, 1374 composition, 12I9 domestic, 1219 fuels, 524 glycerin, 1577 industrial, I374 lignocellulose, I I16 to energy, 799 wood, 465,789, 1364,1396, 1405, 1417 Water consumption, 199 Water processing, 1326 Water vapour sorption, 1550 Wet biofiel, 678 Wheat straw, 564, I06 I washed, 1061 Willow, 524 Wood charcoal, 1396 chips, 678,918 gas stove, 693 particle, 1046 pellet, 867 preservative, 1550 residues, 1179 waste, 465,789,1364, 1396, 1405, 1417 Woody biofuels, 630,758, 1388
XRD, 779 Xylan, 1076
UN classification, I482
1693
