Studies in Surface Science and Catalysis 127 HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS
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Studies
in S u r f a c e
A d v i s o r y Editors:
Science and Catalysis
B. Delmon and J.T. Yates
Vol. 127
HYDROTREATMENT AND HYDROCRACKING OF OIL FRACTIONS Proceedings ofthe 2nd International Symposium/7th European Workshop, Antwerpen, Belgium, November 14-17, 1999
Edited by B. D e l m o n
Universite Catholique de Louvain, Unite Catalyse et Chimie de Mat#riaux Divis#s, Louvain-La-Neuve, Belgium
G.E Froment
TexasA&M University, Department of Chemical Engineering, College Station, Texas, USA
P. G r a n g e
Universite Catholique de Louvain, Unite Catalyse et Chimie de Mat6riaux Divis#s, Louvain-La-Neuve, Belgium
1999 ELSEVIER Amsterdam m Lausanne n N e w York n Oxford ~ Shannon m Singapore m Tokyo
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The Netherlands
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Preface
xl
Keynote lectures Hydrocracking in the year 2000 : a strong interaction between technology development and market requirements J.K. Minderhoud, J.A.R. van Veen, A.P. Hagan Life cycle of hydroprocessing catalysts and total catalyst management S. Eijsbouts
21
Hydrogen spillover and hydrocracking, hydroisomerization K. Fujimoto
37
Conversion of model sulfur compounds to characterize hydrodesulfurization CoMo/AI20 catalysts J. Leglise, L. Finot, J.N.M. van Gestel, J.C. Duchet
51
The Catalyst under Working Conditions Highly active MoS2-based dispersed catalyst with a novel morphology Y. Araki, Y. Iwata, Y. Miki, K. Honna, N. Matsubayashi, H. Shimada
69
Use of noble metals in hydrodeoxygenation reactions A. Centeno, R. Maggi, B. Delmon
77
Influence of the hydrogen sulfide partial pressure on the hydrodeoxygenation reactions over sulfided CoMo/carbon catalysts M. Ferrari, S. Bosmans, R. Maggi, B. Delmon, P. Grange
85
Performance of noble metaI-Mo/7-AI203 catalysts: effect of preparation parameters M.H. Pinz6n, L. Merino, A. Centeno, S.A. Giraido
97
Use of ammonium tetrathiomolybdate as a new precursor for the preparation of hydrodesulfurization catalysts by a sol-gel method L. Le Bihan, C. Mauchauss6, E. Payen, J. Grimblot
105
Influence of sulphidation and fluoridation on the HDN of o-toluidine over tungsten catalysts ex ammonium tetrathiotungstate M. Sun, R. Prins
113
Modeling molybdenum carbide-based hydrodesulfurization (HDS) catalysts using carbon-modified Mo(110) surfaces C.L. Roe, K.H. Schulz
121
TiO2-coated on AI203 support prepared by CVD method for HDS catalysts K. Segawa, S. Satoh
129
Sulfur uptake, exchange and HDS activity of NiMoO#AI203 catalysts T. Koltai, M. Dobrovolszky, P. Tdtdnyi
137
Stability of CoMo/AI203 catalysts : effect of HDO cycles on HDS T.-R. Viljava, S. Komulainen, T. Selvam, A.O.I. Krause
145
CoMo/AI203 and CoMo/TiO2-AI20~ catalysts in hydrodesulfurization : relationship between the promoting effect of cobalt and the nature of the support M. Vrinat, D. Letourneur, R. Bacaud, V. Harl@, B. Jouguet, C. Leclercq
153
Effect of chelating agents on HDS and aromatic hydrogenation over CoMoand NiW/AI203 Y. Ohta, T. Shimizu, T. Honma, M. Yamada
161
Genesis, structural and catalytic properties of Ni-Mo-P-alumina based hydrotreating catalysts prepared by a sol-gel method R. Iwamoto, J. Grimblot
169
Industrial Process Aspects Hydroprocessing kinetics for oil fractions T.C. Ho
179
Molecular base approaches by GC-AED to HDS of gas oil on sulfide catalysts I. Mochida, S. Shin, K. Sakanishi, D. Grudoski, J. Shinn
187
The nitrided CoMo catalysts for hydrodesulfurization and hydrodenitrogenation M. Nagai, H. Koyama, S. Sakamoto, S. Omi
195
HDS of dibenzothiophene and vanadyl porphyrin HDP on bulk Fe-Mo mixed sulphides M.A. Luis, A. Rives, R. Hubaut, B.P. Embaid, F. Gonzalez-Jimenez, C.E. Scott
203
Design of a model activity test for second stage deep HDS catalysts H.R. Reinhoudt, M. van Gorsel, A.D. van Langeveld, S.T. Sie, J.A. Moulijn, J.A.R. van Veen
211
The influence of zeolite introduction on the HDS activity of CoMo catalysts L. Zanibelli, D. Berti, M. Ferrari, C. Flego, R. Riva
219
Hydrodenitrogenation properties of supported metal catalysts in the presence of H2S E. Peeters, C. Geantet, M. Vrinat, J.L. Zotin, M. Breysse
227
vii
Hydrodesulphurisation and aromatics hydrogenation on straight run gas oils of maya crude oil A.A. P6rez, S.G. Marroquin, R.G. Betancourt, T.A. Moreno, R.E. Aguilar
235
Hydrocracking of vacuum gas oil on CoMo/alumina (or silica-alumina) containing zeolite W.-S. Choi, K.-H. Lee, K. Choi, B.-H. Ha
243
Testing and characterisation of Pt/ASA and PtPd/ASA for deep HDS reactions H.R. Reinhoudt, R. Troost, A.D. van Langeveld, S.T. Sie, J.A. Moulijn, J.A.R. van Veen
251
Characterization of Catalysts Probing the electronic state of nickel-molybdenum sulphide catalysts using ortho-xylene hydrogenation L. Fischer, V. Harl6, S. Kasztelan
261
IR study of hydrotreating catalysts in working conditions : comparison of the acidity present on the sulfided phase and on the alumina support A. Travert, F. Maug6
269
Physicochemical characterization of VGO MHCK catalysts and its extrapolation to catalytic activity M.M. Ram#ez de Agudelo, E. Plujic, J.A. Salazar
279
Reactor Modeling Modeling a hydroconversion reactor based on a computational fluid dynamics approach M.M. Carbonell, R. Guirardeilo
289
An integrated approach for hydrocracker modeling C.S.L. Narasimhan, M. Sau, R.P. Verma
297
Fundamentals and Reaction Mechanisms Ab-initio energy profiles for thiophene HDS on the MoS2 (1010) edge-surface P. Raybaud, H. Toulhoat, J. Hafner, G. Kresse
309
Diffusion effects and direct C-N cleavage in the HDN of o-toluidine and methylcyclohexylamine over sulphided NiMo/7-AI203 and Mo(P)/7-AI20~ catalysts F. Rota, R. Prins
319
Theoretical study of benzothiophene hydrodesulfurization on MoS2 S. Cristol, J.F. Paul, E. Payen, D. Bougeard, J. Hafner, F. Hutschka
327
t
Vlll
Posters Effects of alumina-titania supports on the activity of NiMo catalysts J.R. Grzechowiak, I. Wereszczako-Zielfnska, J. Rynkowski
337
Effect of light cycle oil on diesel hydrotreatment J. Ancheyta-Ju~rez, E. Aguilar-Roddguez, D. Salazar-Sotelo, G. Marroqufn-S~nchez
343
Effects of hydrogen sulphide on the hydrodesulphurization of an industrial HDS feedstock in a fixed-bed pilot plant J. Ancheyta-Ju~rez, E. Aguilar-Rodrfguez, D. Salazar-Sotelo, G. Betancourt-Rivera, G. Quiroz-Sosa
347
Catalytic properties of WS2 catalysts prepared by in situ decomposition of tetraalkyl-ammonium thiotungstates G. Alonso, V. Petranovskii, M. Del Valle, J. Cruz-Reyes, S. Fuentes
351
Synthesis, characterization and HDS activity of CoMo/AI203 catalysts prepared by two ways (impregnation of a sol-gel alumina and complete sol-gel synthesis) F. Dumeignil, J. Grimblot
357
On the use of Pco(Ni)Moll heteropolyanions for the preparation of alumina supported H DS catalysts A. Griboval, P. Blanchard, E. Payen, M. Fournier, J.L. Dubois, J.R. Bernard
361
Hydrotreating with mixed Fe-Ni sulphides P. Betancourt, C.E. Scott, J. Goldwasser, F. Gonzalez-Jimen~z, P.B. Embaid, R. Hubaut, A. Rives
365
TPR and NO adsorption studies of Mo, CoMo and NiMo catalysts supported on AI203-TiO2 mixed oxides L. Ceden, J. Ramirez, A .Ldpez-Agudo, M. Vrinat, R. Ldpez Cordero
369
Preparation and characterization of HNaY-alumina supports and their impregnated Mo catalysts T. Klimova, D. Soils, J. Ramffez, A. Ldpez Agudo
373
Modeling of nature and strength of acid centres in ultrastable zeolites as a component of hydrocracking catalysts A. V. Abramova, Ye. V. Slivinsky, Y.Y. Goldfarb, L. Ye. Kitaev, A.A. Kubasov
377
Hydrogenation heavy oil residues under 6 MPa pressure in motor fuels and feedstock for catalytic cracking A.S. Maloletnev, U.P. Suvorov
381
Influence of the nature of the metal and of the acidity of the support on hydrocracking reactions J.-A. Porta, J. Despr~s, F. Garin
385
Hydrogenation of AH-VR using carbon-supported catalysts A. Segawa, K. Watanabe, Y. Shibata, T. Yoneda
389
ix Effects of gaseous and liquid components on rate of deep desulfurization of heavy atmospheric gas oil M.V. Landau, L. Vradman, M. Herskowitz, D. Yitzhaki Catalytic functionalities of TiO2 based SiO2, AI203, ZrO2 mixed oxide hydroprocessing catalysts M.S. Rana, B.N. Srinivas, S.K. Maity, G. M. Dhar, T.S.R. Prasada Rao
393
397
Hydrodesulfurization of dibenzothiophene over Ni-Mo/(P)Ti-HMS catalysts T. Halachev, J.A. de los Reyes, C. Araujo, G. Cordoba, L. Dimitrov
401
The preparation of hydrocracking catalysts using mesoporous aluminosilicates of the MCM-41 - influence of the preparation conditions on the catalytic behaviour A. Klemt, A. Taouli, W. Reschetilowski, H. Koch
405
Selective hydrodesulfurization technology of cracked gasoline for gasoline pool in 2005 M. Li, H. Nie, Y. Shi, D. Li
409
Hydrotreating catalysts on alumina, titania or zirconia from ethanol/water solutions of heteropolyacids L. Pizzio, P. V~zquez, C. C~ceres, M. Blanco
413
A XANES temperature-programmed sulphidation study of modified NiMo/SiO2
421
HY zeolite-based catalysts for hydrocracking heavy oils K. Honna, Y. Araki, Y. Miki, H. Shimada, K. Sato, N. Matsubayashi
427
Hydrodesulphurization of residue-oil over Ni-Mo/HY-zeolite catalyst S. Bhatia, J.K. Heng, A.R. Mohamed
431
Authors i n d e x
435
hydrotreating catalysts R. Cattaneo, T. Shido, R. Prins
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The 2nd International Symposium on Hydrotreatment and Hydrocracking of Oil Fractions, which is also the 7th in the series of European Workshops on Hydrotreatment, took place in Antwerpen, Belgium from November 14 to 17. The Symposium emphasized how oil refining faces increasingly severe environmental regulations. These and the increasing application of heavier crudes containing more S-, N- and metal components call for more efficient hydrotreatment and hydrocracking processes. It is clear from the keynote lectures, the oral contributions and the posters of this meeting that adapting the operating conditions will not suffice. Adequate catalysts need to be developed, with different composition and structure. Surface science techniques and molecular modeling are now well established tools for such a development. They should be of help in widely different aspects, like the role of precursors in the preparation or the modifications undergone by the catalyst under reaction conditions. The improvement of hydrotreatment and hydrocracking also needs accurate modeling of the chemical reactor. This requires more representative hydrodynamics and kinetic models whose validity extend to the very low S-and N-contents. These areas should be vigorously developed. We look back at a successful symposium with contributions from all over the world, reflecting the state of the art in industrial practice, in industrial research centers and in academia. Let these Proceedings dissiminate the information presented at the Symposium also to those who were not able to attend. B.Delmon, Universit~ Catholique de Louvain,Belgium G.F.Froment,Texas A & M University, USA P.Grange ,Universit~ Catholique de Louvain,Belgium
The 2nd International Symposium "Hydrotreatment and Hydrocracking of Oil Fractions" was organized by : The Technological Institute associated with the Royal Flemish Society of Engineers (TI - K VIV). The Technological Institute was founded in 1940 with the aim of disseminating information on scientific and technological development by means of seminars, lectures, courses and conferences. Address :Technological Institute vzw Desguinlei 214, B- 2018 Antwerpen tel: + 3 2 3 2 1 6 0 9 9 6 fax: + 3 2 3 2 1 6 0 6 8 9 e-mai I :
[email protected]
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KEYNOTE LECTURES
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
Hydrocracking in the Year 2000: A strong interaction Technology Development and Market Requirements
between
(J.K. Minderhoud, J.A.R. van Veen and A.P. Hagan)
Summary New developments in hydrocracking technology are increasingly guided by the prevailing market conditions and requirements. This necessitates a very good knowledge of local circumstances in the refinery as well as all integrated approach of the various disciplines involved such as catalyst, process, reactor and engineering technology. Some recent developments, illustrated by a case study, are discussed.
Introduction Starting in the early 1960's, hydrocracking has become one of the major conversion processes in the refinery. It usually converts a rather heavy, low quality feedstock into lighter, highly valuable transportation fuels, contributing significantly to the overall profitability of the refinery. Due to the nature of the process, hydrocracking is predominantly suited to producing middle distillates with excellent product qualities. Jet and diesel fractions can be obtained with very low sulphur contents (often below 20 ppm) and very good combustion properties (kerosine smoke points above 25 rain and diesel cetane numbers above 55). The obvious reason for it is the relatively high hydrogen pressure used, typically above 100 bar, which results in high removal rates of hetero-atoms (sulphur, nitrogen) contained in the feedstock and deep saturation of aromatic compounds. Other typical characteristics of hydrocracking are its flexibility in varying product slate, depending to a large extent on the type of catalysts used, and its potential to produce very good quality feedstock for lube base oil manufacturing, ethylene crackers and fluid catalytic crackers. Full conversion hydrocrackers usually contain at least two catalysts: a pretreatment and a cracking catalyst. Both catalysts are bifunctional: they contain a hydrogenation function and an acidic function (see Table 1). Pretreatlnent catalysts combine a strong hydrogenation function with a rather weak acidic function, whereas for cracking catalysts both functions should be well balanced. By the end of this century, the world hydrocracking capacity will be about 200 million tons per annum. Projections were that more than 225 million tons per annum of processing capacity would be available by the year 2002, with the highest growth rate expected to be in the Asia-Pacific zone, but the recent economic downturns have resulted in a slower increase in capacity. Currently, it is more realistic to expect a 3-5% growth in the next five years. The main driving forces for the expansion of the hydrocracking process are : 9 a steady, but continuing growth in middle distillates consumption in all parts of the world. 9 introduction of ever more stringent automotive fuel specifications (sulphur, aromatics, 95%vol recovery temperature) as prescribed in the USA and Europe. 9 processing synergy in combining catalytic cracking with hydrocracking, two large scale processes 9 increasing development of mild hydrocracking technology [1,2]
Technology Development Over the past forty years, a large number of different process and reactor configurations have been developed to carry out the hydrocracking process (Figure 1) [3-6]. Key differences have been: 9 two stage versus single stage / series flow operation 9 once through versus recycle mode of operation 9 a common versus a separate recycle gas system for first and second stage
9 world-scale single reactors containing multiple catalysts versus two or more reactors containing individual catalysts To comply with the demand for new hydrocracking capacity, it is evident that building new hydrocrackers will continue in the years to come, despite the likely incremental increase of capacity in existing hydrocracking units. In an economic climate of low refining margins and emphasis on high returns on investment, there is a very strong incentive to design and construct hydrocrackers with minimum capital investment. It often means "silnple" units, i.e. single reactors, operating in once-through mode and at low pressures.This is the major reason that the mild hydrocracking process has received increased attention in the last decade. On the other hand, the call for ultra low sulphur and, in particular, very low aromatics levels in the products cannot easily be satisfied by applying a low hydrogen partial pressure process. Moreover, catalyst activities in mild hydrocracking are reduced as well, leading to higher catalyst volumes, i.e. bigger reactors, to achieve the same feed conversion level. These effects are illustrated in Table 2. In selecting the appropriate hydrocracking process, refiners will strive for the most cost-efficient way to achieve their objectives. For hydrocracking process licensors this implies that they need to strike the right balance between expensive, complex, high pressure processes offering much flexibility, products of superior quality, and cheap, simple, low pressure designs with more restricted deliverables. An example of this is the single-stage, single-reactor, stacked bed line-up with optional liquid recycle (Figure 2). With advanced heat integration, a four separator reactor effluent system and less fractionation equipment, this design offers clear advantages over the oldest, conventional two stage processes. In view of strong pressures to increase refinery margins, an increasing interest in optimising and revamping existing hydrocrackers has been shown in the last few years and is expected to continue in the next decade. There are numerous topics that can be addressed to achieve this goal (Table 3). In this article, we will highlight (i) the use of new and improved catalysts, (ii) proper arrangement of different catalysts in stacked bed configuration and (iii) new developments in reactor internals and fouling abatement. Finally, to emphasise the importance of close co-operation between refiner and process developer, we will present a case study in which various process aspects played a very important role in revamp and catalyst selection of the hydrocracker.
New and improved hydrocracking catalysts In present day industrial research, effectiveness and efficiency are key, and catalyst research and development is no exception to this. Target setting and the way R&D programmes are executed have received much attention in recent years. It is crucial that the targets for a new catalyst, in terms of improved performance and/or reduced costs, are valued by the final user. To achieve this, targets are increasingly defined in close co-
operation with refiners in order to include their specific requirements. For locations where a long term relationship has been developed, there is even a trend of tailor made catalysts. As a result, catalyst companies and process licensors continue to introduce new hydrocracking catalysts to the market (7-11). Here, we will discuss some new catalysts that were recently developed by Criterion and ZI, taking into account refiners' wishes.
DN-190: A high activity pretreatment catalyst A considerable number of hydrocrackers are severely constrained by an inability to meet the required nitrogen slip to the cracking catalyst, often aggravated by the processing of heavier feedstocks or feedstocks containing nitrogen species difficult to hydrodenitrogenate. Without better pretreat catalyst, these hydrocrackers suffer from shorter cycle lengths. To develop new pretreatment catalysts three widely different development concepts have been in use: (i) high dispersion, (ii) controlled acidity, and (iii) optimised pore structure. The overall concept is depicted in Figure 3. DN-190 was developed on the basis of a high dispersion of the catalytically active phase. The concept behind DN-190 is to boost the HDN activity by increasing the number of active sites by enhancing the catalyst surface area on a reactor volume basis. This volumetric catalyst surface area can be favoured by: 9 enhancing the CBD by optimising the shape and reducing the size of the extrudates. The latter favours mass transport, too. To avoid excessive pressure drop over the reactor, there is a minimum size and shape to ensure that the void fraction is at least 40%. 9 adjusting
the
SA = F * PV / M P o D
textural
properties
of
the
carrier,
according
to:
(1)
where F is a numerical constant which depends on the shape of the pores (4 in the case of cylindrical pores), PV the pore volume, and MPoD the median pore diameter. DN-190 is based on Century TM technology, a process in which a nano-crystalline phase of alumina is synthesised in-situ on the gamlna-alumina support to generate slit-shaped pores. The nanocrystalline phase suppresses stacking of molybdenum sulphide in the working catalyst. The existence of single layers leads to a very high dispersion of the active phase [12]. The result is a catalyst displaying a very high volumetric surface area. Table 4 clearly confirms that, as a consequence, the RVA of DN-190 is substantially higher than that of a conventional pretreat catalyst such as C-424.
Z-623." A high active, high middle distillate selective zeolitic catalyst It is well known that for cracking catalysts there is a kind of trade-off between high activity, high naphtha selectivity and low activity, high middle distillate selectivity [7]. The objective of new cracking catalysts is often to improve on activity whilst maintaining selectivity or vice versa. This can be achieved by altering the acidic function and/or the hydrogenation function. Zeolite modifications for instance are numerous: dealulnination, insertion of silica, re-insertion of alumina, morphology changes etc [13]. Despite the fact that an overwhelming number of treatments and modification routes have been explored already, there still appears to be room for new and successful zeolite types. This led, as an example, to the development of a dealuminated Y zeolite, which in combination with an amorphous silica-alumina (ASA) and a hydrogenation function, finally resulted in the highly middle distillate selective catalyst Z-603 [14]. Recently, by careful modification of both functions, Z-623 was developed as a result of efforts, stimulated by the market, to obtain a catalyst with higher activity than Z-603 without
comprolnising on selectivity. Figure 4 demonstrates that the Z-623 performance is indeed better than could normally be expected.
503.'A high diesel selective single stage catalyst Catalysts using amorphous silica-alumina as acidic function are very well suited to maximising diesel production. Furthermore, they have the ability for use in pretreating as well, because of the relatively modest acidity. Consequently, this type of catalysts can be employed as the single catalyst in a hydrocracker unit. Improvements in ASA catalysts can be obtained by variations in e.g., composition, synthesis conditions of the ASA, metal emplacement methods, post-treatments. Figure 5 shows how the efficiency of the NiW hydrogenation function can heavily depend on the metal emplacelnent route selected, which ultimately dictates the effectivity of the ASA catalyst in performance terms. Increasing acidity of the ASA is in the first instance beneficial, but finally the performance becomes less attractive due to an imbalance between acidity and maximum achievable hydrogenation power. In the development of 503, both the acidity of the ASA and the hydrogenation activity of the NiW function were improved. The results are shown in Table 5, where the new 503 catalyst is compared with the previous generation DW 800 catalyst: clearly, activity, selectivity and product properties have been improved. The opportunity was also taken to manufacture a Ni/Mo analogue on this new ASA support, designated 505. Due to its better HDN activity, 505 is best used for MHC duty where any ilnprovement in HDN is beneficial in enhancing the cracking conversion. For high pressure units, to make high quality middle distillates, 503 will be the higher performance catalyst, though.
Catalyst stacked bed arrangements To maximise the overall capacity and conversion capability of the hydrocracking catalyst system or to minimise the overall catalyst volume, it is of paramount importance to optilnise the ratio of pretreat over cracking catalyst. A crucial factor in determining the optimum ratio is knowing the hydrocracking activity (reaction rate constant for cracking) of the cracking catalyst as a function of pretreatment severity. Since cracking catalysts are acidic and variations in pretreatment severity result in effluents containing different amounts of organic nitrogen compounds, adsorption of basic nitrogen species can have a dominant effect on the apparent cracking rate constant [15]. This is further illustrated by laboratory tests in which feedstocks with nitrogen contents varying from 20 to 280 ppmwt were processed over a NiW/Dealuminated Zeolite catalyst. Figure 6 shows the following: * there is a significant reduction of the apparent reaction rate constant for cracking with increasing nitrogen content of the feed ,
the penalty of increasing nitrogen content is (slowly) decreasing with increasing operating temperatures
These effects are best understood by inhibition effects of adsorbed nitrogen which are described by Langmuir-Hinshelwood rate expressions. Apart from suppressing the cracking reactions, the nitrogen compounds also cause self-inhibition of the hydrodenitrogenation (HDN) reaction, as shown in Figure 7. This effect in itself further retards the cracking reactions. It is also well known that ammonia, organic sulphur compounds, hydrogen sulphide and (poly) aromatics display inhibitive effects, but their adsorption constants are at least an order of magnitude lower than those of organic nitrogen compounds [16]. The practical implications of the poisoning effects of nitrogen species on cracking catalysts are very dependent on the operating conditions of the hydrocracker. Feedstock type, hydrogen partial
pressure and temperature window between start-of-run and end-of-run of the hydrocracker cycle largely dictate what nitrogen content is acceptable in the feed entering the cracking catalyst. In two stage units, due to the absence or very low levels of ammonia, operating temperatures in the second stage are usually well below 400~ resulting in rather high nitrogen sensitivities. In single stage (series flow) hydrocrackers, however, operating temperatures are generally (much) higher and, hence, the cracking catalysts in those units can tolerate higher nitrogen slips from the pretreating stage. For the same reason, in two stage units, it is useful to consider applying stacked beds in the first stage where temperatures are often higher than in the second stage. This may lead to higher cracking conversion levels in the first stage, off-loading the duty in the second stage, as illustrated in Table 6. Overall, this will result in longer cycle run lengths. Optimising the balance between the various catalysts in a hydrocracker calls for a good description of all kinetic parameters. Process models are therefore an indispensable tool in improving the operation of existing hydrocrackers and designing new ones. The type of models that have been developed, vary from very rigorous and fundamental to simple and correlative [17-22]. Commercial models try to combine simplicity and user-friendliness with accuracy and thoroughness.
Reactor internals
To achieve maximum utilisation of the catalyst inventory in hydrocrackers, it is essential to obtain both an even radial distribution of liquid and gas across the catalyst beds and excellent interbed quench performance [23]. Recently, various studies have been undertaken to better understand and improve the performance of liquid distribution trays [24]. Patel et. al. described the development of the so-called Vapor-Lift Distribution Tray, which is claimed to have a much more stable, low tilt-sensitivity operation over a wide range of vapour/liquid ratios than classic bubble cap trays. The Shell developed, so-called High Dispersion (HD) trays were found to display a much better liquid distribution uniformity than more conventional trays, as shown in Figure 8 [25]. Main reason for the better performance of the HD tray is that contrary to conventional trays, gas and liquid are passing together through the nozzle, which causes an acceleration of the liquid and an intimate mixing of liquid and gas. Because of the excellent distribution at the top of the bed there is no need for a layer of distributive packing above the catalyst bed, resulting in increased catalyst volume per reactor. Due to the exothernlicity of the hydrocracking reactions, it is necessary to apply interbed quenching in order to achieve a safe and controlled operation and to optimise the axial reactor temperature profile. Interbed internals, used to reach this goal, allow injection of a cold gas or liquid medium and also need to provide adequate mixing of reactant liquids and gases toensure a homogeneous, radial temperature profile at the top of the catalyst bed. Many different interbed quench devices have been developed. One of the important aspects in new designs, apart from providing good mixing, is the aim of minimising the height of the internal to maximise the amount of catalyst to be loaded in the reactor. The Ultra Flat Quench (UFQ) internal with a height of only 1 meter, described by Ouwerkerk et. al., elegantly meets this requirement (Figure 9) [25]. In cold-flow testing with an imposed 30~ temperature difference above the internal, they reported a maximum radial delta T of only some 4~ below the internal in case of the UFQ instead of 16~ for a conventional internal (Figure 10). The high performance of the UFQ has been well demonstrated in several commercial operations.
Fouling abatement An important aspect to extend the run length of a commercial hydrocracker is to prevent reactor fouling since it will lead to increase of pressure drop and finally to a premature shut down. Fouling is often caused by inert solid particles entrained in the feed that deposit on and between the catalyst particles (salts, iron scale) or chemical substances that react and deposit on and between the catalyst particles. One of the elements to combat fouling is to apply graded layer loading of inert materials and catalysts which is based on the concept of deep bed filtration [26]. It entails loading higher voidage, larger particle size materials in the top layer, followed by layers of gradually smaller sized materials and finally the hydrocracking catalyst(s). A potential disadvantage of using bed grading is the loss of reactor volume for loading the actual hydrocracking catalyst since (a part of) the grading material can be inert (Raschig rings are often used). To mitigate these effects, materials displaying some catalytic activity are now being used. It should be realised, however, that, due to the larger particle size of the grading materials, diffusion limitational effects are more pronounced, resulting in lower effective reaction rate constants. Clearly, the key issue in optimising bed grading is finding the right balance between fouling prevention and preserving sufficient overall catalyst activity.
Case Study A further illustration of the many factors that count in improving hydrocracker operations is an example of an existing partial conversion hydrocracker for which plans were made to increase feed throughput by some 10%, to include some 10% more aromatic, higher nitrogen feed in the total feed diet and to abandon using a halide agent as activity booster for the catalysts. Moreover, the refinery had stated a number of premises that needed to be met: maintaining the same runlength as before, limiting the extra amount of naphtha and lighter products to maximum 20% more, and producing unconverted oil (called hydrowax) with preferably the same quality as before. It was realised that some revamping of the hydrocracker was inevitable, but the objectives had to be met at minimum capital investment. From a technology point of view, the requirements had the following impact: 9 In order to obtain the same quality hydrowax, feed conversion needed to stay at least at the same level. This is illustrated in Figure 11 showing that BMCI, which is a measure for hydrowax quality, deteriorates, for a given catalyst system and feed composition, with decreasing conversion (note that a higher BMCI value corresponds with a lower quality). 9 The higher throughput, the more difficult feed to process and the abandoning of catalyst activity booster, combined with the need to achieve the same cycle life at equivalent conversion level, called for almost a 100% increase of the intrinsic activity of the catalyst system. 9 Higher catalyst activities could, in principle, be attained by switching to new, more active pretreat and cracking catalysts. For the cracking catalyst, this meant a system containing a zeolite in higher amounts and/or having a higher intrinsic activity. But, as shown in Figure 12, such catalysts would definitely result in higher naphtha yields. The more so, since the extra naphtha make due to the higher amount of feed processed had to be absorbed as well. Moreover, with those catalysts, the hydrowax quality would also deteriorate (Figure 11). In the selection of catalysts, optimising the ratio between pretreat and cracking catalyst played an important role. Not only activity, yield and product quality aspects needed to be considered, but also catalyst activity decline rates and individual bed quench capacities had to be taken into account. Figure 13 indicates that the theoretical optimum is not always fitting with a discrete number of catalyst beds, further complicating the final choice to be made.
On top of these process aspects, additional hardware related equipment became critical too. The capacity of the fresh and recycle gas compressor, the feed furnace, the separators, the fractionator work-up section and the heat exchangers had to be checked in order to review whether they could cope with the more severe duty. The revamp project that started was a multi-disciplinary approach with close co-operation between refiner, technology provider and catalyst vendor. Feedback of detailed data from commercial operation for tuning of the process model and exploiting the new options for the hydrocracker to the limit of equipment capabilities, was a key step in successful execution of the project. The final option selected centred on a revamp of the fractionator allowing more naphtha draw-off, a choice for the more active DN-190 as pretreat catalyst and the more active Z-623 as cracking catalyst, and replacement of one bed of pretreat byr cracking catalyst.
Conclusions Since its origin in the 1960's, hydrocracking has become an important process in the refinery. Over the years many improvements in hydrocracking technology have been implemented, both from a process and a catalyst point of view. Although hydrocracking is a mature process, still new developments come to the fore, stimulated by a steady, further growth of the market and environmental pressures on product qualities. More than ever, hydrocracking technology developments in industry are nowadays based on a thorough knowledge of existing commercial operations. New catalysts are being developed not just because they have a superior performance over older ones, but because they maximise profitability in the refinery. Process developments are geared towards optilnising integration with other refinery processes via processing lower quality feedstocks from catalytic crackers, thermal crackers, residue conversion units etc., and producing high quality transportation fuels and feedstocks for ethylene crackers, catalytic crackers and lube base oil plants. The number of new hydrocrackers to be build in the next few years will be limited, caused by depressed refinery margins. A major part of the market will lie in optimising existing hydrocrackers. This requires attention to optimising catalyst packages, prevention of fouling, maximum utilisation of catalyst reactor volume, careful feedstock selection, but also revamping of equipment such as furnaces, gas compressors, separators, heat exchangers and fractionators. Accurate and detailed process models are a prerequisite to identify tailor-made solutions. Such models need to be based on both kinetic information from R&D experiments and data from commercial operation. Close co-operation between refiner, technology provider and catalyst vendor is a key factor in combining technology developments with market requirements.
Acknowledgement The authors wish to express their thanks to C.E.D. Ouwerkerk, M.C. Zonnevylle and J.R. Newsome from Shell Global Solutions, SIOP BV, Amsterdam, and W.H.J. Stork from Shell International Chemicals BV, Amsterdam, the Netherlands for their contributions to this paper.
References Hunter, M.G., Pappal, D.A. and Pesek, C.L., 1994. Moderate pressure hydrocracking: a profitable conversion alternative. Paper presented at the 1994 NPRA Annual Meeting, March 20-22, 1994, San Antonio. AM-94-21 2.
Dufresne, P., Bigeard, P.H. and Billon, A., 1987. New developments in hydrocracking: low pressure high-conversion hydrocracking. Catal. Today, 1: 367-384.
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Bridge, A.G., Cash, D.R., Law, D.V. and Scotti, L.J., 1994. Paper presented at the Annual International Refining Conference, May 9-12, 1994, Singapore.
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Hoek, A., Huizinga, T., Esener, A.A., Maxwell, I.E., Stork, W.H.J., Van de Meerakker, F.J. and Sy, O., 1991. New catalyst improves heavy feedstock hydrocracking. Oil Gas J., 89 (16): 77-82.
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Martindale, D.C., Abdo, S.F., Antos, G.J., Krenzke, D.K. and Mitchell, D.H.G., 1996. Continuing innovation in hydrocracking technology. Hydrocarbon Asia, 6(6):80-94.
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George, S.E., Boardman, S.R., Foley, R.M., Sanborn, L.J., Johnson, P.S., Webb, A., Gallagher, A., Gualtieri. P.K., Mok, W.S. and Nash, D., 1994. Hydrocracking to achieve product flexibility. Paper presented at the 1994 NPRA Annual Meeting, March 20-22, 1994, San Antonio. AM-94-19. Groeneveld, L.R., Stoop, F., Asim, M. and Brevoord, E., 1997. Akzo Nobel/Nippon Ketjen's hydroprocessing catalysts for hydrocracking applications. Paper presented at the 7th Annual Synaposium of Catalysts in petroleum refining and petrochemicals, Nov. 30 - Dec. 2, 1997, Dhahran.
10. Maxwell, I.E., Minderhoud, J.K., Stork, W.H.J. and Van Veen, J.A.R., 1997. Hydrocracking and Catalytic Dewaxing, in." Handbook of Heterogeneous Catalysis [G. Ertl, H. Kaaozinger, J. Weitkamp, (Eds.)], Wiley-VCH, 4: 2017-2038. 11. Desai, P.H., Gerritsen, L.A. and inoue, Y., 1999. Low cost production of clean fuels with Stars catalyst technology. Paper presented at the 1999 NPRA Annual Meeting, March 21-23, 1999, San Antonio. AM-99-40. 12. Carruthers, J.D. and Shukis, P.J., manuscript in preparation for publication in Hydrocarbon Processing. 13. Scherzer, J. and Gruia, A.J., 1996. Hydrocracking science and technology. Marcel Dekker, Inc. New York., chapter 3.1. 14. Huizinga, T., Theunissen, J.M.H., Minderhoud, J.K. and Van Veen J.A.R., 1995. New hydrocracking catalysts increase throughput, run length. Oil Gas J., 93(26): 40-45 15. Esener, A.A. and Maxwell, I.E., 1989. Improved hydrocracking performance by combining conventional hydrotreating and zeolitic catalysts in stacked bed reactors, in. Advances in hydrotreating catalysts, Annual AIChE meeting, Nov. 27 - Dec. 2, 1988, Washington D.C. [M.L. Occelli and R.G. Anthony (Eds.)],. Elsevier, Amsterdam, 263-271. 16. La Vopa, V. and Satterfield, C.N., 1988. Poisoning of thiophene hydrodesulfurization by nitrogen compounds. J. Catal., 110: 375-387. 17. Stangeland, B.E., 1974. A kinetic model for the prediction of hydrocracker yields. Ind. Eng. Chem. Proc. Des. Dev., 13(1): 71-76 18. Van Zijp, R. and Krishna, R., 1993. A kinetic hydrocracking model for aromatic feed. Erdol, Kohle, Erdgas, Petr., 46(3): 98-102 19. Quann, R.J. and Jaffe, S.B., 1996. Building useful models of complex reaction systems in petroleum refining. Chem. Eng. Sci., 51 (10): 1615-1635 20. Hansen, J.A. and Cooper, B.H., 1992. Process simulation of refinery units including chemical reactions. Computers and Chemical Engineering, 16:$431-$439.
11 21. Laxminarasimhan, C.S.,Verma, R.P. and Ramachandran, P.A., 1996. Continuous lumping model for simulation of hydrocracking. A.I.Ch.E.J., 42(9): 2645-2653. 22. Martens, G. and Froment, G.F., 1999. Kinetic modelling of paraffins hydrocracking based upon elementary steps and the single event concept, in." Reaction kinetics and the development of catalytic processes [G.F Froment and K.C. Waugh (Eds.)], Elsevier, Amsterdam, 333-340. 23. Yeary, D.L., Wrisberg, J., and Moyse, B., 1997. Revamp your internals. Int. J. Hydrocarbon Eng., 2(5): 25-29 24. Patel, R.H. and Bingham, E., 1998. Hydroprocessing reactor and process design to optimize catalyst performance. Paper presented at the First Indian Refining Roundtable, Dec. 1-2, 1998, New Delhi. 25. Ouwerkerk, C.E.D., Bratland, E.S., Hagan, A.P., Kikkert, B.L.J.P. and Zonnevylle, M.C., 1999. Performance optimisation of fixed bed processes. Petr. Tech. Quarterly, 4(1): 21-30. 26. Sanford, E.C. and Kirchen, R.P., 1988. |mproved catalyst loading reduces guard reactor fouling. Oil Gas J., 86(51): 35-41
Hydrogonetion
Acidic*
Table 1 - Optioua for Lyd-
amd emking hactloa
Lorr HC
PPHZ bar
PrrrrrrEigh
HC
M
150
50
50
WHSV, Lgl(ll)
0.5
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Cycle I* yeus Kero smoke point, mm Gas oil cctam index
I
>5
10
30
40
60
VGO comwsion at T
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-
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Hydraerwirtr-biem
-
cfiect 00 b y d m m d p&brmance ~
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UmimprwedandElilolvnadeaa!alysts
Processheawierandlorcheapcrfdstds Prcidtw mom valuable products by s$iffiag the p d u c t slab Use an (on-lme) OprimWon modtl
Pusb Mmequipment to the limiep, but be a m ofpmme drop aad Wing Inmuam the number of swamdays per year by bedm maintmwe and imploved rsliability Low cm?t mvamp of amstmining equipment
Table 3 - Items for optimising existing hydrocrackers
RVA for HDN
100
130
-
Table 4 Colrpnrtbu ofC-424 amd DN-198in VGO kydrocrackhg
- --
C
d first ~f8ge[%MI
25
40
Table 6 - Improving two stage hydrocrackers by optimising first stage catalyst system
14 Once Through[
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Figure 1 - Hydrocracking modes of operation
Single Stage Stacked Bed Hydrocracker Fresh gas
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ench
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15
HIGH DISPERSION
CONTROLLED ACIDITY
Figure
3 -
The concept behind the development of DN-190
Figure 4- Comparison of various hydrocracking catalysts
OPTIMUM PORE STRUCTURE
/-
NiW, e.m. 3
NiW, e.m. 2
NiW, e.m. 1 NiMo
ASA acidity
-
Figure 5 - Balancing hydrogenation function (via emplacement of metals, em.) and acidic function for optimising performance (activity, diesel yield) of ASA catalysts
Feed N content bpmwt]
Figure 6 VGO
-
Inhibition of the cracking reactions by nitrogen in hydrocracking of pretreated
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108
300
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Figure 8 - Comparison of liquid distributions for two types of trays
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Figure 10 - Results from a mixing test on a UFQ internal
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Figure 12 - Effects of the type of zeolite catalyst on naphtha yields in VGO hydrocracking over stacked beds of pretreat and zeolitic cracking catalyst.
Ylnimum tequired
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Figure 13 - Cycle length optimisation for a stacked bed of pretreat and cracking catalyst
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
21
Life cycle of hydroprocessing catalysts and total catalyst management Sonja Eij sbouts Akzo Nobel Chemicals b.v., Research Centre Catalysts, Nieuwendammerkade 1-3, P.O. Box 37650, 1030 BE Amsterdam, The Netherlands;
[email protected] Abstract
Ni/Co-Mo/W hydroprocessing catalysts are commercially used in different refinery applications over a wide range of conditions. Depending on the application, different deactivation mechanisms are of importance (coke formation, active phase sintering, metals deposition, poisoning). Except for extremely contaminated catalysts from residue or heavy vacuum gas oil applications, most catalysts are regenerated and reused. During their life cycle, they undergo several transformations from the oxidic into the sulfidic state and vice versa. Their life cycle is, therefore, very complex and involves many different steps and aspects. Not only it is extremely complicated from the fundamental point of view: There are also numerous technical, environmental as well as purely organizational issues involved. The leading catalyst manufacturers together with specialized firms offer the refineries total catalyst management during the entire catalyst life cycle, starting with the purchase of the fresh catalyst and ending with its final recycling or disposal. Total catalyst management includes a broad range of services, ensuring optimal timing during the change-out process, reliable, smooth and safe operations, minimal downtime and maximum catalyst and unit performance.
1. INTRODUCTION Nowadays, the oil industry deals with more and more demanding economic and environmental requirements. The refineries must operate more efficiently (lower costs, less downtime) and increase the operability of their units (higher throughputs, longer cycle times, more catalyst cycles). The product specifications (e.g. sulfur, aromatics, and cetane number) get stricter due to the environmental regulations, which affect virtually any refinery process as well as any handling of refinery catalysts. These trends have introduced new aspects in the catalyst design and have modified the way the catalyst manufacturers, users and firms providing the logistic, activation, regeneration and reclaiming services work nowadays. The hydroprocessing catalyst life cycle starts with the production of the oxidic catalyst, which is then presulfided. During its use, the sulfidic catalyst deactivates by active phase sintering, coking, metals deposition and by poisoning. The reactor is shut down as soon as the catalyst does not meet the performance targets within the limits of the reactor operating conditions. Depending on the degree and nature of catalyst contamination, the catalyst can either be oxidatively regenerated, has to undergo an additional rejuvenation treatment or has to be disposed of to reclaim metals for re-use. The regenerated (oxidic) catalyst is again
22 presulfided prior to further use in the refinery. The hydroprocessing catalyst life cycle includes thus a large number of separate steps and catalyst handlings. For a catalyst regenerated twice, the price of all catalyst-related services during its two cycles lays in the same order of magnitude as the fresh catalyst price.
2. CATALYST PREPARATION AND ACTIVATION
2.1 Preparation of oxidic catalysts Just like the refineries, also the catalyst manufacturers are limited by economic and environmental requirements. There is only a limited number of raw materials and additives that are well priced and environmentally acceptable. The waste water and gas streams are purified and the solid waste is reprocessed or, if necessary, disposed. To ensure the safe handling of all product and waste streams the production process is carried out in sealed and ventilated production units. The workers involved in the production and further handlings (packaging, transport, loading etc.) use respiratory protection. The catalysts are typically packed and transported in sealed drums. There is a tendency to use larger packaging (e.g. big bags and containers), enabling faster reactor loading and minimizing packaging costs. Generally, the oxidic catalysts are prepared by single or multiple impregnations of a shaped 7-A1203 support with Ni/Co and Mo/W containing solutions. Typical raw materials are nitrates of Ni/Co and ammonium salts of Mo/W [1-3], sometimes in combination with organic or inorganic complexing agents [4-5]. Nitrate, ammonia and the organic complexing agents decompose during the subsequent drying and calcination and their decomposition products are removed from the gas stream in e.g. a scrubber or a DeNOx unit. The fresh oxidic catalysts typically contain well dispersed CoO/NiO and MoO3/WO3 on 7-A1203 support. The oxidic catalysts are stable during the packaging, transport, storage and loading in the reactor. Despite the economic and environmental limitations, leading catalyst manufacturers regularly introduce new products, having improved performance to meet more stringent product specifications and yet competitive prices. The new products not only offer an improved activity and/or selectivity but typically also a good stability and regenerability as well as outstanding mechanical properties. The newly introduced catalysts are usually closely related to the existing hydroprocessing catalysts. This means that these catalysts are 7-A1203 supported Ni/Co-Mo/W catalysts, differing only marginally (by e.g. additives or metal contents) from the existing grades. This conservative approach to the design of improved catalysts is dictated by a number of refinery related factors. Firstly, the existing catalysts have outstanding properties (activity, stability, regenerability, mechanical properties, and price). Secondly, the design of the existing refinery units is based on and the refineries are used to working with this type of catalysts. Thirdly, the application of newly developed catalysts, strongly deviating from the existing ones, includes major risks for the refinery (e.g. unknown stability, i.e. maintenance of performance and mechanical properties during the long-term catalyst use). Fourthly, their price is often much higher than that of existing catalysts and their application may require major changes of refinery units, resulting in additional costs.
23 This has a huge impact on the commercialization of new materials described in the open and patent literature. Besides catalysts prepared in more or less conventional way [1,4-9], there are also many materials, which are very innovative [10-33]. Catalysts prepared directly in the sulfidic form [11-13], new types of supports [17,22-26,32,33] as well as new active components such as niobium trisulfide [10], Mo/W nitrides and carbides [27] or noble metals [14-16,28-31] have been described. Even though these materials offer advantages such as higher intrinsic activity, lower coke propensity, less cracking or less sintering of the active phase, they may be less favorable for commercial application because of their limited stability and regenerability or less good mechanical properties. The fact that the price of such materials is usually much higher than that of the existing commercial catalysts is also very important. This is due to 1. high costs needed for scaling-up of their production process, 2. expensive ingredients and/or 3. high costs of additional environmental measures if unusual toxic components are involved in their production. All in all, it is understandable that the commercialization of new materials is extremely complicated.
2.2 Presulfiding To prevent any fluctuations of the reactor performance and to achieve optimum activity in the initial stage of the refinery operation the oxidic catalysts are presulfided prior to their use. The catalyst manufacturers typically provide a loading advice, performance prediction as well as start-up recommendation and assistance. The choice of the presulfiding procedure depends on the needs and possibilities of the specific refinery unit. When in situ presulfiding is applied, the oxidic catalyst is loaded in the reactor and contacted either with H2 and feed having a high S content (optionally increased by spiking with S compound such as e.g. dimethyl disulfide) [34,35] or with a sour gas (H2S in H2) [36], while gradually increasing the reactor pressure and temperature. Nowadays, there is a clear tendency to use the liquid rather than the gas phase presulfiding. However, especially higher pressure commercial units are suitable only for gas phase start-up. Ex situ presulfiding is carried out partly in the installation of a specialized firm and partly in the refinery reactor [37-41]. The oxidic catalyst is impregnated with a S containing compound such as elemental S [42-43], organic monosulfide [44], disulfide [44] or polysulfide [45-48]. Elemental S can be introduced by melt impregnation [42-43]. The organic S compounds are typically dissolved in an organic solvent [45-47]. After the impregnation with the S compound, the catalysts are usually dried. Under these conditions, S has not yet or has incompletely reacted with metal oxides. There are also processes, enabling a complete ex situ sulfidation of the catalyst [41,49] with e.g. H2S in H2 [49]. Most ex situ presulfiding treatments are followed by a passivation with O containing compounds [50-53]. The passivation treatment reduces the self-heating properties of the materials. Despite that, the S containing catalysts, originating from either of these processes, are classified as self heating and have to be packed and transported in special small sealed (air tight) containers [37,39]. The S containing catalysts are then loaded in the refinery reactor, often under inert atmosphere, and contacted with feed and H2, while gradually increasing the reactor pressure and temperature. As the catalyst already contains S, the S content of the feedstock does not have to be increased, eliminating thus the difficult handling of toxic and odiferous S
24 compounds in the refinery. Moreover, the reactor start-up is less time consuming as the ex situ presulfided catalysts require little or no special handlings during the reactor start-up [37]. At present, about 20 % of hydroprocessing catalysts is sulfided ex situ. The costs of ex situ presulfiding are approximately 10 - 15 % of the fresh catalyst price. The sulfidation behavior of 7-A1203 supported Ni/Co-Mo/W hydroprocessing catalysts has been extensively investigated. Besides studies relating the catalyst performance to the presulfiding conditions [54-57], the sulfiding has been modeled by temperature programmed sulfidation [9,58-61] and the catalysts have been characterized in different stages of the presulfiding procedure [62-67]. For obvious reasons (simple and available equipment), the laboratory presulfiding studies are mostly carried out in the gas phase, using H2S/H2 mixtures. It is, however, reported that the liquid phase presulfiding may provide a better temperature control than the gas phase presulfiding processes [35]. Freshly sulfided hydroprocessing catalysts are usually characterized by high dispersion, homogeneity and absence of crystals. They contain well dispersed MoS2/WS2 slabs and stacks with edges decorated with sulfidic Ni/Co [68-69]. A part of the Mo/W remains oxidic, maintaining the Mo/W-O-A1 linkages, facilitating the high dispersion and the perpendicular position of the MoSz/WS 2 structures on the support. The role of Mo/W-O-A1 linkages is eliminated if the sulfidation is carried out at high temperature (typically above 450 ~ [69]).
Figure 1. TEM micrographs of a Ni-Mo/A1203 catalyst sulfided: a. with spiked feed at a temperature below 350 ~ and b. with H2S/H2 at a temperature above 450 ~
As already mentioned, the aim of the presulfiding is to ensure a stable and high unit performance in the initial part of the run. Thus, the goal is not to fully sulfide all of the metal oxides present but to bring the catalyst in equilibrium with its reaction environment. Forcing the presulfiding conditions in a way assuring a complete sulfidation (high temperature
25 sulfidation) or an inferior temperature control during the sulfidation (temperature runaway due to exotherms) leads to a fully sulfidic but also highly sintered catalyst. Figs. l a-b show the difference between catalysts sulfided at low temperature in the liquid phase and at high temperature in H2S/H2. The exposure to high temperatures not only leads to MoS2 dispersion loss but also to the segregation of NiSx. This altogether stresses the importance of good process control during the presulfiding. This can only be achieved in a close cooperation between the catalyst supplier, the refinery and the company carrying out the presulfiding.
3. CATALYST USE AND DEACTIVATION During their use, hydroprocessing catalysts deactivate due to active phase sintering and segregation [68-69], due to blocking of the catalyst pores [70-72] and active sites by coke and metals deposits [73-75] and due to poisoning of the active sites [73-75]. The catalyst cycle length and regenerability depend on the most dominant deactivation mechanism, i.e. on the exact application and reaction conditions the catalyst is exposed to. The cycle length can be as long as 5 to 10 years for naphtha units, operating on light, metal contaminant free, feed under very mild conditions. For residue units, operating with heavy feeds, containing metals (Ni, V), under very severe conditions, the cycle length can be as short as 0.5 to 1 year.
3.1 Sintering and segregation of the active phase The last Mo/W-O-A1 linkages get broken as Mo/W gets more completely sulfided under the reaction conditions [68-69]. MoS2/WS 2 clusters oriented parallel to the support migrate over the ~-A1203 surface and sinter. The relative amount of NiSx/COS• on the MoSJWS2 edges gets very high. Ni/Co can not be built in the MoSJWS2 matrix and segregates as Ni3SJNiS/Co9S 8. As the segregated sulfides have lower activities than the original mixed phase, this process contributes substantially to the overall catalyst deactivation, especially in catalysts having high metal contents. The sintering and segregation process has been studied by several authors [68-70,76-79]. These studies include characterization of used catalysts at different stages of their life. The sintering and segregation process can only to a limited extent be restored by oxidative regeneration. 3.2 Coke deposition The major part of coke is deposited on the catalyst surface in the initial part of the run [80]. The coke concentration remains more or less constant during the catalyst use but its composition changes. The H rich "soft" coke (- cycloparafins, naphthenes, aromatics...) present at the start of run changes into "hard" coke (= polyaromatics) by dehydrogenation during the further catalyst use. This process is faster at high temperature and low H 2 partial pressure, i.e. at the end of run and in the bottom part of the reactor. The coke deposition has been extensively studied and modeled [70,76,81-87]. In theory, the catalyst can by itself "remove" a part of its coke by hydrogenation under more favorable conditions (increased H2 partial pressure and lower temperature). However, due to unit limitations, this on-stream
26 reductive regeneration usually can not be carried out [88-89]. The oxidative regeneration can nearly completely remove coke and restore the activity loss due to coking [90-92].
3.3 Metals deposition Ni and V porphyrins contained in heavy vacuum gas oil (VGO) and residue feeds decompose rapidly on contact with the catalyst. Small NiSx and VSx crystals deposit on the catalyst surface and can lead to pore blockage. While normal hydroprocessing catalysts can not tolerate more than about 2 - 3 wt.% Ni + V, VGO and residue hydroprocessing catalysts can tolerate up to 100 wt.% Ni + V (based on the fresh catalyst weight). The metals deposition has been studied and modeled by numerous authors [81-83,93-98]. The metals and coke deposition models have been used to explain the so-called S-curves describing the deactivation of residue hydroprocessing catalysts [80,83]. The deactivation by metals deposition can not be reversed by oxidative regeneration. A selective leaching of the contaminants (= rejuvenation) may be carried out in addition to oxidative regeneration [99104]. However, as the activity recovery of the rejuvenated catalyst mostly remains incomplete, highly contaminated catalysts are usually not regenerated. Also other contaminants (e.g. Si, originating from anti foaming agents added in the refinery) may deposit on the catalyst as separate particles blocking the pores [105]. Conventional hydroprocessing catalysts can tolerate up to 15 wt.% Si. Si contamination at levels above 2-3 wt.% can not be restored by oxidative regeneration. Special guard catalysts have been developed to protect the hydroprocessing catalysts from metals and Si fouling. 3.4 Poisoning Also other compounds may contaminate the catalyst during its use. Poisoning by oxygenates [106] and N compounds [107] is in principle reversible under the reaction conditions and can be restored by oxidative regeneration. Elements such as Pb and As specifically poison the active sites [41 ]. Pb, introduced in the refinery as a gasoline additive [(C2Hs)4Pb], poisons the catalyst if a break through occurs. Arsenic, present in some crude oils, is even a more severe poison than Pb. Most catalysts can not tolerate more than 0.5 wt.% Pb and/or 500 ppm As. Pb and As poisoning can not be restored by oxidative regeneration. 3.5 Refinery processes Hydroprocessing catalysts are used in a wide range of applications, under very deviating conditions and with different feeds. Depending on the exact operating conditions and feedstock, one of the above mentioned deactivation mechanisms may become dominant. In distillate hydrotreating, clean light, metal contaminant free, feeds are processed at low reaction temperature (320 - 360 ~ and pressure (20 - 60 bar), using high dispersion medium metal content Co-Mo catalysts. The deactivation is mainly due to coke deposition and the catalyst performance can be nearly fully restored by oxidative regeneration. Multiple regenerations (up to 4 cycles) are possible, provided the catalysts maintain their mechanical properties (crushing strength, particle lengths distribution). In hydrocracker pretreatment, VGO, sometimes containing some Ni and V, is treated at high reaction temperature (380 - 400 ~ and pressure (90 - 200 bar), using medium
27 dispersion - high metal content Ni-Mo catalysts. The deactivation is due to coke deposition and due to the sintering and segregation of the active phase. The catalysts can be reused but their performance can not be fully restored by oxidative regeneration. In FCC pretreatment, VGO feeds, containing some Ni and V, are processed at medium reaction temperature (350 - 370 ~ and pressure (50 - 100 bar), using medium dispersionmedium metal content Ni-Mo catalysts. The deactivation is due to sintering, coke and metals deposition. The performance can not be fully restored by oxidative regeneration. The catalysts can only be regenerated and reused if they are not too much contaminated by metals. In residue hydrotreating, residue feeds, containing much higher levels of Ni and V, are treated at high temperature (380 - 400 ~ and pressure ( 1 0 0 - 160 bar), using high dispersion low metal content catalysts. The deactivation is mainly due to coke and metals deposition. The performance can not be sufficiently restored by oxidative regeneration so that these catalysts mostly can not be reused. 3.6 Used catalysts Used hydroprocessing catalysts are characterized by their lower dispersion, inhomogeneity and presence of contaminants and crystals [68,69,108]. They contain larger MoSJWS2 slabs/stacks still decorated with NiSx/CoSx as well as segregated MoSz/WS 2 and NiSx/CoSx crystals. Crystals of metal deposits (NiSx and VSx from feed + FeS• from unit corrosion) and other contaminants (As, Pb, Si, alkali and alkali earth metals.., etc.) may be present as well.
Figure 2. TEM micrographs of a spent Ni-Mo/A1203 catalyst containing a. crystals of NiS• and VSx metal deposits and b. segregated NiSx crystals. The coke content is usually between 5 and 30 wt.% [109,110]. The used catalysts are toxic, self heating and sometimes even pyrophoric [111 ]. Special unloading techniques have to be
28 applied (e.g. under N2) and the catalysts have to be transported and stored in sealed (airtight) small containers [41,90,109]. As discussed above, the final state of the used catalyst depends on the reaction conditions it was exposed to [68]. Catalysts used under the distillate hydrotreatment conditions contain some coke (ca. 10 wt.% C) and are characterized by their high dispersion and homogeneity and by the absence of crystals. Catalysts used under the hydrocracker pretreat conditions are still homogeneous but their dispersion is lower and they may contain also some large MoS2 or NiSx crystals. Catalysts used under FCC-pretreat conditions have a higher MoS2 dispersion but are contaminated with NiSx, VSx and FeSx deposits (Fig. 2a). Catalysts used at very high temperatures are inhomogeneous and contain large MoS2 or even NiSx crystals (Fig. 2b). Catalysts containing large amounts of metal deposits (Fig. 2a) or many segregated NiSx crystals (Fig. 2b) usually can not be regenerated and have to be disposed. Extensive analysis and a discussion between the catalyst supplier and a professional regeneration company is necessary to decide on how the used catalyst should be further processed.
4. CATALYST REGENERATION AND REJUVENATION Modern hydroprocessing catalysts usually have high activity and stability as well as good mechanical properties and regenerability. The activity restoration on regeneration depends on the degree of the active phase sintering and on the amount of contaminants present in the catalyst [40, 110]. Typically, the used catalyst should not contain more than 2 - 3 wt.% of metal contaminants (Ni + V) to be suitable for oxidative regeneration. After regeneration, the catalyst must reach sufficient activity level and its physical (surface area, pore size distribution) and mechanical (crush strength, length distribution) properties must be on target [40,41,73,110,111]. Especially the deterioration of mechanical properties (particle breakage) is a very important aspect limiting the catalyst regenerability. In general, non-residue catalysts are typically regenerated once or twice. As not all regenerated catalysts can be reused by the refinery where the used catalyst came from, there are special firms operating in the field of catalyst resale and pooling [41 ]. 4.1 Oxidative regeneration Oxidative regeneration removes the coke and transforms Mo/W and Ni/Co sulfides back into oxides [91-92,112-113]. On reoxidation, Mo can be partly redispersed through solid-solid wetting of MoO 3 on 7-A1203 [114,115]. The redispersion of NiO/CoO or WO3 via solid-solid wetting is not possible under typical regeneration conditions [69,116]. Large MoS2/WS2 and NiS]CoSx crystals are oxidized only on the surface. The oxidative regeneration can neither remove the metal deposits nor restore the sintering and segregation of Ni/Co and W components of the active phase. The sintering of the active phase [ 117-119] and of the support [114,115,120-122], Mo losses [123] and the formation of stable/inactive compounds such as Ni/CoA1204 [38,124], Ni/CoSO4 [125], Ni/CoMoO4/WO 4 [57] or A12(MoO4/WO4)3 [69] can be suppressed by a careful temperature control, preventing exotherms.
29 The oxidative regeneration can be carried out in or ex situ [39-41,110,126]. It can be preceded by de-oiling/stripping of the catalyst (mostly treatment with N 2 at high temperature to remove adsorbed hydrocarbons) [41]. The oxidation is usually carried out in two stages [90] 1. at low temperature and 02 concentration to remove the reactive coke and S and 2. at higher temperature and 02 concentration to remove the non-reactive coke. The ex situ oxidative regeneration provides a much better process control and avoids environmental problems at the refinery caused by the regeneration off-gas containing SO x and COx [3941,110,126]. The reactor corrosion due to SOx formation is eliminated and the chance for accidents, hot spots or reactor malfunction is lower. Dedicated catalyst specific regeneration procedures can be applied and the fines can be removed by screening. That is why e.g. in Europe about 90 % of all regenerations is carried out ex situ at present [41 ]. The costs of the oxidative ex situ regeneration are around 20 % of the fresh catalyst price. Various oxidative regeneration procedures are described, from conventional ones [127128] to very sophisticated procedures such as e.g. regeneration of coked catalysts by laser irradiation and oxidizing gas specially designed to suppress exotherms [129,130]. An alternative to oxidative regeneration is on-stream reductive catalyst regeneration, typically carried out at temperatures lower than the reaction temperature [88] and optionally by applying a lighter feedstock [89].
4.2 Regenerated catalysts After resulfidation, the regenerated catalysts comain again MoS2/WS 2 slabs/stacks decorated with NiSx/COSx [68,69] (Fig. 3a). Conglomerates of large MoS 2 stacks (Fig. 3b), segregated crystals of NiSx/CoS• and WS2, crystals of NiS• and VS• deposits and, incidentally, Ni/CoA1204, Ni/CoSO4, Ni/CoMoO4/WO4 or A12(MoOJWO4) 3 may be present.
Figure 3. TEM micrographs of a regenerated Ni-Mo/A1203 catalyst containing a. areas with well dispersed MoS2 and b. conglomerates of large MoS2 stacks.
30 The lower dispersion and homogeneity and the less intimate contact between the active phase components distinguish the regenerated catalyst from the fresh one and explain its lower activity. The temperature and 02 concentration control during the regeneration is critical [57,120,125,131,132]. If the catalyst is exposed to high temperatures at high 02 concentration the S and coke removal is complete but the catalyst is somewhat sintered and contains crystalline Ni/CoMoO4 [57]. If the catalyst is exposed to high temperature at low O2 concentrations, coke will be partly transformed into carbonaceous deposits that can not be removed by oxidation within the temperature range usually applied during the regeneration treatment [120]. If the temperature is too low S2- may be transformed into SO42- and both S and coke removal are incomplete [125]. The proper choice of regeneration conditions is thus determining for the success of the regeneration. Clearly, the chance of choosing the right conditions is much higher if the regeneration is carried out ex situ, i.e. in equipment assuring an optimum temperature control and a good air - catalyst contact and if the regeneration company closely cooperates with the catalyst supplier.
4.3 Rejuvenation An additional rejuvenation treatment can be applied for medium contaminated catalysts. The contaminants must be removed without affecting the underlying catalyst. For example, V (which is usually deposited on the exterior of the catalyst pellets) can be selectively removed by attrition. However, such a treatment affects negatively the catalyst mechanical properties. A selective leaching [99-104] of V is usually even more destructive as it may also affect the active phase components. Another possibility to improve the spent or regenerated catalyst performance is by an additional impregnation with the active components [133,134]. The interstage impregnation of spent and/or regenerated catalysts with group IV [135] or group IIA metal [136] makes the removal of contaminant metals unnecessary and/or increases the attrition resistance of the catalysts during further treatments. The price of such rejuvenation treatments can be relatively high (15 - 35 % of the fresh catalyst price or even higher) and their success may be limited. Moreover, each additional treatment further deteriorates the catalyst mechanical properties so that a large portion of the catalyst may be lost due to excessive breakage. The rejuvenation is, therefore, only rarely applied to increase the activity of regenerated catalysts.
5. CATALYST RECYCLE AND DISPOSAL The used catalyst is disposed of if the performance could not be brought to the desired level or if the mechanical properties would strongly deteriorate on regeneration [111 ]. The disposal has to be carried out according to the regulation for treating of dangerous goods [109]. As the prices of metals fluctuate strongly, the exact way of disposal depends on the economics of the moment [73]. It may vary from reclaiming of expensive components [137] to removal or stabilization of toxic components to make further disposal possible. Alternatives to disposal may be the reuse [101,138] of the catalyst in less demanding refinery operations or
31 its use as hot gas clean-up sorbents (e.g. to remove H2S) [111,139-140]. There are specialized used catalyst brokers, providing different catalyst disposal and recycling services.
5.1 Metals reclaiming The reclaiming of transition metals from spent hydroprocessing catalysts is usually carried out by roasting the catalyst with e.g. NazCO3 and by precipitation [109,111,126]. The precipitate is then treated with different agents in order to come to a selective leaching of metals of interest [111,141]. Typical products of selective leaching are [109] 1. solid containing alumina, silica, Co, Ni and small concentrations of Fe, Mo, W and V and 2. solution containing Mo, W, As, P and low concentrations of Si and A1. The solid is further treated to isolate Ni and Co and the liquid is treated to obtain pure Mo, W and V compounds. The typical costs of the metals reclaiming are about 10 % of the fresh catalyst price. 5.2 Catalyst stabilization If the catalyst is to be disposed as solid waste it must be transformed into an inert nonleachable solid suitable for storage [111]. The roasted catalyst can be e.g. encapsulated in thermoplastic agents (bitumen, paraffin wax, polyethylene etc.). However, the encapsulating agents are flammable and may undergo a long-term deterioration. The catalyst can also be stabilized by the reaction with e.g. alumina, cement or silicate glass at high temperature. In this way, stable/non reactive inorganic compounds (silicates, aluminates) are formed and the material is at the same time encapsulated. The disadvantage of the stabilization is the possible slow devitrification and fracture of the particles. 5.3 Catalyst storage The stabilized catalysts can be stored by so-called landfilling in separate lined cells capped and isolated from each other and from the environment [111]. The above mentioned imperfections of the stabilization treatment together with the more stringent environmental rules have lead to a rapid reduction of landfilling. In 1993, 15 - 25,000 tons of spent HPC catalyst were stored worldwide in an approved way (-- recoverable) and about 10,000 tons were dumped unpacked (--- non-recoverable) [111]. Nowadays in Europe, perhaps only a few hundred tons of spent HPC catalyst a year are stored in approved landfills and there are no more reports of unapproved landfills.
6. CONCLUSIONS The environmental aspects play an important role in all stages of the life cycle of a hydroprocessing catalyst. A responsible care by the catalyst manufactures, refineries and firms providing presulfiding and regeneration services secures that the catalysts are handled in a safe manner, using approved procedures, throughout their entire life, including the final disposal. The catalyst manufacturer can help the refinery to select the best catalyst for the given application and to extend the catalyst life by supplying special guard beds, protecting the catalysts from poisoning. In a close cooperation between the catalyst supplier, the refinery
32 and the firms providing presulfiding and regeneration services, optimum process conditions for activation and regeneration treatments can be selected. Ex situ sulfidation and/or regeneration carried out by a specialized firm may be the best alternative in cases where the process control is really critical. Nowadays, the different steps of the catalyst life cycle are often treated and arranged separately. However, there are also partnerships in the field of catalyst management during the life cycle. These partnerships ensure that not only the individual handlings are optimized but also that the transition from one step to the next one is carried out in an optimum way. Catalyst management means the overall coordination of all elements in the entire catalyst life cycle. The objectives are 1. to increase catalyst and unit performance, 2. to stimulate efficiency and 3. to increase the operability of hydroprocessing units. With product stewardship becoming more and more important in the industry, it can be visualized that the leading catalyst suppliers will no longer offer just a product to the market, but also services like catalyst activation, regeneration, reclaiming, leasing of catalysts, packaging and coordination of reactor unloading, cleaning, inspection and loading activities.
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36
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
37
Hydrogen Spillover and hydrocracking,-hydroisomerization K. Fujimoto Department of Applied Chemistry, School of Engineering, The University of Tokyo, Hongo, Bunkyo-ku, Tokyo 113-8656, Japan
Abstract It was found that proton on Br0nsted acid site of zeolite is easily exchanged with dihydrongen and that the desorption of adsorbed pyridine on Lewis acid site is promoted by gaseous hydrogen when zeolite carded noble metal or was hybridized with supported noble metal. Hydrogen in the gas phase was assumed to bedissociated into atomic fiydfogen on the noble metal and move onto zeolite, where hydrogen atom is converted to either proton (H § and hydride ion (H). It was claimed that the common characteristic feature of hydroconversion (hydrocracking, hydrotreating, hydroisomerization) catalysts is that the catalyst is composed of two functions: hydrogen activation and another one (acid, desulfurization et al.). Hydroisomerization catalyst, which is usually composed of solid acid and supported platinum on it, and has been thought to be a typical bi-functional catalyst, can be substituted by the physical mixture of Pt/Si02 or Pd/Si02 and zeolite. It was suggested that proton on acid activate paraffinic hydrocarbons to carbenium ion to catalyze isomerization or carcking reaction then the spilt-over hydride ion react with carbenium ion to make stable hydrocarbons.
1. INTRODUCTION Spilt-over hydrogen is known to have strong effect on catalytic reaction system on solid acid catalysts. Nakamura et al. found that hybrid catalyst, i.e. the physical mixture of Pt/SiO2 and HZSM-5, was very effective for the isomerization of n-pentane, n-hexane and other paraffinic hydrocarbons under hydrogen atmosphere [1-5]. They concluded that both high conversion and high selectivity were due to the effect of hydrogen spillover, and suggested that spilt-over hydrogen has two forms H § and H, and that the former regenerates Bronsted acid site while the latter stabilizes carbenium ion intermediate by its hydrogenation. Hattori et al. pointed out that hydrogen promoted the activity of cumene cracking o v e r Pt/SO42-ZrOz and inhibited its deactivation. It is expected that Br0nsted acid site generated from spilt-over hydrogen acts as the active site for the catalytic reaction [6]. Hosoi et al. found that when Pt/SOa2-ZrO2 catalyst was used for skeletal isomerization of n-pentane in the presence of hydrogen, it showed not only high activity but also persistence of the activity for a long period, more than 1000 h. They explained that the hydrogen had the effect on the removal of coke formed during the reaction by hydrogenating it [7]. It has been reported that spillover of hydrogen occurs even when physical mixture of supported metal catalyst and zeolite, namely, hybrid catalyst, is exposed to hydrogen atmosphere [8]. Hydrogen molecule in gas phase is
38 dissociated on metal surface to atomic hydrogen at first and then migrates onto the support surface, and further, onto the surface of adjacent zeolite particles. Zhang et al. reported that pyridine chemisorbed on acid sites of zeolite could be hydrogenated into piperidine by spiltover hydrogen [9]. The phenomena of spillover first noticed was the promoted decomposition of GeH4 on a Ge film with a Pt wire [10], the reduction of WO 3 to WO2 by hydrogen at room temperature for a mechanical mixture of WO3 with Pt/A1203 [11] and the accelerated reduction of NiO by hydrogen when mixed with palladium or platinum [12]. Als0, isotopic exchange of OH groups on Al:O3 due to spillover was reported in 1965 [13]. Reverse hydrogen spillover on active carbon or zeolite was claimed as the key step of paraffin dehydrogenation [14]. Important phenomena caused by spillover were summarized in a several reviews. They are (1) enhanced adsorption, (2) surface isotopic exchange, (3) bulk change, (4) strong metalsupport interaction (SMSI). Influences of spillover on catalytic process may be described as (1) spilt-over species keeps catalyst clean, (2) create or regenerate selective sites through a remote control mechanism, and (3) as a regult, catalytic reactions are accelerated and catalyst deactivation is inhibited, effectively. There are a lot of discussion about the nature of spilt-over hydrogen species, such as H atoms, radicals, H § and H ions, ion pairs, H3§ species or protons plus electrons [15,16]. Protons formed from spilt-over hydrogen are suggested to act as catalytic active site for acid catalyzed reaction [17-19]. The present authors have pointed out the possibility of the participation of H § and H which are produced from spilt-over hydrogen in the hydroisomerization or hydrocracking of aliphatic hydrocarbons over Pt or Pd-supp0rted zeolite or physically mixed Pt/SiOE-protonic zeolite system [1,5,20]. Roland et at. have given a clear proof for the electrical charge of the spilt-over species, which was obtained through H-D exchange studies on the influence of a homogeneous magnetic field. The migration of spiltover hydrogen (deuterium) in Pt/NaY-HNaY catalyst was hindered, which was attributed to the influence of the Lorentz force on the electrically charged moving particles [21,22]. Pyridine is a typical organic base and can be chemisorbed on both BrCnsted (B) acid sites and Lewis (L) acid sites in zeolite catalyst while pyridinium ion and coordinately bonded pyridine complexes are formed on B and L sites, respectively-giving different m-adsorption bands on each occasion [23-25]. Zhang et al. reported that they found by FTIR that pyridine strongly adsorbed on acid sites of H-ZSM-5 was hydrogenated over Pt/H-ZSM-5 (0.5 wt%) and a Pt-Hybrid catalyst (a physically mixed catalyst with a weight ratio of Pt/SiO: (2.5 wt%) : H-ZSM-5 = 1:4) to adsorbed piperidine in the presence of gaseous hydrogen at around 473 K, whereas no such phenomena was observed on either H-ZSM-5 or Pt/SiO2126]. Y. Fan et al. reported that this hydrogenation rate was dependent not only on the nature of supported metals in hybrid catalyst system but also on the acidic strength of zeolite catalysts [27]. This paper deals with the hydrogen spillover and its role in the catalytic hydroconversion.
2.
SPILLOVER
AND
REVERSE
SPILLOVER
In the. early period of spillover research, the main experimental method for detecting
39 spillover phenomenon is the adsorption technique, where much more amount of hydrogen, which should be adsorbed on supported metal, was adsorbed on metal-supported catalyst. Especially, for the carbon-supported system, the adsorption and the temperature programmed desorption technique revealed very clearly the reversible adsorption of hydrogen on active carbon through supported metal or metal sulfide and its participation ~ in the catalytic dehydrogenation[35]. Also, the acceptor site of slSilt-over hydrogen has been concluded to be the hydrogen-unsaturated carbon (free radical). However, this method has not successively applied to solid acid such as zeolite, because of extremely small ~tmount of acceptor site. Other clear phenomenon about hydrogen spillover is the reduction of metal oxide at much lower temperature than metal free system in either metal-supported case or physically mixed metal-oxide system. This phenomenon is strongly related to the synergistic effect of cobalt or nickel for Co-Mo or Ni-Mo HDS catalyst[28]. As it has been clearly demonstrated by Delmon et al., hydrogen which is spilt over at Co or Ni-site migrate from the site to Mo st~lfide site to react with sulfide ion and remove it as H2S to increase the sulfur deficiency on Mo, which is the active site of HDS reactionr This pl'fenomenon happens even when Co or Ni site is separated from Mo site by support[36]. This concept also claimed for Pt-Mo system [29]. One of the most important actions of spilt-over hydrogen is the generation and/or interaction with acid site. It has been reported that the protonic acid sites were generated and the Lewis acid sites were weakened on Pt/SO42-ZrO2 by heating in the presence of molecular hydrogen. It means that molecule on Pt, spillover of the H atom onto SOaZ-ZrO2 surface. It is suggested that the shift of the S=O stretching band to a lower frequency by heating in the presence of hydrogen is caused by the electron transfer from the spilt-over hydrogen atom to the Lewis acid sites[18]. In the case of CoMo/SiO2+silica-alumina. (physical mixture), synergy effect in the selective cracking of diphenylmethane to benzene and toluene in the presence of H2 can be interpreted by the creation of BrCnsted acid site from the spillover hydrogen [30]. Recently my group has studied the interaction between spilt-over hydrogen and acid sites on silica supported noble metal+zeolite hybrid catalysts by means of FTIR[32], especially, it was investigated the effect of hydrogen spillover on adsorbed pyridine over Pd/SiO2+H-USY hybrid catalyst. Figure l(a) shows the change of FTIR spectra when hydrogen gas was introduced to the sample with pyridine adsorption. The amount of pyridine left on B (1540 cm 1) and L (1450 cm "1) sites were 85% and 50% to saturated amount at 423 K. At initial stage of hydrogen introduction, L peak decreased and B peak increased. After B peak reached the saturation level, the peaks assigned to piperidine appeared. The results indicated that when vacant B sites exist, pyridine adsorbed on L site migrated to the B site. Figure l(b) shows the spectra of the sample with very low coverage of pyridine (B:55% and L: 25%). Pyridine migration from L to B was clearly observed and pyridine hydrogenation did not proceed at all even after 150-min hydrogen flow. This indicates that the spilt-over hydrogen promotes the migration or the desorption of pyridine on L acid site, suggesting that one of spilt-over hydrogen species can be hydride (H) ion, because this species is high L basic.
40
(a)
(b) 05
O. -z
evacuation at 423 K evacuation at 423 I(
I 1550
I
I
1500 Wavenumbers , cm -~
I
1450"
I
550
1
1500 Wavenumbers ' c m -L
I '450
Fig. 1. FTIR spectra of pyridine adsorbed on Pd/SiO2 + USY hybrid catalyst. (a) Initial coverage of pyridine: B 85% L 50%, (b) B 55% L 25%. Condition of H 2 flow: 423 K, 15 ml/min, pure H 2.
a) L o w
coverage
H2 H$~H
HJ'H Hz
-USY X~
/
sio~
P~
y
H-usY
b) Fti~ coverage
Fig. 2. Model scheme of migration of pyridine adsorbed on L acid site to B acid site promoted by hydrogen spillover effect, and subsequent hydrogenation of pyridine to piperidine.
41 3. HYDROISOMERIZATION HYDROCARBONS
3.1
AND
HYDROCRACKING
OF
PARAFFINIC
Hydroisomerization on supported and hybrid catalyst
It is well known that platinum supported acidic solid catalysts are effective catalysts under hydrogen atmosphere for the paraffin isomerization and that the role of supported platinum and acidic solid are: (1) dehydrogenation of n-paraffins to n-olefins on platinum, (2) isomerization of linear olefins to branched olefins on acid site and (3) its hydrogenation to iso-paraffins. Other interpretation of the platinum role is that the short distance between acid site and platinum, which hydrogenate olefins, is essential for the selective isomerization [33]. In Table 1 it is seen that the hybrid catalyst composed of physically mixed and pressed fine powders of Pt/SiO2 or Pd/SiO2 and H-ZSM-5 show equivalent catalytic activity for n-pentane isomerization to that of Pt-suppQrted H-ZSM-5 in spite of that Pt/SiO2 or Pd/SiO2 shows negligible activity for either dehyhrogenation or isomerization under conditions sdopted. Also, it should be noted that the hybrid catalyst composed of mixed granules(Dp-lmm) of Pt/SiO z and H-ZSM-5 showed poor isomerization activity [34].
Table 1. Isomerization of n-pentane on ZSM-5 [34]. Catalyst Conversion/%
H-ZSM-5
Pt/ZSM-5
Pt/SiO2 a) Pt]SiO 2 Pt/SiO 2 Pd/SiO 2 +H-ZSM-5 b) +H-ZSM-5 c) +H-ZSM-5 d)
7.1
77.2
0.4
68.5
8.1
59.6
C1-C 4
65.7
5.8
0.0
0.1
18.2
0.2
i-C5 C6+ aliphatics
15.2
93.9
49.6
99.4
75.2
99.2
10.1
0.3
50.4
0.5
6.6
0.6
C6§ aromatics
9.0
0.0
0.0
0.0
0.0
0.0
Selectivity /C-mol%
Reaction condition: 423 K, n-Cs:H2=0.1MPa:0.9 MPa, W/F=10 gh/mol. a) Pt (2.5 wt%)/SiO2 0.2 g, b) powdery mixture, Pt (2.5 wt%)/SiO2: H-ZSM-5=I:4, c) granular mixture (0.3 ram), Pt (2.5 wt%)/SiO2: H-ZSM-5=I:4, d) powdery mixture, Pd(2.5 wt%)/SiO2: H-ZSM-5=I:4. If the conventional working hypotheses are correct the hybrid catalyst containing Pd]SiO 2 should show no isomerization activity because Pd/SiO2 has no dehydrogenation activity under these conditions. Also the granular mixture of Pt/SiOz-H-ZSM-5 should show equivalent isomerization activity because the normal olefins formed on Pt/SiO2 can move quickly to the acid site on H-ZSM-5 through gas phase as shown in Fig. 3(a). However, experimental results
42 are not consistent with the expectations. If the hydrogen which has spilt over from Pt/SiO 2 or Pd/SiO2 to acid site act as an acid catalyst for the paraffin isomerization (Fig. 3(b)), all experimental results can be explained quite reasonably.
(a)
Bi-functionai Model
(b) Spiilover Model
Powdery Mixture n-C 5 .
n-C5 =
i-C5 =
/ X/
H2
I-C 5
It
/
No significant difference
§
Granular mixture n-C5 =
-,,/-,,,/
i-C5 =
It
Significant difference
t
n-C 5
Powdery M'lxture _ . .H 2'
Granular mixture
i-C 5
H2 _ PH t~
~t
n-cs, i-cs,
H2
It
Figure 3. Interpretation of catalytic activity on powdary mixture and granular mixture by bifunctional model (a) and spillover model. For example, the excellent activity of Pd/SiO 2 hybrid catalyst even with its negative dehydrogenation ability is quite understandable if it is assumed that the Pd site is the entrance of hydrogen from gas phase to acid site. Also, the poor catalytic activity of Pt/SiOz-H-ZSM-5 granular mixture could be attributed to that, the hydrogen species spilt over from Pt to SiO 2 surface has seldom chance to transfer to zeolite. The experiment of
(Hz+n-Cs)-+(Nz+n-Cs)--~(Hz+n-Cs)which
means the switching of
atmosphere from hydrogen--'nitrogen-*hydrogen was designed and conducted with the two catalysts Pt/HZSM-5 and Pt-hybrid. Figure 4a, and 4b show the results. Under the atmosphere of H2, the conversion and i-pentane selectivity is kept high and stable both for Pt/HZSM-5 and Pt-hybrid, however, when gas flow stream was switched from H 2 to N2, the conversion on Pt/HZSM-5 increased dramatically, almost up to 100%, then deceased gradually. For Pthybrid catalyst, no such enhanced conversion was observed. In contract with the response in conversion, the selectivity lowered rapidly for both cases. The isomerization selectivity values
43 dropped from 98% to 10% just in 30 sec. At this stage, cracking reaction took place, bringing about the production of a large amount of C3, C4 paraffins. The reaction in N2 atmosphere was continued for about 1 h and then the gas flow was switched again from N 2 to H 2, the response of selectivity and activity was described as: the isomerization selectivities were restored rapidly for both catalysts whereas the recovery of conversion for the two c~ttalysts was quite different. The activity of Pt/HZSM-5 could be restored to about 80% of the initial level under Hz but the recovery rate was slower than that of selectivity, for Pt-hybrid, on the other hand, the activity did not recover anymore. Even after the carbon was completely~ removed by air calcination, the activity was not restored. However, if the deactivated catalyst was shaped again (including grinding and pressuremolding), the conversion and selectivity could be restored completely. The amounts of carbon deposited at different reaction stages were measured using the method mentioned above. The vertical lines in the same figures represent the results. Almost no carbon depositing was 100
0o
1
artier
80
-
80
~~2+n-C5
-
re2shapmg
-.d
o60
~: 4O 0 D
nv.%" -
20
~ i40
,l '
Cony. %
1-12 +n'cs
(") 20 A C%
0
50
100
150
200
Time on stream (min) 250~
State
0
50
00
150
0 200
0 0
A 5O
Time on stream (min)
PH2=0.9Mba, PnCs=0.1Mpa.
fresh
deactivated in N., "
after removal of deposited surface carbon
Fig. 4 Effect of atmospheric gas on activity selectivity, coke and catalyst model.
after re-shaped
44 detected in H2, but as long
as
N 2 was introduced, e/,en if only lmin, carbon depositing had
began and came up to 4.25% for Pt/HZSM-5 and 3.42% for Pt-hybrid in one hour. After gas flow was switched from removed.
N 2 to H 2
again, most of the surface hydrocarbon deposit could be
The model of different behavior of the hybrid catalyst is showr~ as in Fig. 4.The
deposited coke, which is formed during the reaction under N 2 should separate the particle of zeolite from Pt/SiO2. Therefore, even if the coke is removed by treating with either split-over hydrogen or air.
The lost contact can never restored and it is restored only when it is pressed
again. The newly postulated reaction mechanism of paraffin hydroisomerization isshown in Fig.5, which involves (1) the dissociation of gaseous hydrogen on metal site to atomic hydrogen and spillover to acid surface as H § and H-, (2) the activation of paraffinic hydrocarbon by H § to carbenium ion on acid site to result in the isomerization to branched carbenium ion and (3) the reaction of isomerized carbenium ion with-H to make isomerized paraffin. This explanation coincides with the fact no olefins are detected in the gas phase on hydroisomerization system even the catalyst is Pt- or Pd containing hybrid catalyst system and that even the hybrid catalyst which contain Pd/SiO 2 shows comparable activity.
+Hso
+
+H+so.-H
[3-scission 1l
+.-so
.•
-sc, ss,on
+H-so ~ ~+H+so, .--I-t2 7
+H'so
.
1l
11
[3-scission fast
(3~
/ •+
~
I+H2/Pd
-.yt
l+H-so
+H+so, -H2
NO13-scission
Fig. 5 Reaction model of n-pentane hydrocracking over Pd/SiO2-DAM hybrid catalyst
3.2 Hydrocracking of n-Heptane on Pd containing mordenite Hydrocracking of normal paraffins on metal-supported zeolites, which only
C3"vC5
45 paraffins as primary products was studied. Fig. 6(a) shows the changes of catalytic activities of a veriety of catalysts containing Pd/SiO2 and/or H-M(mordenite) for n-heptane hydrocracking.
Pd/H-M, which is a typical
dual functional catalyst, showed excellent activity for the hydrocracking. activity of H-M was not affected by the atmosphere and decreased quickly. little activity for both dehydrogenation and cracking of n-C7H16.
80 70
On the other hand, the
o Pd-hybrid H2~)
a
9 Pd-hybrid N21)
~' 50 o
The catalytic Pd/SiOz showed
[5 DAM H22)
40 30
I
20
10
DAM N22)
A Pd/SiO 2 H22)
o
0
0.5 1 1.5 2 2.5 Time on Stream (h)
3
O Pd/DAM H22)
563 K, H2/n-CT--9, 1.1 MPa, 11 Pd/SiO2:DAM=I:I, W/F=2.4 g h mol-1, tool-1
2)
W/F=I.2 g h
Fig. 6 Hydrocracking of n-neptane over Pd-DAM hybrid catalyst; a) conversion as a function of time on stream, b) C-number (TOS=2.5 h), c) distribution of C4 hydrocarbons formed in the hydrocracking of n-C7 (TOS=2.5 h). catalytic activity of a hybrid catalyst comprising Pd/SiO2 and H-M was the highest and its activity was kept constant under hydrogen atmosphere while it was much lower and decreased
46 quickly under nitrogen atmosphere.
This phenomenon clearly shows that the presence of
hydrogen is essential to generate hydrocracking activity. It is well known that the supported platinum shows a high catalytic activity for the dehydrogenation of paraffin whereas the supported palladium does not.
The results shown
in Fig. 6(a) suggest that the dehydrogenation activity of supported metal is not essential for the appearance of the paraffin hydrocracking activity, but the hydrogen migration from Pd/SiOz or supported palladium to acid site should be essential for the high and stable catalytic activity. The characteristic feature of the product distribution is that the reaction products of Pdhybrid catalyst system are only isomerized heptane and propane and equimolar amount of isobutane (small amount of n-C4Hlo was formed), whereas the products ,on ,H-M alone distributed from C3 to C9 and Ca products contained all kind of paraffins and olefins as shown in Fig. 6(b).
The wide product distribution for H-M system should be attributed to the
reaction path comprising oligomerization of cracked fragments and cracking of the oligomers. In hydrocracking of normal paraffin with metal supported acid catalyst, the iso/normal ratios in the paraffinic products generally exceed the thermodynamic equilibrium(Fig. 6(c)). It proves that at least some of the branched paraffins are primary products of the cracking and notthe result of the post isomerization.
This is particularly true in the case of C4, since n-
butane cannot be isomerized under typical hydrocracking conditions with a zeolite catalyst. One probable path is the skeletural isomerization of n-paraffin to branched paraffin and its cracking.
Fig. 7 shows the results of hydrocracking kinetics of C7 isomers.
2-
dimethylhexane was more reactive than n-neptane was less reactive than the other branched isomers.
The reactivities of the C7 isomers can be explained by stability of the
corresponding carbenium ion.
These facts indicate that n-neptane is isomerized to 2-
methylhexane and then it is further isomerized to 2,4-dimethylpentane and both of then were cracked to give propane and iso-butane.
It has been suggested that the formation of
multibranched isomers from the feed and cracking are consecutive reactions [4].
Cracking
of a normal paraffin must thus proceed through the stage of formation of monobranched isomers such as 2-methylhexane, dibranched isomers such as 2,4-dimethylpentane and finally cracked.
In hygrocracking, reaction path which include hydrogen are shown in Fig. 5.
Both proton and hydride ion should be participated in either the activation of hydrocarbon molecule or stablization of acrbenium ion. Hydride ion reacts with the cationic-cracked products on zeolite while and olefinic-cracked products are hydrogenated on palladium to be converted into the less reactive smaller paraffins. Isobutylene, which is one of a pair of the primary cracked products of the 2-methylhezane or 2,4-dimethylpentane hydrocracking, will be hydrogenated to isobutane over palladium catalyst in the presence of hydrogen.
47
n-heptane 1t30 90 ~, 80
2-methylhexane
100 90 ~..,--80
"~ 6o o 50 9 40 o 30 20 10 0
_
7o
"~ 60 50 40
Gr
>
3o
0
0.5
1
1.5
2
W/F(g h tool -l )
2.5
r,.) 20 1o o
0
t 0.5
t 1
I 1.5
i 2
2.5
W/F(g h rnol1 )
[] Conversion(%}' 9 Sel. of Cracking(%) A Sel. of Isomerization(%) 523 K, H2/n-C7---9, I.I MPa, Pd/SiOz:DAM=I'I, Fig. 7 Hydrocracking of C7 isomers over Pd-DAM hybrid catalyst. 3.3'
Effect of hydrogenation
activity on isomerization
and/or cracking
Fig. 8 shows the effect of catalyst composition on the activity and selectivity of the hydroconversion of n-hcxane on Pd/SiOz - H-M hybrid catalyst. As apparently from the data in Fig 8, Pd/SiO2 shows no activity and H-M shows quite low activity. However, their physical mixture shows much higher activity. With increasing Pd/SiO2 content the selectivity of hydroisomerization is increased, while the hydrocracking selectivity increases with increased mordenite content. In terms of TOF (based on acid site), cracking reaction is almost independent on the catalyst composition, whereas that of isomerization reaction is markedly accelerated by increasing Pd/SiO 2 content. These facts suggest that the increased amount of hydrogen promotes the isomerization reaction, while it does not promote the cracking. This concept is also supported by the effect of hydrogen pressure on the hydroconversion on the Pd/SiO2-hybrid catalyst as shown in Fig. 9. In case of the catalyst with the composition of Pd/SiO2 to mordenite ratio is 1:4, the cracking rate increases quickly with the increase in hydrogen pressure to reach maximum at about 1 MPa and then decreases while the isomcrization rate increases mononously up to 2.5 MPa. These phenomena also happens with thc catalyst of highcr Pd/SiO2 contcnt (Pd/SiO2" moredcnite is 3:1), but the maximum rates appcar at lower hydrogen pressure, probably because of higher hydrogen supply.
They arc
explained reasonably by the reaction model which include hydrogen spillover (H § and H) shown in Fig. 5. At the initial stage of the reaction, n-paraffin is first activated to secondary carbcnium ion and then isomcrized to branched carbcnium ion (tertiary carbenium ion).
If
this carbcnium ion reacts with H to form stable paraffin, the reaction is isomcrization.
48 However, the carbenium ion is not stabilized by H , i t reacts further along the O-scission to make cracking product. Thus, hydrogen species (H) on the acid site promote isomerization. It is clear that split-over hydrogen is essential for the activation of paraffins and control the acid-catalyzed reaction. 4. CONCLUSION It is darified that hydrogen spillover onto zeolite for either B-acid site and L-acid site and that the role or noble metal and gaseous hydrogen in the hydroconversion of paraffinic hydrocarbones on metal-supported solid acid.
The similar model should be applied for the
practical hydroconversion catalysts. 280~C, 1MPa, W/F=1 ~40 = 30 O
20 O
r,..) lO . , . . . . . l. . . . . . . . .
i. . . . . . . . .
I. . . . . . . . .
I. . . . . . . .
'
80
80
7O
>., 60
60
.>_
o
40
~-
20
50
lng
4O O
120
I
I
I
I
0 100
kr~
0
[.-,
o
30
9 ~. ....
.~
20 10
60
lsomerlzauon
,,u ~.
""
~---~.___._____~ 1 2 Hydrogen pressure/MPa
_
20
9 Pd/SiO2:Mordenite=l:3 20 PdySiO 2
40
60
80
100
Mordenite content (%) Mor.
Fig. 8 Hydrocracking of n-hexane over mordenite hybrid catalysts
9 Pd/SiO2: Mordenite =4:1 Fig. 9 Effect of hydrogen pressure on hydroisomerization and hydrocracking of n-decane
49 5. R E F E R E N C E S
1 K. Fujimoto, K. Maeda and K. Aimoto, Appl. Catal. A General, 91 (1992) 81. 2 I. Nakamura, K. Sunada and K. Fujimoto, Stud. Surf. Sci. Catal., 105 (1997) 1005. 3 I. Nakamura, K. Sunada and K. Fujimoto, Stud. Surf. Sci. Catal., 106 (1,997~ 361. 4 A. Zhang, I. Nakamura and K. Fujimoto, Stud. Surf. Sci. Catal., 106 (1997) 561. 5 A. Zhang, I. Nakamura, K. Aimoto and K. Fujimoto, Ind. Eng. Chem. Res., 34 (1995) 1074. 6 T. Shishido and H. Hattori, J. Catal., 161 (1996) 194. 7 T. Hosoi, T. Shimadzu, S. Ito, S. Baba, H. Takaoka, T. Imai and N. Yokoyama, Prepr. Syrup. Div. Petr. Chem., Am. Chem. Soc., 562 (1988) in: Successful Design of Catalysts, p. 99, Elsevier, Amsterdam, 1988. 8 S. Ohgoshi, I. Nakamura and Y. Wakushima, Stud. Surf. Sci. Catal., 77 (1993) 289. 9 A. Zhang, I. Nakamura and K. Fujimoto, Stud. Surf. Sci. Catal., 112 (1995) 3~91. 10 J. Kuriacose, Ind. J. Chem. 5 (1957) 646. 11 J. Khoobiar, J. Phys. Chem.,,,68 (1964) 411. 12 B. Delmon, Pouchot, Bull. Soc. Chim., 2677 (1966). 13 J.L. Carter, E J. Lucchesi, P. Corneil, D. J. C. Yates, J. H. Sinfelt, J. Phys. Chem., 69 (1965) 3070. 14 K. Fujimoto, S. Toyoshi, Proceeding of 7th International Congress on Catalysis, (1980) 235. 15 W.C. Conner, G. M. Pajonk and S. J. Teichner, Adv. Catal., 34 (1986) 1. 16 U. Roland, T. Braunshweig, E Roessner, J. Mol. Catal. A: Chemical, 127 (97) 61 17 K. Ebitani, H. Konishi and H. Hattori, J. Catal., 130 (1991) 257. 18 K. Ebitani, H. Konno, H. Konishi and H. Hattori, J. Catal., 135 (1992) 60. 19 H. Hattori, T. Shishido, J. Tsuji, T. Nagase and H. Kita, in: Science and Technology in Catalysis, (1994) 93. 20 K. Fujimoto, M. Adachi and H. Tominaga, Chem. Lett., (1985) 547. 21 U. Roland, H. Winkler, H. Bauch and K. H. Steinberg, J. Chem. Soc., Faraday Trans., 87(1991)3921. 22 U. Roland, R. Salzer and L. Sumrnchen, Stud. Surf. Sci. Catal., 97 (1995) 459. 23 T.R. Hughes and H. M. White, J. Phys. Chem., 71 (1967) 2192. 24 R E. Eberly, J. Phys. Chem., 72 (1968) 1042. 25 J.C. Vedrine, A. Aurox and V. Bolis, J. Catal., 59 (1979) 248. 26 Y. Fan, I. Nakamura and K. Fujimoto, Stud. Surf. Sci. Catal., 112 (1995) 319. 27 A. Zhang, I. Nakamura and K. Fujimoto, J. Catal., 168 (1997) 328. 28 B. Delmon, React. Kinet. Catal. Lett., 13 (1980) 203. 29 P.A. Sermon, K. M. Keryou, Stud. Surf. Sci. Catal., 112 (1997) 251. 30 A.M. Stumbo, P. Grange, B. Delmon, Stud. Surf. Sci. Catal., 112 (1997) 211. 31 E Schuetze, E Roessner, J. Meusinger, H. Papp, Stud. Surf. Sci. Catal., 112 (1997) 127. 32 R.Ueda, K.Tomishige and K. Fujimoto, Catal. Lett., 57 (1999) 145-149. 33 H.Y. Chu, M. P. Rosynek, J. H. Lunsford, J. Catal., 178 (1998) 352. 34 K. Fujimoto, K. Maeda, K. Aimoto, Appl. Catal., A 91 (1992) 81. 35 K. Fujimoto, S. Toyoshi, Proceeding of 7th International Congress on Catalysis, (1980) 235. 36 M. Karroua, H. Matralis, E Grange, B. Delmon, J. Catal., 139 (1993), 371.
This Page Intentionally Left Blank
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
Conversion of model sulfur compounds hydrodesulfurization CoMo/Ai203 catalysts
51
to
characterize
J. Leglise, L. Finot, J.N.M. van Gestel and J.C. Duchet Catalyse et Spectrochimie- UMR.CNRS.6506, ISMRA, Universit6de Caen, 6 Bd du Mar6chal Juin, 14050 Caen Cedex, France - Fax: (+33) 2 3145 2822 -
[email protected]
Abstract A series of commercial CoMo/AI203 catalysts has been tested for thiophene conversion under various operating conditions (300-400~ 1-30 bars). Change in activity and selectivity was monitored by varying independently the H2S content (0-0.9 bar). A kinetic model explains the complex effects brought by H2S. The model allows determining the contributing rates of catalytic sites and superficial active species. Thiophene results were compared to DBT conversion and benzene hydrogenation data. Under the best conditions, thiophene HDS correlates real feed HDS. Requirements for improving the HDS performances of sulfided catalysts are proposed. 1. INTRODUCTION
Tighter emissions regulations entail the development of catalytic hydrotreatment [1]. Within EEC, the sulfur requirement in diesel fuel is 500 wppm now and will target 50 wppm in the year 2005. New catalysts are thus rapidly marketed, and consequently the need to assess their hydrodesulfurization activity is increasingly growing. Among all available techniques, microactivity tests should be more pertaining to predict the industrial HDS activity. In spite of this, literature about the comparison between model sulfur compounds and gas oil HDS is scarce [2-5]. The question is what molecule and reaction conditions should be chosen. To overcome this, several authors have proposed a mixture of model molecules to comply with the complexity of the feeds [6-9]. None of these mixtures has been utilized further. In fact, many authors preferred to use a single compound to test various catalyst preparations [10]. We previously showed that thiophene HDS enabled to model real feed HDS [5]. In the present study, we will discuss the applicability of the sulfur molecule test depending on reaction conditions. The contribution of catalytic sites and superficial H species to global HDS will be determined from kinetics. The objective is to gain some insight on catalyst functions and properties that are required for improving hydrodesulfurization. The study is restrained to CoMo/AI203 catalysts. 2. CATALYSTS AND HYDRODESULFURIZATION ACTIVITY The oxidic precursors are fresh and regenerated commercial CoMo/AI203. They differ in metal composition and in shape. The oxidic solids were crushed and sieved (0.2-0.5 mm) before catalytic testing. Thiophene (T) and dibenzothiophene (DBT) were used to model the
52 HDS properties of the catalysts. The reaction of thiophene was carded out in two lab-scale flow reactors operated under either differential (overall pressure 1-1.5 bar) or integral (30 bars) conditions. The partial pressure of H2S was varied independently of the thiophene pressure. The reaction of DBT was studied in a batch reactor. Benzene hydrogenation was measured in an integral flow reactor. Tables 1 and 2 summarized the conditions of sulfidation and catalytic measurement. Experimental details are reported elsewhere [5, 11]. Table 1 Gas phase sulfidation Model molecule Thiophene Dibenzothiophene Benzene
Sulfiding molecule
Pressure (bar) H2S Overall
Temperature (oc)
H2S H2S H2S
0.15 0.15; 3.6 0.15
1 30 1
400 400 400
CH3SSCH3
0.02
80
400
Overall
Temperature (~
Table 2 Reaction conditions Reactant
Pressure (bar) Added H2S
0.08 0.08-1.30 (0.094) a
0-0.5 0-0.85 none
1-1.5 30 40
300-400 300-380 335
0.70
0.02
80
300-380
Model molecule Thiophene Dibenzothiophene Benzene a
Liquid phase, concentration in mol L-1
The catalytic activities for HDS of several light and heavy gas oils were collected on bench-scale units. The sulfur content of the feedstocks ranged from 0.8 to 3.0 wt.%. From a practical standpoint, the HDS activity is expressed by the rate constant, n-1
S n-1
S~) -1
With n between 1.5 and 1.9, LHSV the liquid space velocity, and S and So the weight concentration of sulfur in the final product and feedstock. Four catalysts were compared using their relative-volume activity (RVA). The RVA values are subject to the feed and operating conditions. Hence, absolute RVA values are not accessible. However, deviation on RVA did not exceed 10% of the averaged value. For example, the activity ratio k2/k~ of two catalysts was equal to 1.18 with a SRGO and 1.20 with a VGO. Averaged RVA (100-180) were then compared to the relative rates determined for model molecule HDS and hydrogenation.
53 3. CHOICE OF THE M O D E L S U L F U R M O L E C U L E
In the desulfurized oils, thiophenes are all virtually absent and benzothiophenes mostly converted [10, 12]. Therefore, DBTs should be preferred as model molecules. In the recent years, the very refractory 4,6-dimethyldibenzothiophene (4,6-DMDBT) has been often used as a model for deep HDS [4, 9]. Before selecting a sulfur compound, it is necessary to examine some basic properties. The polarized C-S bond is shorter in thiophene (1.71 A) than in DBT (1.74 A). Hence, thiophene is intrinsically more refractory. However, in presence of a catalyst, acting as an electron withdrawing or donating material, the reactant is destabilized upon coordination. Much of the surface chemistry can thus be understood in terms of reaction steps involving electrophilic or nucleophilic displacements. Electronic charges of the S and C atoms of the heterodng for thiophene (T), benzothiophene (BT), and DBTs, were thus calculated using the semi-empirical Hamiltonian PM3 (Table3). We included also the congested tetraphenylthiophene (TPhT) for comparison. From charge densities, the reactivity order should fall in the sequence: Electrophilic attack: T > BT > TPhT > DBT = 4,6-DMDBT Nucleophilic attack: T = TPhT > BT > DBT - 4,6-DMDBT In all cases, thiophene appears the most reactive. On the other hand, thiophene would be the less adsorbed compound. Indeed, after Zdrazil [13], the adsorption capacity would increase exponentially with boiling temperature (Table 3). Therefore, inhibition should increase as follows: Adsorption: T << BT < DBT < 4,6-DMDBT << TPhT Conclusively, the known reactivity order of thiophene, BT, and the two DBTs [10, 14, 15] is simply explained by the ratio of attack to adsorption terms. However, TPhT, which is the most hindered compound and has the highest boiling temperature, should not be easier to desulfurize than DBT as Singhal et al. reported [16]. This suggests that electronic effects are more important in hydrodesulfudzalJon than steric effects. Consequently, adsorption should be minimized in order to evaluate better the catalyst chemical functions. In that respect, thiophene may be the preferred model molecule. Table 3 Charge density on the thiophenic ring and boiling temperature of various sulfur molecules. Product Charge density S C~
cI3 Boiling Temp. (~ a
T
BT
DBT
4,6-DMDBT
TPhT
+ 0.304 - 0.300
+ 0.263 - 0.210 (C7a) - 0.267 (C2) - 0.070 (C3a) - 0.098 (C3)
+ 0.244 - 0.191
+ 0.242 - 0.191
+ 0.308 - 0.203
- 0.043
- 0.039
- 0.035
332
371 a
760 a
- 0.122 84
221
Calculated using the Jobak's method [17]
54 Moreover, the reaction networks of thiophene (Fig. 1) and DBTs HDS are remarkably similar [9, 15, 18, 19]. The choice of the sulfur molecule may thus not matter. In all cases, many chemical bonds must be broken to eliminate sulfur from the heteroring. With thiophene, dihydrothiophene and butanethiol intermediates have been reported [20, 21]. Butadiene is detected often at low H2 pressure, whereas thiolane (THT) and butenes are found at high pressure [10]. Hence, depending on reaction conditions, the conversion proceeds through partial or complete hydrogenation of the ring before sulfur removal, so called "hydrogenolysis" or "direct desulfurization" and "hydrogenation" pathways [10, 22]. In the case of DBT, hydrogenated intermediates have been reported, and biphenyl and cyclohexylbenzene formation indicate the pathway follows by desulfurization [15]. At low H2 pressure, the conversion is limited by the dearomatization step [22, 23]. At high hydrogen pressure, the opening of hydrogenated intermediates or the subsequent elimination of SH groups may be limiting.
~+H2
~
+H2 ~
~
<---- Hydrogenation
__},.. +H2 ~-H2S
4 ~ C-Scleavage +H2~-H2S <-----H2Selimination
Figure 1. Reaction network of thiophene hydrodesulfurization In conclusion, we selected thiophene because the inhibitive adsorption should be the lowest among the aromatic sulfur compounds. DBT was also selected because its charge densities are identical to those of 4,6-DMDBT and would be less adsorbed. 4. INFLUENCE OF H2S AS A COMPETITOR The conversion of sulfur compounds produces H2S, which can amount up to 3 bars at the exit of the hydrotreating reactor [24]. Hence, the catalyst surface and the catalytic functions are highly modified along the reactor length. It is therefore critical to elucidate how H2S effects the reaction pathways in relation with the modifications of the sulfide catalyst. Although knowledge about the influence of H2S on real feeds still awaits systematic investigation, a typical example on VGO HDS is shown in figure 2a [25]. H2S strongly inhibits the HDS rate even with such a heavy feed, only one fourth of the activity remains at 4 bars of H2S. Such global response can thus be used to fingerprint the influence of H2S. Therefore, an appropriate model molecule must be affected in a similar manner. Examples on DBT and 4,6-DMDBT HDS (Fig. 2b) are taken from the work of Kabe et al. [26]. At first sight, the variation observed with both compounds does not match that of the VGO. The inhibiting effect was present, but was here too pronounced. Moreover, DBT appeared completely inhibited by H2S. Strikingly, some HDS activity remained using
55 4,6-DMDBT. This thioresistance towards H2S seems a characteristic of sulfide catalysts since it stands also for lighter molecules such as thiophene, toluene, and dimethylaniline [27-28]. That points to the existence of different catalytic sites that operate at high H2S levels [28]. 100
100 T
b- Model molecules
75~
75 L_
K) >
>= 50
50
4,6-DMDBT
I
25 0I 0
i
I
i
0 0.0 0.2 0.4 0.6 0.8 1.0
L
1 2 3 4 5 H2S Pressure (bar)
H2S Pressure (bar)
Figure 2. Influence of H2S on HDS rate over CoMo/AI203: a- LePage et al. [25], VGO (El), 370~ 35 bars; b- Kabe et al. [26], DBT (0) and 4,6-DMDBT (e), 350~ 30 bars. A typical influence of H2S on thiophene conversion measured under pressure is shown in figure 3a. Inhibition and thioresistance effects are both observed; and the variation appears better related to that of the VGO. Hence, thiophene seems more appropriate than the two heavier DBT compounds.
E o
1.0~
-oo |>9 0
5O
a .
8
~
.
(~ 0"4 I E '-'4 ._o o 0.2 LL 0.0 i:3::~=ri i l 0.0 0.2 0.4 0.6 0.8 1.0 H2S pressure (bar) L_
b
~..o~ 40 "7 _-r 30 o E 20
5O
40 30
v
(D
n" 10 0 0.0 0.1 0.2 0.3 0.4 0.5
H2S pressure (bar)
Figure 3. Influence of H2S on thiophene reaction (PT = 0.08 bar) a- fractional conversion over fresh CoMo/AI203, 340~ 30 bars; b- HDS rate over regenerated CoMo/AI203, 400~ (0) and 350~ (c]), 1 bar. On the other hand, thiolane was found to accumulate when H2S was increased (Fig. 3a). Such result was already reported at lower temperature with a higher selectivity into THT [29]. Therefore, as a secondary effect, H2S enhances the hydrogenation function of the catalyst compared to its C-S breakage capacity.
56 When the hydrogen pressure was lowered, the general response to H2S, here at 350~ was again similar with both inhibition and thioresistance behaviors (Fig. 3b). The global HDS rate decreased by a factor conforming to the H2 pressure ratio showing that the reaction is nearly first-order with respect to H2. That indicates that H2 and H2S are not stricto sensus competing on the same sites, and secondly that the adsorption of H2 is weak. At the higher temperature of 400~ (Fig. 3b), the effect of H2S appeared more complex. Inhibition was lower than observed at 350~ what is expected if H2S adsorbs competitively with thiophene. The thioresistance behavior was also found at high H2S pressure. Besides, small amounts of H2S slightly accelerated the HDS rate. The effect was found reversible, so that a destructive reduction of the metal sulfide in absence of H2S could be discarded. Such finding is against the Le Chatelier's principle since the product H2S is able to accelerate the conversion. Prove is thus established that species issued from H2S participate in the reaction mechanism. This unusual behavior was found with every investigated commercial CoMo/AI203 catalyst. In summary, H2S leads to a variety of effects that must be assessed independently if one wants to determine their consequence on the active surface of the sulfide catalyst. Excepting the accelerating effect, the others were noticed with the heavier molecules, DBT and 4,6-DMDBT, as well as with the VGO. In addition, H2S appears to favor hydrogenation reactions. The effects on thiophene HDS are recalled below. - Acceleration at low H2S content ~ superficial active species from H2S - Inhibition ~ H2S and reactant adsorbed on the same site - Stabilization at high H2S content ~ modification of catalytic sites - Similarity whatever the H2 pressure ~ H2 and H2S are not competing on the same site. 5. ELEMENTARY
STEPS
AND
KINETICS
Firstly, catalytic sites and active species were postulated taking into account the above features. Many mechanisms for thiophene HDS have been proposed, namely considering a one-point adsorption through the S atom or a multi-point adsorption by ring ~ electrons [10]. The nature of the CUS site is pertinent to the binding mode of the reactant [30]. For instance, TI1-S coordination requires only one vacancy and TIET three vacancies. ~15T should be the main mode of coordination [31]. It may reflect what occurs at low H2S content. Several binding modes however may occur simultaneously during the reaction because of the competitive adsorption of thiophene and H2S. Therefore, several catalytic sites do probably coexist. Many works are in support of such hypothesis [10]. The superficial active species are thought to come from the dissociation of both H2 and H2S. Hydrides, protons, and S-H groups have been evidenced by means of various spectroscopic or chemical techniques [32-35]. H2 H2S
--
~ Ha* + H~- Ha§ + HS~-
(2) (3)
Without H2S, the concentrations of Ha* and H8 species are equal. They both decrease in presence of H2S because of adsorption blocking part of the dissociating sites. However, the concentration of H~* becomes higher than that of H8 species. This may explain why H2S
57 favors hydrogenation relatively to C-heteroatom breakage [29]. It was thus inferred that hydrogenation requires electrophilic H a* species. However, Kasztelan and Guillaume [27] considered that hydrogenation is limited by the addition of hydride H- species. Therefore, we purposely follow van Parijs and Froment [36] who could not discriminate on kinetic grounds between molecular and atomic hydrogen adsorption. The rate equations are based on the LHHW formalism. We assume that thiophene and H2S adsorb on a vacancy site or, and H2 on a separate site 13. H2 + 13 T + o~ H2S + ~
-----
"~ (H-H).I3 "~ T.~ "~ (H-SH).~
with bH = [H2.13]/PH[13] with b~,T = [m.o~]/PT[~] with bs = [H2S.cz]/Ps[~]
(4) (5) (6)
The o~sites are active for hydrogenation and C-S breakage depending whether H ~* or H ~ species are near the adsorbed reactant. Attack by species issued from H2 gives, T.~ + (H-H).I3 = TH.ot + H.IB ~ TH2.~ + 13 TH2.ot + (H-H).I3 ~ = TH4.ot + 13 The hydrogenated products desorb liberating the o~site. TH2.o~ TH4.(~
TH2 + o~ TH4 + o~
----...-
with br = PTH2H/[T.ot] with b~,TH4= PTH4H/[T.~]
(rds 7)
(s)
(9) (10)
The reaction proceeds further with stepwise C-S bond breakages leaving H2S adsorbed on the same site as T. The products, butenes and butane (Bu), were not distinguished. TH2.o~ + (H-H).I3 TH4.(~ + (H-H).I3
= BuSH.(~ + H.13 = =
= H2S.ot + Bu + 13 H2S.~ + Bu + 13
(11) (12)
Similar steps are written with species issued from H2S. In this case, one S atom will remain on the ~ site. Hence, we supposed that the o~ site was freed by reacting with H2 gas (Eq. 15). T.~ + (H-SH).~ = TH.cz + HS.~ ~TH2.o~ + (H-SH).o~ ~ ~S.o~ + H2 "~ o~+ H2S
TH2.o~+ S.o~ TH4.ot + S.o~ with K = [~] Ps/[S.ot] PH
(rds 13) (14) (15)
Steps 7 and 13 are determining when dearomatization limits the conversion. The two corresponding rate equations account for the continuous decrease of the HDS rate and for the slight maximum, but not for the remaining activity at high H2S pressures. Therefore, we assume that ot site transforms into a sulfur-saturated (~ site. Then, we write a similar set of reaction steps on the (~ sites.
(H-SH).o~ -.
T+~
---
""~
(~ T.(~
with b~,T = [T.~]/PTM
(16)
(17)
58 Thiophene and products adsorb on this site. Hence, a sulfanic S-S bond may be formed. This has been proposed already by Kwart et al. [37] in their study about the mechanism of thiophene conversion, and was recently justified from theoretical calculations [38]. It can otherwise relate a congested site and may reveal the participation of the ~11-Scomplex. T.c + (H-H).I3 TH2.(~ + (H-H).I3 TH2.~ + (H-H).I3 TH4.(~ + (H-H).I3 T.~ + (H-SH).o~ TH2.(~ + (H-SH).o~
= TH.~ + H.13 ~ TH2.~ + 13 ~ = TH4.o- + 13 ~ BuSH.~ + H.13 ~ H2S.~ + Bu + 13 ~ ~ H2S.(~ + Bu + 13 = TH.~ + HS.o~ ~ TH2.~ + S.o~ ~ ~ TH4.~ + S.o~
(rds 18) (19) (20) (21) (rds 22) (23)
Steps 20 and 21 were supposed not to take place since the ~ site is already saturated with H2S. That means that C-S removal is unlikely on H2S-saturated sites, which is conformed to the experimental observation where H2S favors hydrogenation. Steps 18 and 22 are again determining if dearomatization limits the conversion. The total rate R for thiophene disappearance is thus the sum of four terms (Eq. 24), which represent the contributing rates of vacancies o~ and sulfur-saturated sites ~ with species issued from H2 and H2S. For clarity, the initial contributing rates only are given in Eqs. 25-28. (24)
R -- FV,H2 + FV,H2S + FS,H2+ FS,H2S
Vacancies - H2:
ba'TeT bHPH rv'H2= kv'H2 1 + b~,TPT+ bsPs 1+ bliPH
(25)
Vacancies - H2S:
rv,H2s = kv, H2s
(26)
bsPs 1+ b,,TPT + bsPs + Ps 1+ b,TPT + bsPs + - P s ba'mem
KP H
b~,TPT
KP H
bsP s
blipH
Sulfur-saturated sites H2:rS,H2 ks'H2 1+ b~,TP T 1 + bc~,TPT+ bsP s + Ps 1+ bliP H
(27)
KPH
boTP,
(bsP)= PS)=
Sulfur-saturated sites- H2S: rS,H2s=ksH2S ' ' I+b~'TPT(I+b~,TPT +bsPs + K p H
(28)
The model allows determining the contribution of catalytic sites and active species. The contributing rate due to vacancies is given by the sum rv,H2 + rV,H2S and that of sulfursaturated sites by rS,H2+ rS,H2S.Rates due to species issued from H2 and H2S are obtained in a same way. One can thus appraise the superficial modifications brought by a change in conversion and operating conditions. Moreover, the rates following the hydrogenation and hydrogenolysis pathways can be determined when the selectivity has been measured.
59 6. CONTRIBUTION OF CATALYTIC SITES AND ACTIVE SPECIES
We first examine the results obtained at atmospheric pressure, since the amount of added H2S exceeded that produced by the conversion of thiophene. Figure 4a shows the evolution with respect to H2S of the four contributing rates at 400~
50
,e-
4o
50 i "T,o'} 40
3O
',-- 30
2o
E 20
b- 350~
O
s,~
n~1 0 ........ 0 ~
0.0 0.1 0.2 0.3 0.4 0.5
H2S pressure (bar)
(D
rv" lO 0 L--- -'r9-" r r i-~ 0.0 0.1 0.2 0.3 0.4 0.5
H2S pressure (bar)
Figure 4. Influence of H2S on rate contributions for thiophene HDS, 400~ (a) and 350~ (b), 1 bar: vacancies and species issued from H2 (--) or H2S (. . . . ), sulfur-saturated sites and species from H2 (---). Without H2S, the HDS activity results from the sole action of vacancies with species issued from H2. The corresponding rate rv,H2 gradually decreases because of H2S adsorption. In parallel, part of H2S dissociates and yields active species resulting in the fast increase of rv,H2S,which goes through a maximum at about 0.02 bar of added H2S. Beyond that value, vacancies are progressively blocked and the rate rV,H2S decreases regularly. Moreover, H2S transforms immediately some vacancies into sulfur-saturated sites. Consequently, a significant rate rS,H2contributes to global HDS. Finally, the contributing rate rS,H2Sof sulfur-saturated sites with H2S-issued species was nil in the whole range of added H2S, supporting the fact that H2S does not dissociate on the sulfur-saturated site. Vacancies only account for the HDS activity measured with small amounts of H2S. Above 0.15 bar of H2S, the contribution of the sulfur-saturated sites prevails. The majority of active species comes from H2, H2S providing about 30% of species at the rate maximum and 20% at the highest H2S pressure. Those features were found with every CoMo catalyst. We now tum to the reaction under high H2 pressure. There, differential conditions could not be obtained, but the contributions of sites and active species to the conversion and rate were determined from kinetic modeling (Fig. 5). The contribution of H2S issued species was found again negligible. Thus, HDS conversion sums up to the contribution of vacancies and sulfur-saturated sites with species coming from H2 (Fig. 5a). Reaction on vacancies predominates in the whole range of added H2S. When the temperature was increased, the contribution of the sulfur-saturated sites becomes increasingly more important as found at low H2 pressure. At 380~ the rate rS,H2 was faster than the vacancy rate rv,H2 at high H2S, and a small, but non-meaningful, contribution of H2S-generated species with vacancies was calculated. These results were observed for two sulfiding modes (Table 1), and different pressures of thiophene (Table 2).
60 Finally, the variations of initial rate (Fig. 5b) and conversion (Fig. 5a) were very similar, which is indicative that rates determined at low conversion, e.g. under differential conditions, may be relevant to high conversion HDS. 1.0 250 "T 0 r l a ~ 0.8 200 "T h--
..C:
o 150
r-. 0 . 6 o o
v
t-- 0 . 4
E 9 m
._o
100
L_
._~ 50 " ~9
-_
0.0 0.2 0.4 0.6 0.8 1.0
H2S Pressure (bar)
rS,H~
__---. _ _ - -
0 --r---- ' ~- ~--0.0 0.2 0.4 0.6 0.8 1.0
H2S Pressure (bar)
Figure 5. Influence of H2S on thiophene HDS measured at 340~ 30 bars: a- global conversion due to vacancies (V) and sulfur-saturated sites (S), global conversion (0), C4 (A), THT (r-i); b-initial contributing rates. In conclusion, the similarity of the response towards H2S regardless of the H2 pressure is inviting. Temperature appears to be the most important parameter since it fixes the maximal concentrations of catalytic sites and active species. Thiophene and its competitor H2S regulate the concentrations of active sites, and H2 and H2S those of active species. As suggested by Topsee et al. [10], the concentration of SH groups may become rate limiting at high temperatures. In summary, a CoMo catalyst works following three operating modes above 380~ i) at low H2S pressures, HDS results from the attack of the thiophene adsorbed onto vacancies by H species issued from hydrogen; ii) at medium H2S pressures, the conversion is related to vacancies and sulfur-saturated sites, both in comparable concentrations, with nearby species coming from H2 and H2S again in comparable amounts; iii) at high H2S pressures, sulfur-saturated sites with H2-issued species contribute mainly to the global rate. Below 360~ the HDS rate simplifies principally to the contribution of vacancies with species issued from hydrogen. In this case, H2S acts mostly as an inhibitor. 7. CATALYST FUNCTIONS- COMPARISON OF CATALYST Comparison of catalysts from literature data leads to discrepancy results [10]. Conversions are generally high, and thus the amount of H2S varied in a large extent. Most of the data are further obtained at low temperature where adsorption is strong. Therefore, correlation of model molecule activity with industrial HDS may be casual [2]. Moreover, it is not often straightforward which catalyst function has been measured in the model molecule conversion. With thiophene HDS under high pressure, C4 hydrocarbons are produced through both hydrogenolysis and hydrogenation pathways, but hydrogenolysis
61 largely prevails although hydrogenation contributes more upon adding H2S [5]. Kinetic simulation shows that the hydrogenation pathway always stays minor even at H2S pressures up to 3 bars Hence, thiophene conversion under pressure measures essentially the catalyst capacity for C-S bond breakage, whilst thiophene HDS at atmospheric pressure measures the catalyst capacity for the heteroring dearomatization [22, 23] 250
l_
A
-Q9 200 d
15o
m
100
b
B
!
|
50 o
I
I
I
I
50 100 150 200 250 Rate - 400~ 1 bar Figure 6 Relation between initial rates for thiophene HDS over various fresh and regenerated CoMo/AI203 catalysts 0
This is obvious when one compares the relative initial rates for the series of CoMo catalysts (Fig 6) Three catalysts deviate from the ideal line with slope unity For instance, catalyst B was found as active as catalyst A for thiophene HDS under pressure and more active at atmospheric pressure This does not arise from a change in site environment, since the product selectivities for each test were the same regardless of the CoMo catalysts Therefore, the catalysts function differently (n 250 t.. ..Q o oo
t,=.
o~ .Q 200
200
d 15o b
B
co 100
J
I
I1)
~
250
a- Hydrogenation
d ~
A
0
I
50
I
I
I-rn 121
I
100 150 200 250
Thiophene-
400~
100 !
50 0
150
O3
N t-
| rn
b- Hydrogenolysis
1 bar
50 0
I
I
I
I
0 50 100 150 200 250 Thiophene - 380~ 30 bars
Figure 7 Relation between initial rates for: a- benzene hydrogenation (380~ 80 bars) and ~iophene HDS (400~ 1bar); b- dibenzothiophene HDS (335~ 40 bars) and thiophene HDS (380~ 30 bars)
62 Indeed, the rates for benzene hydrogenation correlate the rates for thiophene HDS determined at atmospheric pressure (Fig 7a), but not those under pressure This further proves that thiophene HDS is limited by the preliminary ring dearomatization Rates determined under pressure are correlated by the rates for DBT HDS (Fig 7b) Here, DBT conversion proceeds mainly through the hydrogenolysis pathway [5] Figure 8a shows a representative product distribution versus reaction time, where DBT is essentially converted into biphenyl (BP) through hydrogenolysis, and to a lower extent into cyclohexylbenzene (CHB) through hydrogenation The excellent correlation between DBT and thiophene, although the reactions were studied in liquid and gas phase reactors, is in support of C-S breakage being rate-limiting for thiophene under pressure 1.0 250 250 r
0.8
(•a
AB P
O .m
o 0.6
E 150 L_ ,9.0 rn 100 "1O 50
L_ t,--
0.4
O
0.2 0.0
rO 200
0
60 120 180 240 Reaction time (min)
0
r
- 200 o -
150
E:
|
O L_
- 100-o >,, t-
-50
Qnn
= i i i 0 50 100 150 200 250 Industrial HDS activity
Figure 8 Reaction of DBT (335~ 40 bars): a- fractional conversion versus time; b- relation between industrial RVA and initial rates for CHB formation (n), and for BP hydrogenation at 10% (A) and 50% (~ conversion We then attempt to establish what function was relevant to industrial HDS This was done by comparing the relative rates for model molecules and real feed The industrial RVA do not correlate the thiophene HDS rates under pressure [5], indicating that C-S hydrogenolysis is not limiting real feed HDS The comparison with the rate for hydrogenation through THT was imprecise because this pathway contributes here little to global HDS However, with DBT, the initial rate for cyclohexylbenzene formation could be assessed. A very good correlation was thus obtained with the RVA values (Fig 8b) The correlation fails with the rates for biphenyl formation, in agreement with the conclusion draw from thiophene Furthermore, the correlation stands also with the calculated rate for hydrogenation of biphenyl into CHB (Fig 8b) We can therefore conclude that real feed HDS is limited by the hydrogenation capacity of the sulfided catalysts The question now is to ascribe the type of catalytic sites and that of species coming from H2 or H2S needed to enhance the hydrogenation properties of the catalysts This was investigated with thiophene HDS at atmospheric pressure and 400~ where the individual contributing rates could be easily determined Figure 9 shows the response to added H2S of the global rate and the three significant contributing rates Thiophene HDS is there totally controlled by dearomatization, and, as expected, the rates rank satisfactorily with industrial
63
RVA in the whole range of added H2S (Fig. 10a). Hence, the thiophene test corroborates what was obtained with biphenyl hydrogenation and CHB formation (Fig. 8b). 80
80 o) "T t--
60
O
40
v
E |
n,"
"T tO v
E
b- Vacancies - H2
60
4O
20
-,-, 20
0 40
0 20
-
I
c- Vacancies - H2S
"T
I
-
~
.
"
I
.
=
I
d-Sulfur sites- H2 r-
20
o v
E
10
~
n,,
,
.....-..=
...~--~'T"-'T-". ~
.
,-
"'.--
~.-.--,-
,
o
0.0
0.1
0.2
0.3
0.4
0.5
0.0
H2S pressure (bar)
0.1
0.2
0.3
0.4
0.5
H2S pressure (bar)
Figure 9. Influence of H2S on global rate and contributing rates for thiophene HDS (400~ 1 bar) over various catalysts; RVA 100 (~), 135 (A), 140 (D), 180 (O). 250 G)
200
l a- Global
250
f S
~ 150 (D
. B
100 50 0 if
~
/
'- 150 G) > r 100 (l) iv' 50
(D . B
n,'
(D
_b- Vacancies -
200
i 50
I 100
i i 150 200 250
Industrial HDS activity
0
0
I
I
50
100
I
I
150 200 250
Industrial HDS activity
Figure 10. Correlation between industrial RVA and relative rate for thiophene HDS (400~ 1 bar) at 0.02 (0), 0.2 (D), and 0.5 (A) bar of H2S: a- global HDS; b- contributing rate due to vacancies and species from H2.
64 When examining the individual relationship, the contributing rates due to H2S species (Fig. 9c) or sulfur-saturated sites (Fig. 9d) fail to correlate the industrial RVA. The rates due to vacancies and H2-generated species are only representative of the RVA (Fig. 10b); the correlation was excellent comparing to that obtained with global HDS (Fig. 10a). Conclusively, improvement of real feed HDS should compel to an increase in the concentration of vacancies of the sulfided CoMo phases. Furthermore, the catalyst should be able to dissociate hydrogen into H species active for hydrogenation. This orientates future research aimed to catalyst development. 8. CONCLUSION
We have shown that the reactivity of model sulfur compounds is controlled by their adsorption capacity and charge density of the heteroring. The product H2S behaves as a competitor, and modifies the concentrations of active sites and surface H species of the sulfided CoMo catalyst. The superficial modifications brought by a change in thiophene conversion and operating conditions were evaluated from kinetics. Above 380~ HDS is due mainly to the activity of vacancies and sulfur-saturated sites with nearby species issued from H2 dissociation, the action of vacancies prevailing at low H2S content. Below 360~ HDS resumes to the activity of vacancies with species coming from hydrogen. In this case, the major role of H2S is to inhibit the conversion. The contribution of both hydrogenation and hydrogenolysis pathways to global HDS was appraised. Either benzene hydrogenation or DBT hydrogenolysis could correlate thiophene HDS. Under suitable conditions favoring hydrogenation, a correlation was found between thiophene and real feed HDS. It is concluded that the hydrogenation capacity due to vacancies with H2-dissociated species should be developed to obtain better hydrodesulfurization CoMo/AI203 catalysts. Acknowledgement Two of us acknowledge financial support, J.N.M. van Gestel from Elf-Aquitaine and L. Finot from EIf-Antar. We thank A. Travert who performed the theoretical calculations. 9. REFERENCES
1. 2. 3. 4. 5.
G. Heinrich, S. Kasztelan and L. Kerdraon, Rev. Inst. Fr. Petr, 49 (1994)475. S. Betteridge and R. Burch, Appl. Catal., 23 (1986) 413. H. Tanaka, M. Boulinguiez and M. Vrinat, Catal Today, 29 (1996) 209. E. Lecrenay, K. Sakanishi and I. Mochida, Catal. Today, 39 (1997) 13. J. Leglise, J.N.M. van Gestel, L. Finot, J.C. Duchet and J.L. Dubois, Catal. Today, 45 (1998) 347. 6. L.D. Rollmann, J. Catal, 46 (1977) 243. 7. P. Zeuthen, P. Stolze and U.B. Pedersen, Bull. Soc. Chim .Belg., 96 (1987) 985. 8. C. Sudhakar, L.T. Mtshali, P.O. Fritz and M.S. Patel, Proc. of the 10th Int. Cong. on Catal., Akademiai Kiad0, Budapest, 1993, p. 1418. 9. G.F. Froment, G. A. Depauw and V. Vanrysselberghe, Ind. Eng. Chem. Res., 33 (1994) 2975.
65 10. H. Topsee, B.S. Clausen and F.E. Massoth, in "Hydrotreating Catalysis, Science and Technology", and references therein, Springer-Verlag, Berlin, 1995. 11. J. Leglise, A. Janin, J.C. Lavalley and D. Comet, J. Catal., 114 (1988) 388. 12. X. Ma, K. Sakanishi and I. Mochida, Ind. Eng. Chem. Res., 33 (1994) 218. 13. M. Zdrazil, Coll. Czech. Chem. Comm., 42 (1977) 1484. 14. M. Houalla, D.H. Brodedck, A.V. Sapre, N.K. Nag, V.H.J. de Beer, B.C. Gates and H. Kwart, J. Catal., 61 (1980) 523. 15. M.J. Girgis and B.C. Gates, Ind. Eng. Chem. Res., 30 (1991)2021. 16. G.H. Singhal, R.L. Espino and J.E. Sobel, J. Catal., 67 (1981) 446. 17. R.C. Read, J.M. Prausnitz and B.E. Poling, in "The properties of gases and liquids", 4t" Edition, McGraw-Hill, New York, 1987, p. 11. 18. N.K. Nag, Appl. Catal., 10 (1984) 53. 19. D.L. Sullivan and J.G. Ekerdt, J. Catal., 178 (1998) 226. 20. E.J.M. Hensen, M.J. Vissenberg, V.H.J. de Beer, J.A.R. van Santen and R.A. van Santen, J. Catal., 163 (1996) 429. 21. H. Schulz, N.M. Rahman, Proc. of the 10~ Int. Cong. on Catal., Akad~miai Kiad6, Budapest, 1993, p. 585. 22. M. Zdrazil, Appl. Catal., 4 (1982) 107. 23. J. Leglise, J.N.M. Van Gestel and J.C. Duchet, Am. Chem. Soc. Div. Petr. Prep., 39 (1994) 533. 24. W.H.J. Stork, Stud. Surf. Sci. Catal., 106 (1997) 41. 25. J.F. Le Page et al., in "Applied Heterogeneous Catalysis, Design, Manufacture, Use of solid Catalysts", Technip, Pads, 1987, p. 387. 26. T. Kabe, K. Akamatsu, A. Ishihara, S. Otsuki, M. Godo, Q. Zhang and W. Quian, Ind. Eng. Chem. Res., 36 (1997) 5146. 27. S. Kasztelan and D. Guillaume, Ind. Eng. Chem. Res., 33 (1994) 823. 28. J.N.M. Van Gestel, L. Finot, J. Leglise and J.C. Duchet, Bull. Soc. Chim. Belg., 104 (1995) 189. 29. J.N.M. Van Gestel, J. Leglise and J.C. Duchet, Chem. Ind. Ser. 58, (1994) 357. 30. T.L. Tarbuk, K.R. McCrea, J.W. Logan, J.L. Heiser and M.E. Bussell, J. Phys. Chem. B, 102 (1998) 7845. 31. R.J. Angelici, Bull. Soc. Chim. Belg., 104 (1995) 265. 32. H. Jobic, G. Clugnet, M. Lacroix, S. Yuan, C. Mirodatos and M. Breysse, J. Am. Chem. Soc., 115 (1993) 3654. 33. N.Y. Tops~e and H. Topsee, J. Catal., 119 (1989) 252. 34. J. Miciukiewicz, W. Zmierczak and F.E. Massoth, Bull. Soc. Chim. Belg., 96 (1987) 915. 35. L. Portela, P. Grange and B. Delmon, Bull. Soc. Chim. Belg., 100 (1991) 985. 36. I.A. van Parijs and G.F. Froment, Ind. Eng. Chem. Prod. Res. Dev., 25 (1986) 431. 37. H. Kwart, G.C.A. Schuit and B.C. Gates, J. Catal. 61 (1980) 128. 38. T.S. Smit and K.H. Johnson, Chem. Phys. Lett., 212 (1993) 525.
This Page Intentionally Left Blank
THE CATALYST
UNDER WORKING CONDITIONS
This Page Intentionally Left Blank
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
69
Highly active MoS2-based dispersed catalyst with a novel morphology Y. Araki a, Y. Iwata a, Y. Miki a, K. Honna a, N. Matsubayashi b, and H. Shimada b* aTsukuba-branch, Advanced Catalysts Research Laboratory, Petroleum Energy Center, 1-1 Higashi, Tsukuba, Ibaraki 305-8565, Japan bNational Institute of Materials and Chemical Research, 1-1 Higashi, Tsukuba, Ibaraki 305-8565, Japan
Abstract The present paper introduced the preparation and characterization of a highly active unsupported MoS2 catalyst for hydrogenation of aromatic rings. In the present method, ammonium-tetrathiomolybdate was decomposed in an autoclave under water and a high pressure of H2. The high hydrogenation activity of the catalyst was attributed to very low stacking of MoS2 layers with much curvature that was confirmed by transmission electron microscopic observation, low (002) reflection peak intensity by X-ray diffraction, and other spectroscopic techniques. The surface area of the freshly prepared catalyst was 250 m2/g, but significantly reduced after the reaction at a high temperature. The hydrogenation activity of the present catalyst was further enhanced by the addition of a Ni promoter and exceeded that of oil-soluble MoDTC in the reaction of tetralin.
1. INTRODUCTION For the upgrading of heavy oils with residues, increasing attention has been paid to the slurry phase system that can avoid serious plugging problems of both catalyst bed and catalyst pore-mouth in the conventional fixed bed reactor system. In the slurry phase system, the use of dispersed catalysts with high hydrogenation activity is essential to stabilize the thermally formed radicals at high reaction temperatures. * corresponding author
70 In the past studies, several kinds of molybdenum-containing compounds that were water- or oil-soluble were proposed as the origin of the dispersed catalysts [1,2]. These molybdenum-containing compounds are decomposed into very small MoS2 particles under the reaction conditions and exhibit high catalytic activities. From the viewpoint of catalyst design, however, it is difficult to control the structure of the active MoS2 phase originated from these complexes. From this point of view, we have prepared e x s J t u some kinds of MoS2 fine particles and studied the structure-activity relationship [3]. The results suggested that the curvature of the basal planes of MoS2 crystallites were catalytically active for the hydrogenation of aromatic rings. A recent paper [4] also reported that MoS2 crystallites with low stacking were more active for the hydrocracking of 4-(1-naphthylmethyl)bibenzyl than those with high stacking. In the present study, we prepared MoS2 fine particles starting from ammonium-tetrathiomolybdate (ATTM) under the presence of water to prevent the stacking of MoS2 layers [5]. The catalysts prepared by this method showed a significantly higher hydrogenation activity as compared with those prepared by conventional thermal decomposition. In addition, we examined the effects of a Ni promoter on the activity and structures of the unsupported molybdenum sulfide catalysts prepared by the present and conventional methods.
2. EXPERIMENTAL Three kinds of unsupported molybdenum sulfide catalysts were prepared by thermal decomposition of ammonium tetrathiomolybdate (ATTM, (NH4)2MoS4). In the first method, an autoclave (35 cm ~) was charged with 3 g of ATTM, 15 g of water and 8 MPa of hydrogen. The autoclave was then heated to 673 K and kept at the temperature for 1 h. After cooling, the fine powders of molybdenum sulfides (named MS-W) were collected by filtering. The other two catalysts (MSG and MS-GR) were prepared by conventional thermal decomposition of ATTM under a flow of 10%H2S/H2 at 673 K. In the preparation of MS-G, the temperature was raised at 2 K/min, while MS-GR was obtained at a heating rate of 50 K/min. Molybdenum dithiocarbamate [MoDTC, (R2NC)2S~Mo202, R=CsHIT] was used as an oil soluble catalyst. Three kinds of nickel-molybdenum sulfide catalysts were prepared in the following procedures. The aqueous solution of a mixture of ATTM and ammonium nickel sulfate [(NH4)2Ni(SO4)2"6H20] was evaporated at 343 K on an evaporating dish. The black powder obtained was then dried at 383 K and decomposed at 673 K for 1 h under the following two conditions. NM-W was obtained under the same conditions in the same m a n n e r as described in the preparation of MS-W. For NM-G, the decomposition was carried out under a
71 flow of 10%H2S/H2. N-MS-G was p r e p a r e d by i m p r e g n a t i o n of MS-G with a solution of nickel (II) nitrate h e x a h y d r a t e [Ni(NO3)2"6H20] in acetone. Before the reactions, N-MS-G was sulfided under a flow of 10%H2S/H2 at 723 K for 1 h. The hydrogenation activities of the catalysts were e x a m i n e d using 1m e t h y n a p h t h a l e n e (1-MN) and t e t r a l i n (TLN) as model compounds. The reaction of 1-MN was carried out in an autoclave (inner volume of 35 cm 3) containing 10 cm 3 of the feed (25% 1-MN/tetradecane) and hydrogen with an initial pressure of 6 M P a at 603 K for 1 h. The reaction of TLN was carried out in the same autoclave containing 5 cm 3 of TLN and hydrogen with an initial pressure of 6 M P a at 663 K for 2 h. Each reaction rate constant was obtained using the relationship between the conversion and the catalyst weight t h a t was varied from 0.005g~0.20g. BET surface area of the catalysts was m e a s u r e d with a Micromertics ASAP 2010C i n s t r u m e n t . XRD p a t t e r n s were recorded using a Phillips P W l 8 0 0 diffractometer using Cu K a radiation. TEM observation was performed with a Hitachi 9000UHR operated at an accelerated voltage of 200 kV. XPS spectra were obtained on a PHI 5500 photoelectron spectrometer w i t h monochromated A1 Kcz excitation (1486.6 eV). EXAFS spectra were recorded at the Photon Factory (beam line 10B) of the National Laboratory for High E n e r g y Physics (Tsukuba, Japan).
3. R E S U L T S AND D I S C U S S I O N 3.1. Unpromoted catalysts Table 1 s u m m a r i z e s the hydrogenation activity and surface area of the u n s u p p o r t e d m o l y b d e n u m sulfide catalysts. In the reaction of I-MN performed at 603 K, MS-W was more active t h a n the other catalysts. In contrast, MoDTC Table 1 Specific surface area and rate constants for l - m e t h y l n a p h t h a l e n e and t e t r a l i n of Mo sulfide catalysts Catalyst MS-W MS-G MS-GR MoDTC
Rate constant .1 (1-MN) 43 7.8 9.2 14
*' 10 .5 mol/(mol of Metal)/s ,2 after the reaction of TLN
Rate constant*' (TLN) 3.9 1.3 1.3 5.4
Surface Area (m2/g) 251 9 61 -
Surface Area .2 (m2/g) 34 7 11 223
72 was the most active in the reaction of TLN at 663 K. The catalytic activities of MS-G and MS-GR were much lower t h a n those of MS-W and MoDTC. The different orders of the catalytic activities in the reactions of 1-MN and TLN can be attributed to the difference in the reaction temperature. According to literatures [6], MoDTC was decomposed at 578 K. The reaction temperature for 1-MN (603 K) was probably too low for MoDTC to completely decompose into MoS2 fine particles during the reaction. The present results suggest t h a t the high activity of MoDTC in the reaction of heavy feedstocks [7] is attributed to the formation of ultra fine MoS2 particles, not to the transient species before the formation of MoS2. The surface area of the catalysts decreased in the order of MS-W > MS-GR > MS-G. Some previous studies [8-10] reported considerably high surface area of 50-100 m2/g for the thermally decomposed catalysts of ATTM, which was comparable to that of MS-GR in the present study. Previous papers [8,11] reported t h a t the surface area of these catalysts were dependent on the heating rate during the decomposition. This discussion was consistent with the present results showing a higher surface area of MS-GR than MS-G. In spite of the high surface area, the catalytic activity of MS-GR was almost the same as that of MS-G. The structure of MS-GR was quite unstable and easily aggregated during the reactions as evidenced by the low surface area after the reaction. It may be worth to note that the surface area of the MS-GR was not very reproducible. The surface area of MS-W decreased to a large extent after the reaction of TLN. This may be partly due to the carbonaceous deposition, but as a major part due to the aggregation of small particles. As the result, the catalytic activity of MS-W was lower than that of MoDTC in the reaction of TLN. In addition, the superiority of MS-W to MS-G relatively decreased from 5.5-fold to 3-fold with ~200 ~175 ~150 ~125 I00 < 75 50 25 o
!
,/Mo3d-../~ _
_
_
0
0.2 0.4 0.6 0.8 Relative P r e s s u r e (P/Po)
Fig. 1 N2 isotherm of the catalysts. 9 :MS-G, 9 :MS-W, 9 :MS-GR, A:MoDTC after the reaction of TLN
1
236
pure
, ,
" j/
234
~
232
/
~
230
~
M
S
228
-
W
226
224
222
Binding Energy (eV) Fig. 2 Mo3d and S2s XPS spectra of the catalysts; * after the reaction of TLN.
73 the increase in the reaction temperature from 603 K to 663 K. The structure of MS-G with a low surface area was relatively stable during the reaction (Table 1). Fig. 1 shows the nitrogen adsorption-desorption isotherms of the catalysts. MoDTC after the reaction exhibited type IV isotherm, indicative of mesoporosity. The other three catalysts exhibited type II isotherm suggesting almost no microor meso-porosity. Fig. 2 shows the Mo3d and S2s XPS spectra of the catalysts. All the spectra exhibited similar profiles to that of pure MoS2. This suggests that all the active species in these catalysts were with the chemical form of MoS2. The Mo/C ratio, calculated from the intensities of Cls and Mo3d spectra, of the used MoDTC was 0.59, which was comparable to that (0.92) of the fresh MS-W. This suggests that the catalytically active Mo species was not covered by carbon but was rather supported on the porous carbonaceous species formed during the reaction. Fig. 3 represents the Fourier Transform EXAFS of MS-W and MS-G. Both of the curves exhibited two main peaks corresponding to Mo-S (2.4 A) and Mo-Mo (3.2 A). The local structures around Mo were the same as that of pure MoS2 for both of the catalysts. This was quite consistent with the above XPS results. Fig. 4 shows the XRD patterns of the catalysts. All the patterns were characteristic of highly distorted MoS2-1ike structures. The (002) reflection of MS-W was much less intense than that of MS-G, indicating lower stacking of MoS2 layers in MS-W. The patterns of MoDTC and MS-W after the reaction resembled that of the fresh MS-W. Using the Scherrer's equation, the average stacking of MoS2 layers in MS-W was estimated at 2.2 layers, while that of MSG was at 5.6 layers. The above discussions on the structures of the catalysts were confirmed by the TEM photographs of the catalysts (Fig. 5). The (002)
.
~
r~
v
MS-G ............
0
1
2
3 4 R(A)
5
6
Fig. 3 Fourier Transformed EXAFS of the catalysts.
7 5
15
25 35 45 2 0 (degree)
55
Fig. 4 XRD patterns of the catalysts. *after the reaction of TLN
65
74
Fig. 5 TEM photographs of the catalysts; (a) MS-W, (b) MS-G photograph of MS-W displayed only a few stacking of MoS2 layers with morphology like a soft feather, while stacking of several MoS2 layers was observed in the photograph of MS-G. Summarizing the above results and discussions, MS-W with low stacking was highly active for the hydrogenation of aromatic rings. As discussed in the previous paper [3], the high catalytic activity of the molybdenum sulfide catalysts with low stacking was ascribed to the much curvature and high surface area. However, the aggregation of the crystallites during the reactions at high temperatures reduced the activity. The superiority of MoDTC was attributed to the high surface area even after the reactions. This is probably due to the formation of fine MoS2 structures supported on porous carbon during the reactions. It might be essential to support the ultra fine particles of molybdenum disulfide on very fine particles in order to suppress the aggregation of the active species. 3.2. Ni promoted catalysts Fig. 6 shows the hydrogenation activities of the NiMo catalysts for 1-MN. The catalytic activity of MS-W was greatly enhanced by the Ni promoter with Ni/(Ni+Mo)=0.2. On the other hand, the activity of MS-G was hardly improved by the two conventional methods of the Ni addition. The surface area of NM-W was 213 m2/g, which was comparable to that of MS-W. In the reaction of TLN at 663 K, the activity of NM-W [rate constant: 6.7)< 10 .5 mol/(mol of Metal)/s)] exceeded that of MoDTC (Table 1). Fig. 7 shows the TEM photographs of NM-W [Ni/(Ni+Mo)=0.2] and N-MS-G [Ni/(Ni+Mo)=0.7]. The picture of NM-W exhibited a few stacking of MoS2 layers with curvature that was similar to MS-W. On the other hand, the TEM
75 ~.40 /
/
i 20
\
+,
. N M - W
.NM-a
_
80 70%" 60 ~
A
50 ~ 4o ~
30 ~ 20 ~ 10 ~ 0
~ ~1o % 0 0
0.2
0.4 0.6 Ni/(Ni+Mo)
0.8
1
Fig. 6 Hydrogenation of the NiMo sulfide catalysts for of I-MN.
0
1
2
3
4
5
6
7
R(A) Fig. 8 Fourier Transformed EXAFS of the catalysts.
Fig. 7 TEM photographs of the catalysts; (a) NM-W : Ni/(Ni+Mo)=0.2, (b) NMS-G : Ni/(Ni+Mo)=0.7 photograph of N-MS-G displays l i n e a r MoS2 layers that are surrounding large nickel sulfide particles. A similar trend was observed in the EXAFS spectra (Fig. 8) that showed an increase in the peak corresponding to the Mo-Mo scattering by the Ni addition. MS-G without Ni had a layer structure with much curvature (Fig. 5b), which provided high hydrogenation activity [3]. The Ni addition reduced the curvature that resulted in the decrease in the activity. In summary, a highly active hydrogenation catalyst was obtained by the addition of Ni to MS-W. This was because the Ni addition did not destruct the basic feature of MS-W, in contrast to the case of N-MS-G.
76 4. CONCLUSION An unsupported molybdenum sulfide catalyst prepared from ATTM with water and a high H2-pressure showed a very high hydrogenation activity for 1MN and TLN. The high hydrogenation activity was mainly attributed to the low stacking structure with much curvature and the high surface area. The catalytic activity of the present molybdenum sulfide catalyst was further enhanced by the addition of Ni promoter, in contrast to that of the conventionally prepared molybdenum sulfide catalysts. The superiority of MoDTC to the other unsupported molybdenum sulfide catalysts was attributed to the high surface area of the used catalyst after the reactions. For the further improvement of unsupported molybdenum sulfidebase catalysts, it is necessary to establish a method to suppress the aggregation of the molybdenum sulfide particles during the reaction at high temperatures.
5. ACKNOWLEDGMENTS This work has been carried out as a research project of the Petroleum Energy Center with the subsidy of the Ministry of International Trade and Industry. EXAFS measurements were carried out under the approval of the Photon Factory Advisory Committee (proposal PF98G333).
6. REFERENCES
1 N. Rueda, R. Bacaud and M. Vrinat, J. Catal., 169 (1997) 404. 2 B.R. Utz, A. V. Cugini and E. A. Frommell, Am. Chem. Soc. Symp. Ser. 437 (1990) 289. 3 Y. Iwata, K. Sato, T. Yoneda, Y. Miki, Y. Sugimoto, A. Nishijima and H. Shimada, Catal. Today, 45 (1998) 353. 4 Y. Yoneyama, C. Song and K. M. Reddy, Prepr. Am. Chem. Soc. Div. Pet. Chem., 42 (1997) 550. 5 P. Joensen, R. F. Frindt and S. R. Morrison, Mater. Res. Bull., 21 (1986) 543 6 H. Isoyama and T. Sakurai, Tribology international, Aug., (1974) p.151 7 K. Sato, Y. Iwata, K. Honna, Y. Araki, T. Yoneda, Y. Miki and H. Shimada, Stud. Surf. Sci. Catal., 121 (1998) 411. 8 D.G. Kalthod and S. W. Weller, J. Catal., 95 (1985) 455. 9 R.R. Chianelli, M. Daage and M. Ledoux, Adv. Catal., 40 (1994) 177. 10 C. Calais, N. Matsubayashi, C. Geantet, Y. Yoshimura, H. Shimada, A. Nishijima, M. Ledoux and M. Breysse, J. Catal., 174 (1998) 130. 11 A. W. Naumann, A. S. Behan and E. W. Thorsteinson, Proc. Fourth Int. Conf. on The Chemistry and Uses of Molybdenum, (1982) p.313.
Hydrotreatment and Hydrocrackingof Oil Fractions B. Delmon, G.F. Fromentand P. Grange (Editors) 91999Elsevier ScienceB.V. All rights reserved.
77
USE O F N O B L E M E T A L S IN H Y D R O D E O X Y G E N A T I O N R E A C T I O N S Arist6bulo Centeno (b), Rosanna Maggi ~a) and Bernard Delmon ta) (a) Unit6 de Catalyse et Chimie des Mat6riaux Divis6s, Universit6 Catholique de Louvain, Place Croix du Sud 2/17, 1348 Louvain-la-Neuve, Belgium. (b) Present address: Centro de Investigaciones en Catfilisis, Escuela de Ingenieria Quimica, Universidad Industrial de Santander, A.A. 678, Bucaramanga, Colombia. Fax: 57 7 6350540. e-mail:
[email protected] Abstract Two series of catalysts supported on active carbon containing noble metals (NM) Pt, Pd, Ru and Rh were prepared (bimetallic NM-Mo/C and trimetallic NM-CoMo/C catalysts) and tested in their sulphided form in hydrodeoxygenation (HDO) reactions of model molecules containing carbonyl, carboxyl, hydroxyl and methoxy groups. Results show that hydrogenation reactions of the carbonyl and carboxyl groups are favoured by the presence of NM in the case of the bimetallic catalysts. The influence of cobalt in the decarboxylation reaction was confirmed. Bimetallic catalysts without cobalt have no decarboxylation activity. A mutual interaction between noble metal sulphides and cobalt sulphide was observed for trimetallic catalysts. Effects of this interaction are different for each metal, which indicates that the noble metal nature is implicated in this phenomenon.
1. I N T R O D U C T I O N The present work corresponds to part of a program aimed at upgrading oil obtained by pyrolysis of biomass by hydrotreating (HDT). In previous studies [1] we reported that coke formation on typical HDT catalysts is the principal limitation to the use of these conventional HDT catalysts in hydrodeoxygenation (HDO) of biomass pyrolysis oil. In order to improve the catalytic performance of catalysts to upgrade biomass oil, two possibilities have been considered. The first one is to change the alumina support for an inert support like active carbon. The other one is to find a catalytic system that is more active for HDO. This would permit work in a lower temperature range, where coking reactions take place more slowly. A combination of these two solutions is understandably preferable. It is known that noble metals (NM) have good hydrodesulphurizing (HDS) properties and excellent hydrogenating (HYD) activities. The use of (NM) in HDT of model molecules has been reported in the literature especially for HDS and hydrodenitrogenation (HDN) [2, 3]. The general conclusion is that NM sulphides are more active in HDS, HYD and HDN reactions compared to Mo or CoMo catalysts and that a maximum of activity as a function of composition is present for some of them [3]. Recently some works concerning the use of bimetallic and trimetallic catalysts in HDS and HYD have been reported in the literature. Xiao et al [2] reported that Ru-CoMo/~,AIzO3 is more active in the HDS of thiophene than CoMo/~-AIzO3. Vrinat, Breysse et al [4] reported that unsupported NiRuS and FeRuS catalysts were more active than NiMoS, CoMoS
78 and FeMoS catalysts. These authors also found a synergetic effect in RuMo/7-AI203 and NiRu/T-AI203 catalysts in HDS of thiophene and HYD of biphenyl. Information concerning HDO for this subject is scarce. Shabtai et al [5] used bimetallic NM-Mo catalysts supported on alumina in the simultaneous hydrogenolysis of C-O bond of diphenyl ether and naphthalene hydrogenation. They found that sulphided bimetallic catalysts presented a more important hydrogenolysis activity compared to monometallic ones. We present in this work results concerning the use of sulphided bimetallic (NM-Mo/C) and trimetallic (NM-CoMo/C) catalysts supported on active carbon in the HDO of model molecules containing carboxyl, carbonyl, hydroxyl and methoxy groups. XPS measurements were made in order to characterise catalysts. 2. E X P E R I M E N T A L
2.1. Catalysts Two series of catalysts containing NM (Pt, Pd, Rh and Ru) supported on active carbon were prepared. The first one consisted in bimetallic NM-Mo/C catalysts containing 1 wt% NM and 15 wt% MOO3. The second one consisted in trimetallic NM-CoMo/C catalysts containing 0.5 wt% NM, 3 wt% CoO and 15 wt% MOO3. Monometallic Mo/C and bimetallic CoMo/C were also prepared as reference catalysts. Wet impregnation was always used. Volume of solution was 20 times the pore volume of the support. Solids remained in the solution for 12 h. Afterwards, they were dried in a rotovapor at 303 K and subsequently dried under a flow of air at 403 K during 12 h. Impregnation order was Mo-NM for bimetallic catalysts and Mo-Co-NM for trimetallic ones. Heating rate was 3 K min 1. A standard reduction-sulphidation procedure [6] was realized after each impregnation. Ammonium heptamolydate, cobalt nitrate, tetra-aminoplatinum (II) nitrate, tetra-aminopalladium (II) nitrate, ammonium hexachlororhodate and ammonium hexachlororuthenate were the salts used. Active carbon (Merck) of 1300 m2.g -1 surface area was the support. 2.2.Reaction system A 570-ml batch reactor was used for the catalytic tests at 70 bar of Ha and 553 K. Reaction procedure has been reported in a previous work [6]. The model reactants were 4methylacetophenone (4-MA, 0.218 mol 1-1), 2-octanone (2-octa, 0.229 tool 1-1), ethyldecanoate (edec, 0.147 mol 1-1), 2-methoxyphenol (guaiacol: gua, 0.237 mol 1-1), 4-methylphenol (pcresol, 0.110 mol 1-1). Two different ketones were added to this solution in order to study the performance of the catalysts in reducing carbonyl groups at two different reactivity levels. They were dissolved in p-xylene; n-pentadecane was added to this solution as an internal standard for the chromatograph analysis. CS2 was added as a precursor of HaS to maintain catalysts in a sulphided state. 2.3. Activation procedure The standard activation procedure used in our previous work [6] was modified in the present study. Now a 50 vol% HaS in Ha mixture was used during reduction-sulphidation at 673 K for 3 h. 2.4. Analysis Liquid samples were analysed in a Chrompack model CP-9001 gas chromatograph equipped with a split injector and an FID detector. Samples of the catalysts after sulphidation
79 were recovered in isooctane to avoid any oxidation by air in order to make XPS analysis. XPS measurements were made in an SSI X-probe (SSX-100/206) instrument. Peaks of C1s, O1s, Sap, Mo3a, Co2p and Pt4f, Pd3a, Ru3d or Rh3a were registered. The peak of C ls was used as reference for the binding energy calculations. 2.5. Expression of the results The activity of the catalysts, characterised by disappearance of the reactants, is reported as a pseudo first-order rate constant, (k). Based on HDO reaction schemes [1] of the different model compounds used, two kinds of selectivities were defined: Sdec and phe/cat ratio. The first one represents the capacity of catalysts to develop decarboxylation reactions related to hydrogenation reactions in the HDO of carboxylic esters. The second one represents the importance of Caromatic - O bond hydrogenolysis related t o C m e t h y l i c - O bond hydrogenolysis in the HDO of guaiacol. 3. RESULTS 3.1. Performance of bimetallic catalysts Important changes in activity and selectivity compared with reference catalysts were found when bimetallic catalysts containing noble metals were used in HDO reactions. Table 1 presents the results of activity and selectivity for bimetallic and reference catalysts. These results correspond to the conversion of carbonyl groups of 4-MA and 2-octa, carboxyl groups of ethyldecanoate and methoxy groups of guaiacol. Table 1 Activity and selectivity of bimetallic and reference catalysts in HDO of 4-MA, 2-octa, edec and gua. Activity is expressed by a pseudo first-order rate constant (k in: min-l.g cat.cm3). Temperature 553 K, pressure 70 bar of H2. 4-MA 2-octa edec gua Catalyst k Sdec k phe/cat Mo/C 4.8 1.4 0.25 0.56 0.57 PtMo/C 14.7 5.7 0.37 0.62 0.87 PdMo/C 5.8 1.7 0.27 0.51 0.62 RhMo/C 9.7 3.1 0.34 0.65 1.10 RuMo/C 8.6 2.2 0.39 0.63 0.85 CoMo/C 9.9 4.5 0.32 0.22 0.33 0.38 Table 1 shows that, except for palladium catalysts, addition of NM increases significantly the activity of catalysts to convert carbonyl groups compared with Mo/C. Activities of NMMo/C catalysts are always higher than that of the CoMo/C catalyst. The activity of the PtMo/C catalyst is the highest and that of PdMo/C is the lowest. It is possible to indicate a decreasing activity order for the different catalysts tested in the carbonyl groups conversion: Mo-Pt/C >> Mo-Rh/C > Mo-Ru > Mo-Pd/C = Mo/C Table 1 also shows that activity and selectivity changed in HDO of carboxyl group when NM was used. An interesting result is that bimetallic NM-Mo/C catalysts and monometallic Mo/C catalysts have no decarboxylation activity. Decarboxylation is only present in the case
80 of CoMo/C catalyst. All bimetallic catalysts, except PdMo/C, present a higher activity than that of the Mo/C catalyst. Activities for all bimetallic catalysts are similar. Activities of all NM-Mo/C catalysts for guaiacol conversion are similar to the one of the Mo/C catalyst, but higher than that of CoMo/C. The phe/cat ratio for NM-Mo/C catalysts is always higher than those of Mo/C and CoMo/C catalysts. This ratio is highest in the case of RhMo/C. 3.2. Performance of trimetallic catalysts Table 2 presents the results of activity and selectivity for trimetallic catalysts in HDO reactions. Table 2 Activity and selectivity of trimetallic and CoMo/C catalysts in HDO and gua. Activity is expressed by a pseudo first-order rate constant Temperature 553 K, pressure 70 bar of H2. 4-MA 2-octa edec Catalyst k Sdec PtCoMo/C 13.6 6.5 0.20 0.27 PdCoMo/C 13.5 6.7 0.17 0.26 RhCoMo/C 12.6 6.0 0.26 0.25 RuCoMo/C 11.6 6.0 0.32 0.20 CoMo/C 9.9 4.5 0.32 0.22
of 4-MA, 2-octa, edec (k in: min-l.g cat.cm3). gua k 0.29 0.26 0.35 0.27 0.33
phe/cat 0,27 0.28 0.65 0.39 0.38
Concerning results presented in Table 2, it is possible to remark the following: 9 Trimetallic catalysts present activities for the conversion of ketones, always higher_than that of CoMo/C. In spite of the fact that activities are similar, it is possible to discern a decreasing order of activity for this group: PtCoMo -- PdCoMo > RuCoMo> RhCoMo > CoMo. 9 Trimetallic catalysts do not improve activity of CoMo/C for the conversion of ethyldecanoate. In the case of PtCoMo/C and PdCoMo/C a decrease of activity is observed. The presence of decarboxylation products is detected, but the selectivity of decarboxylation is the same for all catalysts. The presence of decarboxylation products is the fact that points to the differences between bimetallic and trimetallic catalysts containing noble metals. 9 Addition of a noble metal, except Ru, to the CoMo/C catalyst, leads to a diminution of activity for guaiacol conversion. A diminution of the phe/cat ratio, except in the case of RhCoMo/C catalyst, is observed. For this catalyst this ratio is twice the one presented by CoMo/C. 3.3. XPS measurements Binding energies for Mo and for most noble metals bimetallic or trimetallic catalyst if compared with those CoMo). Nevertheless, for Pt a difference of 0.4 eV PtCoMo/C. In all cases binding energy values correspond
do not change significantly for any of the reference catalysts (Mo, and is observed between PtMo/C and to those of the metal sulphides.
81
Table 3 presents quantitative XPS results of different metals on the surface of the catalysts. Molybdenum/carbon, cobalt/carbon, noble metal/carbon and sulphur/carbon atomic ratios are presented for bimetallic and trimetallic catalysts. As observed in Table 3, the molybdenum/carbon atomic ratio for trimetallic catalysts is similar to that of CoMo/C and Mo/C catalysts. That means that the presence of a noble metal does not significantly change the dispersion of molybdenum at the surface of carbon in these cases. Nevertheless, bimetallic catalysts show two different tendencies. This atomic ratio is higher for RuMo/C and RhMo/C catalysts and lower for PtMo/C and PdMo/C catalysts when compared with the one of the reference catalyst Mo/C. Only the presence of Pt significantly changes cobalt dispersion at the surface of catalysts. The Co/C atomic ratio of the other trimetallic catalysts is similar to that of the reference catalysts CoMo/C. Table 3. Metal/carbon and sulphur/carbon atomic ratios measured by XPS. A similar decreasing order of XPS signal intensity
Catalyst Mo/C PtMo/C PdMo/C Ru/Mo/C RhMo/C CoMo/C PtCoMo/C PdCoMo/C RuCoMo/C RhCoMo/C
Molybd./C 0.017 0.014 0.013 0.019 0.019 0.016 0.016 0.016 0.016 0.016
Cobalt/C
NM/C 0.085 1.043 1.420 0.407
0.005 0.012 0.003 0.004 0.004
0.029 0.578 0.803 0.193
Sulphide/C 0.048 0.045 0.052 0.116 0.062 0.051 0.077 0.066 0.078 0.058
A similar decreasing order of XPS signal intensity of noble metals is observed in bimetallic and trimetallic catalysts. This decreasing order is: Ru > P d >> Rh > Pt. As expected the NM/C atomic ratio for bimetallic catalysts is higher than that of the corresponding trimetallic catalysts. Noble metals do not change significantly the relative content of sulphur at the surface in the case of bimetallic catalysts, except for RuMo/C. This catalyst presents an S/C atomic ratio twice that of Mo/C. In the case of trimetallic catalysts, this atomic ratio for PtCoMo/C and RuCoMo/C is higher than that of CoMo/C. 4. DISCUSSION Noble metal addition leads to important changes in the performance of catalysts in HDO reactions. Characterisation and catalytic tests developed in this work are not sufficient to determine the origin of the main changes. Nevertheless, understanding some effects might contribute to obtaining more appropriate catalysts for hydrotreating biomass pyrolysis oils.
82 We will discuss each chemical group separately, always comparing with the behaviour of reference catalysts. We begin the discussion with a qualitative comparison of XPS results attempting to correlate them with the catalytic performances. 4.1. Carbonyl group. If we want to understand the effect of noble metals on catalytic systems it is necessary first to understand the bimetallic ones. It is possible to find answers to explain the catalytic behaviour in reduction of carbonyl groups in the following two hypotheses: a good dispersion of molybdenum on the active carbon surface or a direct effect of noble metal sulphides. XPS results show that the first hypothesis is not correct. It is not possible to correlate molybdenum dispersion with catalytic activity for carbonyl groups reduction. Molybdenum dispersion in both PtMo/C and PdMo/C catalysts goes in a direction opposite to that of RhMo/C and RuMo/C, but their catalytic activities do not go in the same direction. PtMo/C is the most active and PdMo/C is the least active. If this second hypothesis were considered we should conclude that noble metals in surface do not correlate in any way with catalytic activity. We conclude that the nature of the noble metal is the important parameter. We have reported previously [6] that the limiting stage in the reduction of carbonyl groups is the hydrogenation of the C - O bond. In other words, the capacity of catalysts to reduce carbonyl groups is related to their hydrogenating capacity. Literature reports that noble metals have an important hydrogenating capacity in both reduced and sulphide forms [6 - 8]. Noble metals are used to reduce C = C [9, 10] and C = O [11] bonds, but when they are in metallic form. Use of sulphides in these last reactions is not reported in the literature. We interpret oiJr results thinking that noble metals increase the catalyst capacity to hydrogenate C - O bonds. Differences in activity are related to the nature of the noble metal. We consider now trimetallic systems to discuss the combined effect of NM and Co. Compared with the reference catalyst CoMo/C, there is an increase in activity of trimetallic catalysts to convert carbonyl groups. It is not possible to explain this increase in activity considering differences in molybdenum or noble metals dispersion on the surface of carbon. XPS measurements show that the Mo/C atomic ratio for trimetallic catalysts is similar to that of the CoMo/C catalysts. NM/C atomic ratios do not correlate with catalytic activity. Except for the PtCoMo/C catalyst, Co/C ratios are similar to that presented by the CoMo/C catalyst. This ratio is two times higher for the catalyst containing Pt if compared with CoMo/C. There is a risk in considering that this is the determining parameter to increase activity because results from the use of other metals show opposite tendencies and activity always increases. It is important also to point out the important increase in the activity of a trimetallic catalyst containing Pd compared with the bimetallic one with the same noble metal. Molybdenum dispersion does not explain differences in activity between trimetallic and bimetallic catalysts. The only possibility to explain this apparent inconsistency might be found in an interaction between cobalt sulphide and noble metal sulphides [12]. This interaction is different for each noble metal. It would be strongly favourable in the case of Pd and negative in the other cases.
4.2. Carboxyl group. Our results show that bimetallic catalysts are active for ester conversion but do not present any decarboxylation activity. We have pointed out in the results section that decarboxylation products are only detected if cobalt is present in the catalysts. Weisser and Landa have reported that cobalt and nickel sulphides are not active for carboxyl group
83 conversion but, if they are associated with Mo, decarboxylation is catalysed [13]. This agrees with our observations. The influence of noble metals on the conversion of carboxyl group is as important as for carbonyl groups especially for most of the bimetallic catalysts. The fact that bimetallic and Mo/C catalysts do not present a decarboxylation activity and the fact that two simultaneous reactions take place during ester conversion (decarboxylation and hydrogenation) [1], indicate that the increase in activity is directly related to an increase in hydrogenation activity. Molybdenum or noble metal dispersions cannot explain this behaviour as in the case of carbonyl group hydrogenation. Again noble metal nature is implicated. It is understandable that tendencies must be the same because hydrogenation reactions are involved in both cases. It is necessary to refine the discussion considering either a simple addition of individual activities or a more complex interaction between sulphides, but our present results are not sufficient to conclude definitely. Nevertheless, previous work in our laboratory showed that Pd, Pt and Rh sulphides increase the activity of MoS2 and WS2 in HDS of thiophene and HYD of cyclohexene when mechanical mixtures were used [8, 14]. The synergetic effect found has been explained by the presence of "spillover" hydrogen. 4.3. Methoxy group. Bimetallic catalysts showed a slight increase in activity to convert the methoxy group of guaiacol compared with the Mo/C catalyst. This increase is associated to an important increase in the phe/cat ratio. These results are in agreement with those of Shabtai et al [5] who report an important increase in activity for the hydrogenolysis of the C - O bond compared with activity of Mo/],-A1203 when RhMo/],-A1203 and RuMo/],-AI203 were tested. These authors reported also a decrease in activity in the same case for PdMo/~,-AI203 and PtMo/~,A1203 catalysts. It is remarkable that the presence of noble metals in bimetallic catalysts has a positive effect on the conversion of guaiacol contrary to the negative effect presented by Co [15]. If we now consider the reaction mechanism proposed in the previous work [15], we might interpret our results. Two reaction pathways are always possible: elimination of a methyl group by hydrogenolysis of the CH3 - O bond and a direct elimination of a methoxy group by hydrogenolysis of the Caromatic - O bond. Noble metal sulphides favoured these two hydrogenolysis reactions. It is possible that the second one was the most favoured. This hypothesis might explain the high phe/cat ratio observed when noble metal bimetallic catalysts are used. 4.4. General c o m m e n t .
It is important to remark two common facts during the conversion of the three groups: a general effect on activity and selectivity, and an interaction between noble metal sulphides and cobalt sulphide in trimetallic catalysts. An explanation is that reaction conditions, namely H2S/H2 ratio, have a different influence on the performance of each noble metal sulphide. In the case of the interaction between sulphides, it is possible that cobalt sulphide decontaminates noble metal sulphides of part of the sulphur on their surface.
84 5. CONCLUSIONS The use of noble metals in bimetallic catalysts increases the activity of Mo/C in the HDO of carbonyl and carboxyl groups. This increase in activity must be explained by the increase in the activity of these catalysts to hydrogenate C = O bonds. Noble metal nature is the most important parameter in this case. Bimetallic catalysts NM-Mo/C are not active for the decarboxylation reaction. This fact contributes to confirm the role of cobalt in this reaction. The conversion of the methoxy group is favoured by the presence of noble metal sulphides in the catalysts. Noble metal sulphides increase activity for the hydrogenolysis of CH3 - O and Caromatic- O bonds. In trimetallic catalysts a mutual interaction between noble metal sulphides and cobalt sulphide always exists. Effects of this interaction are different for each metal. This indicates that noble metal nature is implicated. ACKNOWLEDGMENTS This work was made possible thanks to the financial support of the European Union (Contract No. JOR3-CT95-0025). A. Centeno specially thanks the Universidad Industrial de Santander, Bucaramanga, Colombia for its help that made possible his stay in Belgium. REFERENCES 1. 2. 3. 4. 5. 6. 7.
E. Laurent, and B. Delmon, Appl. Catal. 109 (1994) 77. F.S. Xiao, Q. Xin, and X.X. Guo, React. Kinet. Catal. Lett., 46 (1992) 351. T.A. Pecoraro, and R.R. Chianelli, J. Catal., 67 (1988) 430. M. Vrinat, M. Lacroix, M. Breysse, L. Mosoni, and M. Rouban, Catal. Lett. 3 (1989) 405. J. Shabtai, N.K. Nag, and F.E. Massoth, J. Catal., 104 (1987) 413. A. Centeno, E. Laurent, and B. Delmon, J. Catal., 154 (1995) 288. M. Lacroix, N. Boutarfa, C. Guillard, M. Vrinat, and M. Breysse, J. Catal. 120 (1989) 473. 8. S.A. Giraldo, PhD. Thesis (Universit6 catholique de Louvain) 1993. 9. L. Schmitt, and P.L.Jr. Walker, Carbon 9 (1971) 791. 10. L. Schmitt, and P.L.Jr. Walker, Carbon 10 (197) 87. 11. W.F. Maier, W. Roth, L. Thies, and P.V. Ragu6 Schleyer, Chem. Ber. 115 (1982) 808. 12. X. Vanhaeren, PhD. Thesis (Universit6 catholique de Louvain) 1997. 13. O. Weisser, and S. Landa, "Sulphide Catalysts, their Properties and Applications". Pergamon, Oxford, 1975. 14. S.A. Giraldo, P. Grange, and B. Delmon, in "New Aspects of Spillover Effects in Catalysis" (T. Inui, K. Fujimoto, T. Uchijima and M. Masai, Eds.), Studies in Surface Science and Catalysis, vol 77, p. 345. Elsevier, Amsterdam, 1993. 15. A. Centeno, PhD. Thesis (Universit6 catholique de Louvain) 1987.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
85
I n f l u e n c e of the h y d r o g e n sulfide partial p r e s s u r e on the h y d r o d e o x y g e n a t i o n reactions over sulfided C o M o / C a r b o n catalysts M. Ferrari, S. Bosmans, R. Maggi, B. Delmon, P. Grange Unit6 de Catalyse et Chimie des Mat6riaux Divis6s, Universit6 Catholique de Louvain, P1. Croix du Sud 2/17, B-1348 Louvain-la-Neuve, Belgium Abstract
This work concerns the influence of the hydrogen sulfide partial pressure on the activity in hydrodeoxygenation and associated reactions of a CoMo/Carbon catalyst for the transformation of guaiacol, ethyldecanoate and 4-methylacetophenone. This has been studied over a wide range of hydrogen sulfide partial pressures, from 10 to 150 kPa, under a total pressure of 7 MPa, at reaction temperatures of 270 and 200~ An inhibition effect is observed in the hydrogenolysis pathway in guaiacol conversion and in ethyldecanoate dehydroxylation. The conversion of 4-methylacetophenone is also inhibited by hydrogen sulfide. The results have been interpreted on the basis of the existence of different active sites responsible for hydrogenolysis, hydrogenation and acid catalysed reactions. The increase in hydrogen sulfide partial pressure would decrease the average degree of reduction of the active sites and inhibit the hydrogenolysis pathway, more than the hydrogenation one. The results are compared to those observed with alumina supported catalysts in hydrodeoxygenation and hydrodesulfurisation. 1. I N T R O D U C T I O N Hydrodeoxygenation (HDO) reactions have not been studied as much as hydrodesulfurisation (HDS) and hydrodenitrogenation (HDN). This may be attributed to the very low amounts of oxygen found in conventional fossil crudes and to the fact that oxygencontaining compounds are not as harmful for the catalysts and for the environment as sulfur and nitrogen ones (1). Problems associated with the presence of oxygen arise when alternative sources of energy, like liquids derived from the pyrolysis of ligno-cellulosic biomass, namely bio-oils, are considered (2). These liquids are characterised by a high oxygen content (O = 26-47%wt), which is the cause of high viscosity, low thermal stability, poor heating value and corrosivity. Bio-oils properties should be upgraded by partial or total oxygen elimination in order to enable their storage, transportation and easy utilisation. Catalytic hydrotreatment (3), more precisely hydrodeoxygenation (HDO), performed at moderately high temperature and under hydrogen pressure could be adapted for this purpose. The reactions involved are quite similar to those occurring in the hydrotreatment of petroleum fractions; they consist in the elimination of oxygen as water and in the hydrogenation-hydrocracking of various chemical functions contained in the molecules constituting the pyrolysis products. Some authors (4-6) have demonstrated the feasibility of bio-oils deoxygenation by catalytic hydrotreatment. Laurent et al. (7) have elucidated the reaction pathways of model compounds which contain the oxygenated functions responsible for bio-oil instability (carbonyl, carboxylic esters and phenolic ethers) over sulfided CoMo/AI203 and NiMo/A1203. Although alumina supported catalysts presented a good activity, they were quickly deactivated by coke formation, due to the acidity of the support (8, 9). This tendency to coke formation was found to be particularly due to phenol type molecules containing two or more oxygenated substitutes,
86 like guaiacol and catechol. The utilisation of neutral supports, such as activated carbons, has been shown to limit or to avoid the formation of condensation products acting as coke precursors (8, 9) and to facilitate the elimination of oxygenated groups from guaiacol and catechol. Activated carbons as catalyst supports for hydrotreating processes have received much attention in the last years (10, 11). Some advantages of carbon over alumina based catalysts have been highlighted, like a lower coking propensity (12-14) combined with a higher catalytic activity (14-18). Our work deals with the influence of the hydrogen sulfide partial pressure on the HDO activity of a CoMo catalyst supported on a commercial activated carbon (CoMo/C). As biomass derived liquids do not contain sulfur compounds, hydrogen sulfide is likely to be required to create the catalytic sites and to preserve the sulfided catalyst from oxidation by oxygenated compounds and water, or reduction of the sulfided phase by hydrogen, according to operating conditions. It was suspected that the hydrogen sulfide partial pressure could alter in different ways the catalytic functions as it does in HDS and HDN. It is known that hydrogen sulfide inhibits the hydrogenolysis of C-S bonds in HDS, but enhances the hydrogenation and breaking of C-N bonds in HDN (19). Former works, performed with alumina supported catalysts, have shown that hydrogen sulfide and other competitors for catalytic sites, like ammonia and water, strongly affect the HDO of model compounds, like 4-methylpheno| (20) and of mixtures of 4-methylacetophenone, diethylsebacate and guaiaco| (21). Hydrogen sulfide inhibits the hydrogenolysis pathway of 4-methylphenol and guaiacol (cleavage of CAromatic-O bond) and the conversion of 4-methylacetophenone. Even if our interest is focused on a very specific field, namely HDO of bio-oils, which is usually not of much concern for the petroleum industry, the results obtained contribute to the acquisition of new knowledge on the behaviour of the CoMo/C catalytic system. This kind of catalyst can find larger applications, in other hydrotreatment reactions, with other feedstocks, in particular with petroleum fractions, where similar reactions (hydrogenation, breaking of heteroatoms-carbon bonds) are involved. The understanding of the behaviour of CoMo/C catalysts can thus shed light on that of other CoMo supported catalysts. 2. E X P E R I M E N T A L The catalyst was prepared using a commercial activated carbon, BKK-100 (specific surface area 1070 m 2 g-l, pore volume 0.66 cm 3 g-l). The support was impregnated with aqueous solutions of ammonium heptamolybdate and cobalt nitrate, using the incipient wetness method. Molybdenum was impregnated first and cobalt second. After each impregnation, the sample was dried overnight at 130~ and at 400~ for 3 hours under argon flow. The reaction tests were carried out in a fixed bed continuous flow tubular reactor. 1.5 g of catalyst (particle diameter between 0.3 and 0.5 mm) was diluted with glass spheres (diameter between 0.2 and 0.45 ram) to reach a volume of 18 cm 3. The catalytic bed (8 cm height) was loaded in the reactor between two plugs of glass wool. The rest of the reactor was filled with 1 mm diameter glass spheres. Prior to reaction, the catalyst was dried at 130~ overnight, under nitrogen flow. The sulfidation mixture, 15%vol H2S in H2, was then introduced, the pressure set at 0.3 MPa and the flow at 150 m| min-1; after a stabilisation period of 30 min, the temperature was increased to 400~ (3~ rain-l). These conditions were held for 3 hours. The temperature was subsequently decreased to 270~ The gas was then switched to pure hydrogen and the liquid feed introduced. The pressure was progressively raised to 7 MPa, in 1 hour. The liquid feed rate was 45 ml h -1 and the hydrogen flow rate 24 1 h -1. The tests were performed at a total pressure of 7 MPa. The liquid feed contained oxygenated model compounds selected on the basis of an in depth chemical characterisation of bio-oils (22). The composition was as follows: guaiacol, GUA (3%wt, 0.21 mol 1-1), ethyldecanoate, ED (3%wt, 0.13 mol 1-1) and 4-methylacetophenone, MA (3%wt, 0.20 mol 1-1). Pentadecane (2%wt, 0.082 mol 1-1) was added as internal standard
87 for the chromatographic analysis and p-xylene was used as solvent. Dimethyldisulfide (DMDS), which rapidly decomposes in the presence of hydrogen, was added as a precursor of H2S. The reactants were from Aldrich and they were of a purity of 98% at least. Liquid samples were analysed by a gas-chromatograph equipped with an FID detector and a capillary column (stationary phase CP-Sil-8CB, length = 25 m). Conversion, yield and selectivity data reported hereafter were obtained after 18-24 hours of reaction, based on the analysis of at least three samples of the converted mixture. Two reaction temperatures were tested: 270~ in order to evaluate the reactivity of GUA and ED, 200~ in order to evaluate the reactivity of MA. Three series of experiments were done. In the first, four tests were performed at different hydrogen sulfide partial pressures (50, 75, 100, 150 kPa), using a fresh catalyst for each test. In the second, called "PHzS increase", the test was done at increasing hydrogen sulfide partial pressures (10, 25, 35, 50, 75 and 150 kPa) without changing the catalyst and always waiting 24 hours before modifying the operational conditions (temperature and hydrogen sulfide partial pressure). After having measured the activity at 150 kPa, the hydrogen sulfide partial pressure was decreased to 10 kPa (10 kPa bis) and the activity was measured. In the third experiment series, called "PHzS decrease", the test was done at decreasing hydrogen sulfide partial pressures (150, 100, 50, 25 kPa) without changing the catalyst and waiting 24 hours before modifying the operational conditions. After having measured the activity at 25 kPa the hydrogen sulfide partial pressure was increased back to 150 kPa (150 kPa bis) for a last series of products analyses. The HDO pathways of the model compounds have been established in a previous work (7). They are recalled in figure 1. Guaiacol can be demethylated to give catechol (breaking of the O-methyl bond) which can be transformed to phenol with the hydrogenolysis of the CAromatic-O bond (dehydroxylation); benzene, cyclohexene and cyclohexane can be obtained from phenol (dehydroxylation reaction). Phenol can also be directly formed from guaiacol by a demethoxylation reaction (hydrogenolysis of CAromatic-O bond). The ratio between the concentration of the products which have undergone the hydrogenolysis of the CAromatic-Obond (phenol, benzene, cyclohexene and cyclohexane) and catechol is calculated as follows: Ph
m
C phenol +
C b . . . . . . -I- C cycloh. . . . . -Jr- C cycloh. . . . .
cat -
Ccatechol
Ethyldecanoate can react following three pathways. The first [1] is the hydrogenation which produces decanol and ethanol, followed by a dehydration of decanol to give olefins (decene) which are subsequently hydrogenated to decane; the second [2] is the direct decarboxylation to produce nonane. The carboxylic acid can also be formed via a de-esterification reaction [3]; it is an intermediate product that can follow both the hydrogenation and decarboxylation pathways. The decarboxylation selectivity is calculated as the ratio between the concentration of nonane and the sum of the concentration of all the hydrogenated and decarboxylated products (nonane + decane + decene + decanol)" Sdecar b -_
C nonane C no. . . .
xl00
q- C decane "1- C decene q- C d. . . . ol
The carbonylic group of 4-methylacetophenone can be hydrogenated to the corresponding alcohol, a,4-dimethylbenzylalcohol, which can be dehydrated to 4-methylstyrene and finally hydrogenated to p-ethylmethylbenzene.
88
OH
OH
OOCH3
~OH
~
0
[
OH
{benzene cyclohexene cyclohexane
~ ~
guaiacol
catechol
ethanol C2H 5 - OH
decanol +
o II
C10H21 - OH
[11
C9H19 - C - OC2H 5
ethyldecanoate
[2]
phenol
dehydration
.---
decene C8H17 - H C ~
hydrogenation decane CH z
C10Hzz
hydrogenation ~ de-esterification [3]
O
II
C9H19 - C-- OH
decanoic a c i d ~ , ~ r b o x y l a t i o n C9H20
decarboxylation O
II
nonane
OH H3C
H3C
H3C
4-methylacetophenone p-methylstyrene a,4-dimethylbenzyl alcohol p-ethylmethylbenzene Figure 1. Hydrodeoxygenation pathways of guaiacol, ethyldecanoate, 4-methylacetophenone. 3. RESULTS The conversion data concerning GUA are reported in table 1; the figures in parentheses indicate the order in which the hydrogen sulfide partial pressure was increased (2nd series, namely "PHzS increase") or decreased (3rd series, namely "PHzS decrease") in the case of the experiments carried out with the same catalyst in different conditions. GUA conversion is not affected by hydrogen Table 1. GUA conversion. Reaction temperature = 270~ sulfide and similar values are obtained in the three H2S partial 1st series 2nd series 3rd series series of experiments; after pressure "PHzS increase" "PHzS decrease" about 300 and 200 hours of 21%(1) 10 kPa reaction, the conversion is 23% (7) 10 kPa bis unchanged (10 kPa and 10 19% (4) 19% (2) 25 kPa kPa bis, 150 kPa and 150 19% (3) 35 kPa kPa bis). 19% (4) 22% (3) 50 kPa 19% An important material 20% (5) 75 kPa 18% balance default has been 19% (2) 100 kPa 21% reported; the sum of the 21%(1) 20% (6) 150 kPa 17% product yields calculated on 19% (5) 150 kPa bis the basis of gaschromatograph analyses
89 corresponds to 50-60 % of the reactant conversion. Actually, when the liquid samples were withdrawn from the reactor, the presence of a white dense phase, which could not be identified, was been observed. The evolution of the yields in catechol 10 ~|m .. ~ - ~ ~ ~ ~ ~ ~-,"~~ Cat- P n2s and phenol (which ..~ Cat-P H2S inc also i n c l u d e the .... -~ Cat- P H2S dec yields in benzene, / 9 Cat -P H2S inc - 10 kPa bis c y c l o h e x e n e and .].~ -t::k~ ~ j, 9 Cat-P H2S dec-150 kPa bis cyclohexane) as a 5 __ 2;i1~~~.!~ ----O---Ph- P H2S function of PH2S is ~""~"~'" ----D---Ph-P H2S inc shown in figure 2. For the three series 2,5 ~ 9 ----A----Ph- P H2S dec of experiments [] Ph -P H2S inc - 10 kPa bis 0 I I I A Ph -P H2S dec - 150 kPa bis pheno1 y i e Id s decrease with the 0 50 100 150 increase of PH2S, while catechol yields P H2S (kPa) increase. In the case of the "PH 2S Figure 2. Evolution of catechol and phenol yields as a function of PH2S increase" series, the for the three series of experiments, catechol yields are higher and phenol ones are lower, compared to those obtained with the two other series of tests. A s a consequence of the decrease of phenol yields and increase of catechol ones, the phenol to catechol ratios decrease with the increase of PHzS (figure 3). Only in the case of the "PHzS increase" series, is the Ph/Cat ratio lower than 1. For the two other P H2S series of tests, phenol is produced in higher amounts than catechol, ---I--p H2S inc o 3 except at high hydrogen sulfide partial pressure. & P H2S dec 2 In the "PH2S increase" series, the Ph/Cat ratio after 300 hours of 9 P H2S inc reaction (10 kPa bis) is lower than 10 kPa bis the one observed at the beginning 9 P H2S dec of the test (10 kPa). This is due to a 150 kPa bis 0 higher catechol and a lower phenol 0 50 100 150 yield at 10 kPa bis than at 10 kPa. On the contrary, in the "PHzS P HZS (kPa) decrease" series, the Ph/Cat ratio after 200 hours of reaction (150 Figure 3. Evolution of the phenol to catechol ratio as a kPa bis) is higher than at 150 kPa; function of PHzS for the three series of experiments. in this case, the catechol yield is lower at 150 kPa bis than at 150 kPa. The results of the conversion of ED are reported in table 2. The conversion of ED is not affected by hydrogen sulfide; similar values are obtained for the three series of experiments. No differences are observed after 300 hours of reaction for the "PHaS increase" test (10 kPa and 10 kPa bis) and after 200 hours for the "PHaS decrease" test (150 kPa and 150 kPa bis). 1
90 Table 2. ED conversion. Reaction temperature - 270~ H2S partial 1st series 2nd series 3rd series pressure "PH2S increase . . . . PH2S decrease" 10 kPa 26% (1) 10 kPa bis 28% (7) 25 kPa 25% (2) 24% (4) 35 kPa 25% (3) 50 kPa 23% 25% (4) 25% (3) 75 kPa 23% 26% (5) 100 kPa 23% 24% (2) 150 kPa 23% 25% (6) 23% (1) 150 kPa bis 23% (5)
The sum of the product yields always corresponds to 95% (_+ 5%) of ED conversion, except for the "PH2S increase" series, where the sum of the product yields was about 80% of the r e a c t a n t c o n v e r s i o n . T h i s was probably due to a problem in the quantification of d e c a n o i c acid by gaschromatographic analyses. Decanoic acid yields are constant with the increase of the hydrogen sulfide partial pressure, but in the series "PHzS increase" they are much lower (0.02%) than in the two other series of tests (about 2.5%). The evolution of the yields in n o n a n e , 16 decanol and molecules "C9- PH2S containing ten atoms of C9 - PH2S inc carbon (C10), which 12 .A C9 - PH2S dec include both decane and decene, as a function of O C10 - PH2S PH2S, is represented in [] C10- PH2S inc --(D 8 figure 4. Nonane yields A ~z C10 - PH2S dec show a m a x i m u m at P H a S = 25-50 kPa, -- --0-- -decanol - PH2S 4 C10 yields increase, -- -El-- -decanol - PH2S inc while decanol yields -'-~-- -decanol - PH2S dec decrease w i t h the I I I increase of PH2S. The 0 50 100 150 sum of the yields of decane, d e c e n e and P H2S (kPa) decanol is constant. At PH2S = 25-50 kPa, a maximum of the Figure 4. Evolution of C9 (nonane), C10 (decane + decene) and decarboxylation decanol yields as a function of PH2S for the three series of selectivity is observed experiments. (figure 5). The products yields and the decarboxylation selectivity for the experiments done at 10 kPa bis and at 150 kPa bis are reported in table 3. Table 3. Nonane (C9), decane + decene (C10), decanol yields, decarboxylation selectivity, for the conversion of ED. nonane decane + decene decanol decarboxylation H2S partial Yield % Yield % Yield % Selectivity % pressure 10 kPa 10 kPa bis 150 kPa 150 kPa bis
4.7 6.3 4.7 3.4
9.2 9.6 12.7 13.9
7.1 6.8 1.1 1.0
22 28 26 19
91 In the "PHaS increase" " series, a higher nonane yield is observed at 10 kPa bis than at 10 kPa; this .o ,...~ results in a higher value for the 30 -" PH2S inc decarboxylation selectivity. In the o "PHaS decrease" series a lower --" PH2Sdec 20 nonane yield is observed at 150 kPa ~9 than at 150 kPa; the 9 PH2S incd e c a r b o x y l a t i o n selectivity is 10 kPa bis 10 consequently lower. 9 PH2S dec150 kPa bis The results for the conversion of MA are reported in table 4. The 0 50 100 150 hydrogenation of the carbonylic P HZS (kPa) group of MA is inhibited by h y d r o g e n sulfide. The results obtained in the three series of Figure 5. Evolution of the decarboxylation selectivity experiments are similar. In the as a function of PHzS for the three series of " P H 2 S d e c r e a s e " series the experiments. conversion is unchanged after 200 hours of reaction (150 kPa and 150 kPa bis). Table 4. MA conversion. Reaction temperature = 200~ The sum of the product yields always corresponds H2S partial 1st series 2nd series 3rd series pressure "PH2S increase . . . . PH2S decrease" to 95-100% (+ 5%) of the reactant conversion. 34% (4) 31% (2) 25 kPa T h e s e l e c t ivities in 28% (3) 35 kPa e t h y l m e t h y l b e n z e n e and 26% (3) 22% 23% (4) 50 kPa a,4-dimethylbenzylalcohol, 19% 23% (5) 75 kPa as a f u n c t i o n of the 18% (2) 19% 100 kPa hydrogen sulfide partial 12% (1) 11% 15% (6) 150 kPa pressure, are represented in 12% (5) 150 kPa bis figure 6. In the case of the series of tests performed with a fresh catalyst for each H2S partial pressure, the 100 evolution of the selectivity does not EtMeB show any clear tendency. P H2S "2- EtMeBFor the "PHaS increase" series, ~. 75 P H2S inc the ethylmethylbenzene selectivity .= EtMeBslightly decreases, while the a , 4 9- 50 P H2S dec dimethylbenzylalcohol selectivity is (D -- --O-- -MeBeA1 constant. In the case of the "PH2S ~D P H2S m 25 d ecrease" series the -- -m-- -MeBeA1 ethylmethylbenzene selectivity is mmmm-mm.- .ram. . . . 9 P H2S inc higher at higher H2S partial I I I -- --&-- -MeBeAI pressure, while the selectivity in P H2S dec 50 100 150 a,4-dimethylbenzylalcohol shows P H2S (kPa) the opposite behaviour. 40
--O--PH2S
A
Figure 6. Evolution of ethylmethylbenzene (EtMeB) and a,4-dimethylbenzylalcohol (MeBeA1) selectivity as a function of PHaS for the three series of experiments.
92 4. D I S C U S S I O N In what follows, we shall discuss the influence of the hydrogen sulfide partial pressure on the HDO of GUA, ED and MA over our CoMo/C catalyst. We will compare the results of this work with those obtained over alumina supported catalysts, as presented in previous papers (7, 20, 21). We will show that a good agreement exists between the different series of experiments. CoMo catalysts supported on alumina and on activated carbon show many similarities, but also some differences. The study performed on a neutral support, such as activated carbon, allows to distinguish the features which are due to the behaviour of the metal sulfides from those linked to the acid properties of alumina. The conversion of GUA initially involves the breaking of the O-methyl bond (demethylation) to give catechol. Then the reaction proceeds through the hydrogenolysis of the CAromatic-O bond (dehydroxylation) to phenol first and further to benzene, cyclohexene and cyclohexane (7, 23, 24). In this work, we have shown that, with the CoMo/C catalyst, the total GUA conversion is constant, but the formation of phenol, benzene, cyclohexene and cyclohexane decreases with the increase of the hydrogen sulfide partial pressure. Laurent et al. (21) have shown that on sulfided CoMo/AI203 and NiMo/Al203 catalysts, the breaking of the O-methyl bond was not affected by water or hydrogen sulfide, but was inhibited by ammonia. Alumina alone had some activity (about half that of the sulfided CoMo catalyst) for GUA conversion, giving catechol as single product. This result, combined with the fact that hydrogen sulfide had no influence on this reaction, while it generally has one on the hydrogenation and hydrogenolysis reactions occurring on the sulfided phase, indicated that acid sites were involved in the demethylation of GUA. It was concluded that the acid sites of alumina were mainly responsible. However, the higher activity of bimetallic sulfided catalysts suggested that the sulfided phase also played a role in the catalytic activity, probably through the acidity of the -SH groups. When alumina was replaced by activated carbon (9), the production of catechol decreased and a higher Ph/Cat ratio was observed; this increased as the total conversion increased. Compared to alumina, carbon gives catalysts which are less active for the breaking of the O-methyl bond, relatively to CAromatic-O bond hydrogenolysis. These results confirmed the role of acid sites in the demethylation reaction performed with alumina supported catalysts. But, at the same time, they showed that the sulfided phase also posses some activity in the demethylation of GUA. The dehydroxylation of catechol to give phenol (21) and the HDO of 4-methylphenol (20) over sulfided CoMo/A1203 and NiMo/A1203 catalysts decreased with the increase of hydrogen sulfide partial pressure. The dehydroxylation of 4-methylphenol to give toluene (hydrogenolysis of the CAromatic-O bond) was strongly inhibited by H2S. The hydrogenation of the aromatic ring was much less sensitive to H2S; it slightly decreased when the H2S concentration increased with the CoMo catalyst and it decreased over the NiMo one. Ammonia also had a strong inhibiting action, which was more marked for the dehydroxylation reaction than for the hydrogenation one. The inhibition by H2S in the dehydroxylation of GUA and methylphenol is in agreement with numerous data concerning C-S hydrogenolysis activity in HDS (19, 25). Such reactions are always more inhibited than hydrogenation ones. The cleavage of the CAromatic-O bond thus seems to correspond to a hydrogenolysis reaction. The different sensitivity to poisoning of the two reaction pathways suggests the presence of different active sites responsible for the dehydroxylation and for the hydrogenation of phenols. This has to be compared with the effect observed in HDS. It is generally accepted that the active sites in hydrotreatment catalysts are uncoordinated unsaturated molybdenum surface atoms and the distinction between hydrogenation and hydrogenolysis sites is thus likely to be found in the degree of uncoordination, the acidic character, the environment of the vacancy (for example sulfydryl groups) and/or the degree of reduction. The active sites are depleted in electrons and the electronic density and availability of the reactant molecules determine the adsorption. Laurent et al. (20, 21) have shown that hydrogenolysis sites have a higher affinity for electrons than hydrogenation sites. The surface sites associated with the molybdenum
93 sulfide phase can be unreduced, weakly reduced (partially uncoordinated), or strongly reduced (highly uncoordinated) molybdenum atoms, depending on the experimental conditions (H2/H2S ratio) and catalyst type. Unreduced sites would be inactive for adsorption and reaction (blocked sites), weakly reduced sites would be active for hydrogenation and strongly reduced sites would be active for hydrogenolysis reaction. Upon increasing the hydrogen sulfide partial pressure, the average degree of reduction of the active sites decreases and the hydrogenolysis of CAromatic-heteroatoms bonds is consequently inhibited. This accounts with the effects observed in our case for the hydrogenolysis of CAromatic-Obonds. In our work we have also observed that the Ph/Cat ratio after 300 hours of reaction for the "PHzS increase" series (10 kPa bis), is lower than the one measured at the beginning (10 kPa). It could be speculated that at increasing hydrogen sulfide partial pressure a progressive adsorption or reaction of hydrogen sulfide takes place on the hydrogenolysis sites; uncoordinated sites would be progressively transformed into blocked sites. When the hydrogen sulfide partial pressure is abruptly decreased, the hydrogenolysis activity cannot be restored rapidly. On the contrary, in the case of the "PHzS decrease", the Ph/Cat ratio after 200 hours of reaction (150 kPa bis) is higher than the initial one at 150 kPa. When the hydrogen sulfide partial pressure is decreased stepwise from 150 to 25 kPa, hydrogenolysis sites are progressively generated; the more reductive atmosphere (decrease of H2S/H2 ratio) could render them more stable. A longer exposure (compared to the duration of our experiment) would be necessary for observing the results of the re-increase of the hydrogen sulfide partial pressure. Two main reactions have been identified for the conversion of the carboxylic ester group: the hydrogenation-hydrogenolysis, which gives saturated alcohols and subsequently unsaturated and saturated hydrocarbons containing ten carbon atoms, and the decarboxylation which forms hydrocarbons with nine carbon atoms. Decanoic acid is formed by deesterification, it is an intermediary product and it can follow the hydrogenation as well as the decarboxylation pathway. Our results show that the ED conversion and the decanoic acid yields are constant, but the decarboxylation selectivity decreases with the increase of the hydrogen sulfide partial pressure. In a previous work (21), done with sulfided CoMo/AI203 and NiMo/AI203 catalysts, ammonia had a strong inhibiting effect on the conversion of carboxylic esters" the decarboxylation reaction was more affected than the hydrogenation one. No hydrogenation or decarboxylating activity of the support alone was detected (7). It was suggested that acidic (electrophilic) sites, located on the metal sulfides, were responsible for the reaction of the carboxylic group. The decarboxylation sites seemed to be characterised by a more acidic or electrophilic character. Hydrogen sulfide promoted the conversion of diethylsebacate, especially with NiMo active phase. Hydrogenation and decarboxylation reactions were both increased, with a more intense effect on the decarboxylation pathway. The promotional effect on the decarboxylation reaction was compared to the breaking of CAIiphatic-Nbonds in the presence of H2S. This effect is generally attributed to an increase of Bronsted acidity (an increase of the amount of protons available). The fact that the hydrogenation reaction was not poisoned by hydrogen sulfide was in contrast with the hydrogenation of aromatic or olefinic hydrocarbons and with the hydrogenolysis of CAromaticO bonds. Less specific uncoordinated molybdenum atoms were indicated as responsible for this reaction. In the case of our CoMo/C catalyst, the increase of the H2S partial pressure does not influence the conversion of ED. The hydrogenation pathway and the dehydration of decanol to give decene and decane are favoured by an increase of the H2S partial pressure, while the yields in decarboxylated products decrease at H2S partial pressure higher than 25 kPa; the decarboxylation selectivity consequently decreases. With the carbon supported catalysts the influence of hydrogen sulfide on ED decarboxylation and on GUA dehydroxylation is very similar. It seems that the two reactions take place on the same kind of active sites. The different behaviour of carbon and alumina supported catalysts could be due to their different acid properties. The acidity of alumina could partially alter the adsorption of the molecules. At
94 present we have no more elements to discuss and explain these differences. Some additional tests with the CoMo/C catalyst in the presence of ammonia could help to understand the details of acid-basic properties of the sulfided phase, without the influence of additional acidity apported by alumina. In the reaction of MA, the conversion of the carbonyl group proceeds through hydrogenation to a,4-dimethylbenzylalcohol, which is subsequently dehydrated to 4-methylstyrene (this intermediary product has never been detected, in our reaction conditions) and finally hydrogenated to 4-methylethylbenzene. In contrast to what was observed by Laurent et al. (7, 21), the rate limiting step in our reaction conditions is the dehydration of the alcohol to give olefins. MA conversion is the only one to be inhibited by hydrogen sulfide. In a former work (21), it was shown that on sulfided CoMo/AI203 and NiMo/A1203 catalysts, MA conversion was the only one not affected by ammonia; diethylsebacate and GUA were both strongly inhibited. Similarly to our results, hydrogen sulfide showed an inhibition action, more evident on NiMo than on CoMo catalysts. These results suggest that MA does not adsorb on electrophilic sites, as its conversion is not inhibited by ammonia, and that nucleophilic species (like a nucleophilic sulfur atom, or a hydride species) could be involved; the increase of hydrogen sulfide could affect the formation of these species and inhibit the conversion. As for the product selectivity, no clear tendency appears. For the "PH2S decrease" series, it seems that an increase of the hydrogen sulfide partial pressure would favour the conversion of the alcohol. 5. C O N C L U S I O N S Hydrogen sulfide does not affect the activity of the CoMo/C catalyst for the overall conversion of GUA. As the products are concerned, hydrogen sulfide inhibits the direct hydrogenolysis of the CAromatic-O bond, the phenol to catechol ratio is consequently decreased. These results are in good agreement with those obtained over alumina supported catalysts, for the HDO of 4-methylphenol and GUA. Acid sites situated on both alumina and sulfided phase (-SH) groups seem to be involved in the demethylation of GUA. Highly reduced uncoordinated sites, located on the metal sulfides, would be responsible for the hydrogenolysis reaction. The decrease of the H2/H2S ratio would decrease the average degree of reduction of the active sites; phenol production (hydrogenolysis of CAromatic-O bond) is consequently inhibited. As for ED, the conversion is not inhibited by hydrogen sulfide, but the decarboxylation to nonane is decreased, while the formation of hydrogenated products is favoured. On our carbon supported catalyst the influence of hydrogen sulfide on GUA dehydroxylation and ED decarboxylation is quite similar. The two reactions seem to take place on the same kind of active sites. MA conversion is the only reactant molecule to be inhibited by hydrogen sulfide. In a previous work, performed with alumina supported catalyst, the same effect was observed; at the same time, ammonia did not show any inhibition action. These results suggest that nucleophilic species (like a nuc|eophilic sulfur atom, or a hydride species) could be involved in the HDO of MA; the increase of hydrogen sulfide partial pressure could affect the formation of these species and inhibit the reaction. ACKNOWLEDGEMENTS The financial support of the European Union (Contract No. JOR3-CT95-0025) is gratefully acknowledged.
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
97
PERFORMANCE OF NOBLE METAL-Mo/~'-AI203 CATALYSTS: EFFECT OF PREPARATION PARAMETERS M.H. Pinzfn, L.I. Merifio, A. Centeno and S.A. Giraldo Centro de Investigaciones en Catalisis, Escuela de Ingenieria Quimica, Universidad Industrial de Santander. A.A. 678, Bucaramanga, Colombia. e-mail:
[email protected] Abstract
The catalytic performance of bimetallic catalysts NM-Mo (NM = Pt, Pd, Ru) supported on ~/-alumina-was evaluated in simultaneous hydrodresulfurization (HDS) and hydrogenation (HYD) reactions. Nature of noble metal, calcination temperature after noble metal impregnation, activating mixture composition and the effect of the impregnation of noble metal on MoS2/~,-A1203 were the parameters studied. Results show that the noble metal nature and impregnation of noble metal on MoS:h/-AI203 are the principal factors that influence the catalytic performance. PtMo/'y'-AI203 and RuMo/7'-A1203 display a high synergetic effect. Catalysts prepared impregnating noble metal on MoS2h/-AI203 present higher activities than the corresponding ones prepared impregnating noble metal on MoO3/y-A1203. Calcination temperature after noble metal impregnation does not have an important influence on catalytic activity. 1. INTRODUCTION New catalytic systems for hydrotreatment (HDT) based on the use of noble metal (NM) sulphides as active phases have been recently proposed [1-3]. It is known that noble metals have hydrodesulfurizing properties and excellent hydrogenating properties [1,4], but until now, it is not clear how these systems work. Finding the best activating and preparation conditions for these new catalytic systems, becomes a promissory alternative to be explored, in order to obtain catalysts for deep HDT. The best activating conditions for the traditional CoMo or NiW catalysts for HDT reactions have been studied and reported in the literature [5,6]. But these activating conditions do not necessarily correspond to those of catalytic systems containing others active phases. Literature reports differences in the activating conditions for catalysts containing noble metals compared with those used for the CoMo or NiW systems, e.g. different temperatures and different ratios of activating mixtures are necessary to activate catalysts containing ruthenium [2,3]. In this paper the results of the catalytic performance of bimetallic catalysts NM-Mo (NM = Ru, Pt, Pd) supported on alumina in HDS of dibenzothiophene and HYD of naphtalene are presented. The parameters studied were the following: nature of the noble metal, calcination temperature after noble metal impregnation, activating mixture composition and the effect of the impregnation of noble metal on MoS2/y-AI203. A temperature-programmed reduction (TPR) was made in order to characterise the catalysts.
98 2. EXPERIMENTAL 2.1. Catalysts Two series of catalysts supported on alumina were prepared. In the first one, wet impregnation of the MoO3/y-AI203 with a aqueous solution of the corresponding noble metal salt was used and in the second one the wet impregnation was made on the MoS2/y-AI203. The MOO3(15 wt.%)/y-Al203 was prepared by dry impregnation of the alumina (Procatalyse, BET surface area = 244 g m 2 and pore volume = 0.55 cm3.g1) with an aqueous solution of ammonium heptamolybdate (Merck). The preparation procedure was as follows: the alumina was previously calcined in air at 773 K. ARer the impregnation a weight of approximately 10 g of this solid was dried under a flow of air (100 ml.min-1) at 393 K for 12 h. Then it was calcined in air at 773 K for 4 h. This part of the preparation procedure was the same for both series. The MoO3/~/-AI203 obtained was impregnate directly with the corresponding noble metal salt (first series) or it was sulphided before the noble metal impregnation (second series) using a standard reduction-sulphidation procedure [5]. Wet impregnation was made with a volume 6 times the pore volume of an aqueous solution of the noble metal salt: PdCI2, H2PtCI64.5H20 or RuC130.5H20 (Sigma). Different temperatures were used to calcinate catalysts of first series at~er noble metal impregnation, namely 523 K, 623 K and 723 K. Catalysts without this second calcination were also tested. Catalysts of second series have not been calcinated aider noble metal impregnation. The composition of the catalysts for the two series was 0.5 wt.% of noble metal and 15 wt.% ofMoO3. A conventional CoMo/7-m1203 (15 wt.% MOO3, 3 wt.% CoO) catalyst was also prepared for comparison purposes. 2.2. Activating procedure for the catalytic tests The catalysts were activated in situ (TPR apparatus or catatest). In general the activating procedure was the same as the standard one described elsewhere [5] but now activating mixture composition was changed. Three different activating mixtures were used: 15 vol.% H2S in N2, 15 vol.% HzS in I,/2 and pure H2S. H2S/N2, H2S/H2 and H2S will denote them respectively. After activation and before the catalytic test, catalysts were maintained in the reactor under the corresponding activating mixture. 2.3. TPR measurements The TPR measurements of activated catalysts were made in a ChernBet 3000 of Quantachrome working with a TCD detector. After reactor and before detector, outlet gases were cooled at 210 K with a solution of ethanol-solid CO2 in order to trap any trace of H2S and H20, which might be formed during reduction. The TPR procedure used was as follows: 10 mg of catalyst, in particles ranging from 0.3 mm to 0.6 mm were put into a U shaped quartz reactor used for the activation and TPR measurements. The samples were activated following the same activating procedures used for the catalytic test described before in 2.2. The activating temperature was always maintained for 1 h. The samples were cooled under the corresponding activating mixture until 383 K. They were then flushed with 1',I2 (10O ml.min-1) at 383 K for 30 rain and then cooled until room temperature. At~erward, N2 was replaced by the reducing mixture (15 ml.min1) of 5 vol.% H2 in N2. The temperature was then 1 increased at a rate of 20 K.min" until 1123 K and the hydrogen consumption was registered. 2.4. Reaction system Catalytic tests were made in a fixed-bed, high-pressure flow reactor at 5 MPa and 583 K working in steady state. The model reactants were dibenzothiophene (2 wt.%), nat~halene (3
99 wt.%) solved in cyclohexane and hexadecane (2 wt.%) as internal standard for the chromatographic analysis. A liquid flow rate of 30 ml.hq and an H2 (SPT)/liquid volume ratio of 500 were used. 0.5 g of catalyst (0.3-0.5 ram) diluted with glass spheres (1 ram) was used for the entire catalytic tests. The absence of any diffusion effects was experimentally verified by showing that similar conversions, as a function of contact time, were obtained for two different weights of catalysts. 2.5. Analysis The liquid samples were analysed in a HP Model 6890 gas Chromatograph, equipped with a capillary column (phenyl xyloxane 5%) of 30 rn, a split injector and a flame ionisation detector. 2.6. Expression of the results HDS and HYD catalytic activities were expressed both by the respective total conversion of dibenzothiophene (%CHos) and nafthalene ( % C ~ ) after steady state was obtained (6 h on stream).
3. RESULTS 3.1. TPR measurements In most of the cases the TPR profiles present a well-defined peak that correspond to hydrogen consumption. This peak is present at a temperature range of 550 to 750 K. Typical figures showing results are presented below. Figure 1 presents the TPR profiles corresponding to PdMo/y-Al203, PtMo/y-A1203, RuMo/y-A1203, CoMoh/-Al203 and Mo/7-AI203 catalysts when H2S/N2 activating mixture was used. As observed in Figure 1, the peak of the Moh/-Al203 is modified by the presence of the noble metal. The nature of the metal is involved. Hydrogen consumption is different for each catalyst. The area of this peak for the PtMo/y-AI203 catalyst is the biggest and that of the CoMo/y-Al203 is the smallest. It is possible to point out an increasing order of temperatures: Mo < CoMo ~ PtMo < RuMo
100
mixture. A bump at high temperature is present in these three TPR profiles. If TPR profiles of catalysts of the first series (d, e and f) are now considered, we observe similar shift of peak temperature. Nevertheless, the peaks from the second series catalysts are presents at higher temperature than those ones from the first series catalysts. The bumps disappear in the cases of catalysts of the first series.
3.2. Catalytic performance Table 1 presents the catalytic performance for lIDS and HYD of bimetallic catalysts NMMo/7-A1203 of first and second series and of the reference catalysts (CoMo/~/-AI203 and Mo/7-AI203) when different activating mixtures were used. In general the catalysts prepared by impregnation of noble metal on MoS2/7-AI203 present a higher activity for both HDS and HYD than those ones of the corresponding catalysts prepared by impregnation on the MoO3/3t-A1203. The results presented in b Table 1 show that for each noble metal the f ~ 7"--x.-~_.. trends are the same in both series. The PtMo/7-AI203 catalyst is among the NMD Mo/7-AI203 catalysts the most active for (g both HDS and HYD reactions. The other noble metal bimetallic catalysts (RuMo/~,A1203 and PdMo/7-AI203) present lower conversions than the reference catalyst d CoMo/3t-AI203. Based on results showed in Table 1 Ill it is possible to point out the following: The activating mixture influences in ~n' ~' 7'oo ' oo' ' different ways the performance of the Tempemt~e catalysts. This influence is related with the nature of the noble metal contained in the Figure 1. TPR profiles when H2S/N2activating catalyst. mixture is used. a) PdMo/7-A1203, b) PtMo/7For ~Moh/-Al:O3 catalyst, activation A1203, c) RuMo/'y-A1203, d) CoMo/'y-AI203 without H: in activating mixture increase and e) ~do/~/-Al203. the catalytic activity for both HDS and HYD reactions compared to the use of the traditional H2S/H2 mixture. For the PtMo/7-A1203 catalyst the HYD is the most favoured. Use of H2S/N2 activating mixture gives the highest activities in this case. For RuMoh/-Al203 the use of different activating mixtures does not have an important effect on the catalytic performance. For PdMo/7-AI203 catalyst the activating mixtures without H2 increase the catalytic activity for both HDS and I-PfD compared with those when the H2S/H2 mixture was used. For this catalyst the highest activities are obtained when pure H2S is used. ...... "1
t--
m
-
-
d
-
101
For Mo/y-Al203 and CoMo/y-AI203 catalysts the best catalytic performance is obtained when activating mixture is the traditional H2S/H2. Table 2 presents the results of the activity of NM-Mo/y-Al203 catalysts of the first series calcinated at different temperatures after noble metal impregnation and non-calcinated. These catalysts were activated using H2S/N2. As observed in Table 2, calcination temperature after noble metal impregnation does not have an important influence on activity for both HDS and HYD. Nevertheless, some tendencies are observed: PdMo/y-AI203 and RuMo/y-AI203 catalysts show slight increment in HDS activity when calcinaton temperature decreases and PLMo/3r-AI203 increases activity in HDS when the calcination temperature increases. The best HDS activity for PdMo/y-Al2Os and RtLMo/yA1203 catalysts is when they are not calcinated after the noble metal =i impregnation. Hydrogenation does not present a clear tendency in any of the cases. s
o-__Y
v
f
l--
4. DISCUSSION We will discuss the imquence of noble metal nature, the effect of different activating mixtures, the effect of calcination temperature after noble metal impregnation and the effect of the noble metal impregnation on both HDS and HYD reactions. We try to interpret TPR results.
d
e
f
oo'
i
8113
I
,
I
900
I
4.1. T P R interpretation
All the TPR profiles show a welldefined peak in the range of 550 to 750 K. This well-defined peak corresponds to the partial reduction of molybdenum sulphide Figure 2. TPR profiles PtMo/~t-Al203 [9]. It is attributed to non-stoichiometric catalysts: (a), (b), (e) Pt impregnated on species located on the surface of micro MoS2/~/-Al203, and (d), (e), (f) Pt impregnated crystals active in HDS [10]. Reduction on MoOJT-AI203. (a) and (d) activated with temperature of this non-stoichiometic H2S/N2, 03) and (e) with H2S, (r and (0 with species is reported to correlate with H2S/H2. activity for HDS [10]. The TPR profiles of Mo/7-A1203 catalyst show that the temperature of this peak changes depending on the activating mixture used. That means that the formation of different non-stoiehiometrie species on the surface of crystals depends of the activating mixture. It is observed that there are displacements of the temperature of the peak in most of the cases of bimetallic catalysts when
102
different activating mixtures were used. These changes in temperature may explain the differences in catalytic activity for some of the cases. We think that maybe the temperature of the peak correlate with catalytic activity in a limited temperature range and that a maximum of activity could be obtained in this interval depending of the noble metal nature and its content in the catalyst [ 1,2]. We interpret this fact considering that activating conditions and noble metal nature change the statistical distribution of species. We believe that not all these species have the same participation in the catalytic phenomenon during HDS and HYD. Our results show that in the case of PtMo/3,-Al203 and RuMo/~t-Al203 catalysts the conditions used favoured the most active species. This hypothesis might explain the synergetic effect that was found. In the case of PdMo/~t-Al203 catalyst we think that activating conditions are not the best to generate the most appropriate species. The presence of the bumps at high temperature, in some of the TPR profiles, confirms the idea that activating conditions influence the formation of different sulphide compounds. Table 1
Catalytic performance in HDS and HYD of NM-Mo/~,-AI203 (first and second series) and of the reference (CoMo/y-A1203 and Mo/~,-A1203) catalysts activated with different activating mixtures at 673 K. i i First series* Second series Impre~nated on MoO~J~'-A1203 Imprej~nated on MoS2/y-AI203 Catalysts/7-A1203 (Activating mixture) 0~CHDs O~CHYD 0~CHDs ~ D i 16.0 4.0 Mo (H2S/H2) 9.0 4.0 Mo (H2S/-N2) 65.0 21.0 CoMo (H2S/H2) 52.0 16.0 CoMo (H2S/N2) 52.0 25.0 37.0 11.0 PtMo (H2S/H2) 64.0 41.0 35.0 14.0 PtMo (H:S/N:) 62.0 29.0 PtMo (H2S) 10.0 9.0 PdMo (H2S/H2) 12.0 13.0 10.0 5.0 PdMo (H2S/N2) 16.0 18.0 PdMo (H2S) 24.0 15.0 RuMo (H2S/H2) 22.0 17.0 10.0 6.0 guMo (H2S/N2) 24.0 12.0 i RuMo 0-I~S) i i i , , * Calcination temperature- 723 K. 4.2. Influence of noble metal nature
We discuss the results for the catalysts corresponding to the second series. Tendencies are the same for the first series. It is clear that nature of noble metal is the most important factor that influences the catalytic performance in the HDS of dibenzothiophene and HYD of naithalene. There are important differences in the catalytic behaviour among the three noble metal bimetallic catalysts.
103
Table 2 Catalytic activity for HDS and HYD of N]VI-Mo/y-AI203 catalysts of the first series. Activat!n~ mixture: H2S/N2 at 673 K. Calcination Temperature (K) Non Calcinated. 523 623 723 Catalysts %CribS ~ O~CHDs ~ 0~CHDs ~ ~ 0~CHYD PtMo 29.0 11.0 26.0 12.0 30.0 14.0 35.0 14.0 PdMo 14.0 7.0 13.0 4.0 11.0 8.0 10.0 5.0 RuMo 14.0 3.0 11.0 4.0 8.0 6.0 10.0 6.0 ,
,
,
i
,
i
i
Catalyst containing platinum increases the catalytic activity for HDS of molybdenum sulphide and promotes an important increment in its HYD activity. We have reported for Pt/7A1203 catalyst activated with H2S/N2, conversions at the same reaction conditions, of 14 % and 17 % for HDS of dibenzothiophene and HYD of naf~halene respectively [8]. We point out that platinum concentration in the catalyst is only 0.5 wt.% compared to the typical 3 wt.% of Co in the traditional CoMo/~,-AI203 catalyst. If we take into account that the Pt/Co atomic weight ratio is more than 4, we conclude that the effect per atom of platinum is much higher than that corresponding to cobalt. This synergetic effect found in our platinum-molybdenum catalyst agrees with the results reported in literature for model catalysts [ 1,7]. Palladium does not have an influence for both HDS and HYD reactions. In this case a synergetic effect is not present. Similar conversion to that of the Mo/7-AI203 catalyst is obtained. This behaviour is in agreement with results reported in the literature [1]. Ruthenium slightly increases the HDS activity of the molybdenum sulphide but it promotes an important increament in its hydrogenation capacity. A synergetic effect, especially for HYD, is also present when this bimetallic catalyst is tested. We have reported [8] very low activity for Ru for both, HDS and HYD in the conditions of our catalytic tests. We interpret our results considering that each noble metal sulphides have different capacity to activate hydrogen. It is possible that "spillover" hydrogen might be involved [ 1,9]. 4.3. Influence of the activating mixture Our results show that use of different activating mixtures influences the catalytic performance in different way for each bimetallic catalyst. The influence of this parameter on catalytic activity is more important in the case of the PtMo/y-Al203 catalyst for both HDS and HYD reactions. Activity for PtMo/y-Al203 catalyst increases when it was activated with H2S/N2 or pure H2S if compared with that corresponding to the activation with the traditional H2S/H2. Activity for the HDS and HYD reactions are not influenced significantly by the activating mixture in the case of RuMo/7-AI203 and PdMoh/-Al203 catalysts. Nevertheless, the activities tend, in the case of P@clo/y-A1203 catalyst, to change slightly when activating mixtures without H2 are used. Our results show that in the case of bimetallic catalysts containing noble metal the catalytic behaviour and activating conditions are related, but also show that the best activating conditions do not correspond to those reported for the CoMo/y-Al203 catalyst [5]. We have interpreted this fact thinking that the formations of sulphide active phases are related with the H2S/H2 ratio at equilibrium [ 11]. We know that depending on this H2S/H2 ratio and on the others activating conditions, namely temperature, different kinds of species might be formed.
104
The sulphidation degree of each metal is also influenced by these conditions. Literature reports that different degrees of sulphidation are obtained for each transition metal when the same activating conditions are used [11]. This last fact might explain one part of the difference in the catalytic behaviothr between the different tested catalysts. We speculate that it is possible that the presence of the noble metal in bimetallic catalysts might decompose the H2S and activate H2 and that it changes the H2S/H2 ratio on the surface of the catalyst during the activation. 4.4. Influence of the impregnation procedure Impregnation of noble metal on MoO3/~/-A1203 or on MoS2h/-A1203 is an important factor that influence catalytic behaviour. All catalysts prepared impregnating noble metal on MoS2/~'-Ai203 present better activities than the corresponding ones impregnated on MoO3/~/Al203 (Table 1). We interpret this fact like the result of two possible contributions. A good dispersion of noble metal sulphide on the surface of alumina and a direct effect of noble metal during the activation. If we consider that MoS2/~/-Ai203 is well dispersed on alumina, it is probable that the noble metal sulphide can be well dispersed in the flee sites of the alumina surface resulting after sulphidation of molybdenum. Literature reports that strong interaction are possible between Ru and alumina support [2]. In the case where noble metal is impregnated on MoO3/y-Al203 we think that interaction between molybdenum oxide and noble metal salts are not strong to avoid sintering in the next stage of the activating procedure. It is also possible that noble metal influences the formation of species on the surface of the MoS2 micro-crystals. This interpretation agrees with results showing that different TPR profiles are obtained when different noble metals were used. ACKNOWLEDGMENTS
This work was made thanks to the financial support from COLCIENCIAS, a government institution that promotes science in Colombia, in the frame of the project: 1102-08-271-94. REFERENCES
1. S.A. Giraldo de Le6n, P. Grange, B. Delmon, in "New Aspects of Spillover Effect in Catalysis", (T. Inui, K. et al. Eds.), Studies in Surface Science and Catalysis, Vol. 77, Elsevier, Amsterdam, 1993, p. 345. 2. C.E. Scott et al. in "XVI Simposio Iberoamericano de Cat/disis" (A. Centeno, S.A. Giraldo, E.A.Paez Mozo Eds.) Vol. 1, Cartagena de Indias (Colombia) 1998. 3. C. Creantet, J.A. De Los Reyes, M. Vrinat, M. Breysse, in "JECAT'91", (J. Hiraishi, Chairman; A. Nishijima et al. Eds.), JITA and PATE, Tokyo, 1991. 4. S.A. Giraldo, P. Grange, B. Delmon, Appl. Catal. A 107 (1993) 101. 5. R. Prada Silvy, J.L.G. Fierro, P. Grange, B. Delmon, in "Preparation of Catalysts IV" 03. Delmon, P. Grange, P.A. Jacobs, G. Poncelet Eds.), Elsevier, Amsterdam, 1987, p. 605. 6. B. Scheffer, P.J. Mangnus, A. Moulijn, J. Catal. 121 (1990) 18. 7. S.A. Giraldo et. al, in "XIV Simposio Iberoamericano de Cat/disis", (P. Reyes Nunes, R. Cid Araneda Eds.), Soc. Chilena de Quimica, Vol. 3, p. 1357. Concepci6n (Chile) 1994. 8. L.I. Merifio, A. Centeno, and S.A. Giraldo, Appl. Catal. Submitted for publication. 9. S.A. Giraldo, P. Grange, B. Delmon, Catal. Lett. 47 (1997) 51. 10. Portela, P. Grange, B. Delmon, J. Catal. 156 (1995) 243. 11. P.J. Mangnus, A. Riezebos, A.D. van Langeveld, J.A. Moulijn, J. Catal. 151 (1995) 178.L.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
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Use of ammonium tetrathiomolybdate as a new precursor for the preparation of hydrodesulfurization catalysts by a sol-gel method L. Le Bihan, C. Mauchauss6, E. Payen and J. Grimblot Laboratoire de catalyse h6t6rog~ne et homog~ne, URA CNRS 402, Universit6 des Sciences et Technologies de Lille, 59655 Villeneuve d'Ascq C6dex, France Abstract MoS2-alumina catalysts, with various amounts of Mo, were prepared by a sol-gel method using aluminium-tri-sec-butylate (ASB) in 2-butanol and ammonium tetrathiomolybdate (ATTM). After hydrolysis, the dried gels were sulfided and tested in hydrodesulfurization (HDS) of thiophene (T) or dibenzothiophene (DBT). This method of preparation allows to increase the DBT conversion by reference to results obtained on classically sulfided solgel oxides.
1. INTRODUCTION MoS2-based catalysts promoted by Co or Ni are largely used to clean petroleum feeds that contain heteroatoms such as S, N, and O. These catalysts are generally obtained by sulfidation of an oxidic precursor CoO(NiO)-MoOs/AI203, synthesized by impregnation of an alumina support with solutions containing ammonium heptamolybdate (AHM) and Co or Ni nitrate (coimpregnation or successive impregnation). It is well known that the oxidic precursor features and the activation conditions influence the performances of the nanocrystallites of the MoS2 active phase well dispersed on the support [1, 2]. Even if some Mo-O-A1 bonds remain in the sulfided catalysts [3, 4], the exact nature of the interaction between the active phase and the support and the orientation of its basal plane relative to the support are not well established; they seem to depend on the sulfidation temperature [5]. On the contrary, it is well accepted that the active sites are located on the edge of the MoS2 slabs and correspond to coordinative unsaturations created by the removal of labile S atoms [1, 2] and their number is influenced by the MoS2 morphology. Indeed Kasztelan et al [6]
106
showed a direct correlation between the size of the crystallites and the HDS activity. On the other hand, Daage et al [7] established, through the study of bulk MoS2, the importance of the layer stacking on the HDS and hydrogenation (HYD) selectivity during DBT conversion. The precise control of the MoS2 layer stacking appears, therefore, as a means to control catalyst activity and selectivity. In previous works [8, 9] we reported a sol-gel procedure to prepare oxide precursors. Alumina was obtained by hydrolysis of ASB and AHM was incorporated during the alumina gel synthesis This procedure allows to increase the Mo loading up to 30 wt% Mo with preservation of the good Mo dispersion and to increase the activity per gram of catalyst in HDS of T. High Resolution Electronic Microscopy (HREM) and X-Ray Diffraction (XRD) studies showed that the active phase of these catalysts also consists of well dispersed MoS2 nanocrystallites as those present on the classical catalysts. The aim of the present work was to use ATTM as a starting material during the sol-gel procedure to further increase the catalyst performances. However the low solubility in water of ATTM precludes valuable preparation through the conventional incipient wetness impregnation. Indeed impregnation of alumina with ATTM solutions has already been proposed and has not given interesting results [10, 12]. In the present study, ATTM is therefore incorporated directly in the alumina gel which allows to overcome the problem of solubility. 2. E X P E R I M E N T A L 2.1. S a m p l e s p r e p a r a t i o n and a c t i v a t i o n ATTM has been prepared by sulfidation of an ammonia solution containing AHM [13] The procedure of preparation of the catalysts has been already reported [9]. The corresponding catalysts will be referred as XA1T, where X represents the Mo loading in wt% of Mo, A1 the alumina obtained by the sol-gel method and T the ATTM. These solids were compared with those obtained by incorporation of AHM in the gel. These latter ones, which were characterized in previous works [8, 9], will be referred hereafter as XA1H with H for AHM. The sulfidation by a H2/H2S (90/10) mixture at 400~ for 3 h was carried out on the dried XA1T and on the calcined XA1H samples. For comparison purposes, bulk MoS2 has been prepared by sulfidation of ATTM in the same conditions. 2.2. C h a r a c t e r i z a t i o n s and H D S t e s t s The chemical composition was determined by CNRS (Vernaison, France) and the texture of the sulfided solids, previously degassed at 300~
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(2 h), was determined by N2 adsorption at 77.5 K. Procedures of characterization by XRD, HREM and XPS have already been described [8, 9]. In particular, XPS provides results about surface composition, extent of Mo sulfidation and identification of the Mo chemical states (MovI in MoO3 or in MoS42-, Mo v in oxysulfides or Mo TMin MoS2) [14-16]. HDS of T was conducted at atmospheric pressure in a flow type reactor packed with 200 mg of catalyst. The solids were sulfided with a H2/H2S (90/10) mixture at a flow rate of 100 cc/min at 400~ for 2 h and then cooled down to 300~ After purification by vaccuum distillation, T was introduced in the reactor at constant pressure (50 torr) in a flow of pure H2 (20 cc/min). The reaction products were analyzed by gas chromatography. HDS of DBT was carried out at the Exxon Research Laboratory (Annandale, N. J., USA) with a procedure described in [7]. 3. RESULTS AND DISCUSSION 3.1. D r i e d s o l i d s The XA1T solids have been characterized by XPS and XRD before sulfidation in order to check the preservation of the ATTM upon its introduction in the alumina gel. Sample 46A1T has a pattern similar to that of ATTM whereas samples 37A1T and 33A1T exhibit the ATTM peaks on broad underlying peaks which characterize the formation of an amorphous phase similar to the one observed for the dried XA1H solids [8]. Incorporation of ATTM during the alumina preparation prevents thus formation of well crystallized boehmite which is observed without introduction of ATTM. Width, intensity and position of the diffraction peaks indicate the presence of a weakly crystallized A1 oxyhydroxyde. Sharp peaks corresponding to well crystallized compounds are also observed at low Mo loading but their origin is not yet understood. However, by taking into account the sensitivity of XRD to crystalline compounds compared to amorphous oxyhydroxyde, we can consider that this crystalline phase is only present as a minor component. The XRD features os ammonium oxothiomolybdates, ammonium sulfate or A1 sulfate have not been observed. Presence of Mo vI oxo-species (MoVr)ox is evidenced by XPS [Mo 3d binding energy (BE) = 236.3 and 233.3 eV] at low Mo loading whereas at high Mo loading the Mo 3d lines of ATTM (BE = 233.6 and 230.4 eV) are detected. The Mo 3d features of ATTM are not observed on the spectrum of the 13A1T sample, which suggests that ATTM is transformed upon its introduction in the gel. A new doublet characteristic of Mo TM sulfided species (MoIV)sulf a s in MoS2 or in (NH4)2(MoaSla).nH20 [16] (BE = 232.1 and 229.1eV) appears ; it is not possible to distinguish between these two species by XPS but both can result from the decomposition of ATTM. A doublet
108
inserted between these three XPS features is also present, and is generally ascribed to (( oxysulfides )) species [ 15, 16]. The amounts of the various types of Mo species present in the samples have been determined by integration of the XPS peaks and are reported in Table 1. It appears that the total amount of the oxide phase increases upon decreasing the Mo loading. This is confirmed by the XPS S2p peaks. Indeed, the intensity ratios ($1/$2) of the peaks corresponding to sulfates ($1) and sulfides ($2) also decrease upon increasing the Mo loading. It can therefore be suggested that oxidation occurs at low Mo loading. It may occur during the preparation or in air after transfer of the dried samples. At high Mo content, large ATTM crystallites dominate, as evidenced by XRD. As they are not totally visible by XPS a non uniform evolution of their abundance with the Mo loading is therefore obtained. Table 1 Quantitative XPS analysis of the dried XA1T solids wt% Mo % (MoVr)~u~ % oxysulfides % (MoVI)ox 13 (a) 25 59 18 26 42 32 33 57 43 37 76 15 9 46 76 24 (a) 16% was attributed to the presence of (Mo~V)~,lfspecies
82/81 1.1 0.95 0.83 0.35 0.23
3.2. Sulfided Solids The XRD pattern of the sulfided XA1T solids without Mo shows the features of an amorphous alumina [8] with no evidence of the presence of A1 sulfide. When Mo is introduced, the XRD features of MoS2 are observed on broad underlying peaks characteristic of rather amorphous alumina. Broadening of the (200) peak at 20 = 14 ~ is observed upon decreasing the Mo loading. From its full width at haft maximum (FWHM), the layer stacking of the MoS2 slabs has been determined (Table 2) and are compared with the sulfided 30AIH solid. It appears that the mean number of layers increases with the Mo loading and for a similar Mo content, the MoS2 stacking in solids prepared with ATTM is higher than the one of solids prepared with AHM. The mean number of slabs deduced from the XRD measurements do not take into account the single slab crystallites and therefore correspond to the upper limit of the actual mean stacking taking into account the MoS2 nodules when they exist.
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Table 2 Mean layer stacking (n) of the MoS2 crystallites in sulfided XA1T and XA1H samples [from the FWHM of the (200) diffraction peak] Sample 13A1T 18A1T 33A1T 37A1T 46A1T 30AIH n 3.2 3.7 4.8 4.2 5.5 1.1
Concerning the HREM observations, whatever the Mo loading, the samples are heterogeneous which precludes any statistical analysis. At low Mo loading, the pictures show the presence of both well dispersed single and stacked slabs which proportion depends on the Mo content. At high Mo loading, stacked MoS2 crystallites of various lengths are mainly present. Their length is however smaller than that observed on bulk MoS2. Some MoS2 nodules are also evidenced at very high Mo content. The Mo 3d XPS spectra exhibit mainly the Mo TM features of MoS2. with a weak contribution of (MoVl)ox; t h e remaining part (less than 25 %) is due to the oxysulfide phase. The sulfidation degree is close to 90%, which corresponds to higher values than those obtained for the XA1H series. Whatever the Mo loading, the S/Mo atomic ratios obtained by XPS or by elemental analysis are always slightly higher than 2 and are ascribed to the small size of the MoS2 particles [6]. From the variation of the intensity ratio IMo]IA1versus the bulk atomic ratio nMo/nA1for the XA1T samples it appears that the limit of good dispersion is reached for a Mo content of 13 wt% Mo. At higher loadings, MoS2 is therefore not present as a monolayer at the surface of the support, but is present as crystallites which have been evidenced by HREM. The specific surface area (SSA) of the solids after sulfidation are reported in Table 3. After correction of the weight contribution of MoS2 contained in the solids, these values are quite similar and should correspond to the SSA of alumina. Alumina plays mainly the role to disperse the MoS2 phase, even at high Mo contents. For catalysts prepared with ATTM solutions (incipient wetness impregnation or equilibrium adsorption) [10-12], it was noted a very low Mo dispersion with a high degree of sulfidation which is the consequence of the absence of Mo-O-AI bonds. The method of preparation used in this work allows to increase the dispersion up to values similar to the one obtained by dry impregnation of alumina with AHM solutions. However the stacking of the MoS2 crystallites deduced from XRD measurements is higher than for these latter ones. This is probably due to the presence of a residual amount of ATTM in these dried solids. Upon increasing the Mo content, the MoS~ layer stacking increases. This can be correlated to the preservation on the dried solids of the ATTM well dispersed in the alumina precursor and which
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Table 3 9Specifi c surfac e area (SSA) of the sulfided XA1T catalyst s . Sample wt % MoS2 SSA (m2.~9 13XA1T 22 617 18XA1T 31 363 33XA1T 55 151 37XA1T 62 179 46XA1T 77 120 MoS2 100 18 transforms upon sulfidation into MoS2 crystallites distributed in alumina. 3.3. C a t a l y t i c r e s u l t s
Figure 1 shows the HDS activity of T versus the Mo loading. The best activities are obtained for the samples 13A1T (22 wt% MoS2) and 30A1H (45 wt% MoS2), these Mo loadings corresponding to the limit of good dispersion for each series. After these limits, a slight activity decrease is observed for the XA1T series which is however by far less than for the XA1H series. This could be ascribed to the high degree of sulfidation of Mo in the XA1T series alumina. In counterpart, large MoS2 crystallites have been evidenced on the sulfided XAIH solids with high Mo loading. Their formation originates from the suliidation of the MoO3 oxide evidenced on the oxidic precursor [9].
lo4/.
1.2 1.0 0.8 0.6
i . XAla
0.4
o.2
U
0.0
I 0
r 20
I 40
I 60
I 80
100
%MOS2
Figure 1. Conversion of T (in mol.gcata'l.s"1xl08) vs the MoS2 content (wt%).
111
60
2O
5O 15
40
10
30
5
20 I0
0
0 0
20
40
60
80
100
%MoS2
Figure 2. Formation of BP (in mol.gcata'X.S"1 xl0 8) VS the MoS2 content (wt%) during HDS of DBT.
0
20
40
60
80
100
%MoS 2
Figure 3. Formation of H4DBT (in mol.gc,t,'l.s "1 xl0 7) vs the MoS2 content (wt%) during HDS of DBT.
Figures 2 and 3 show the catalytic results for the conversion of DBT respectively into biphenyl (BP) and in tetrahydrodibenzothiophene (H4DBT) as a function of the MoS2 content. The evolution of H4DBT and BP products are similar. The H4DBT and BP productions on the XA1T catalysts are higher than on bulk MoS2 or on the XA1H catalysts. From these results it can be deduced that HDS of T, whatever the method of preparation, proceeds on well dispersed MoS2 single crystallites for which the number of active sites (at the edges of the slab) per gram of Mo is optimum i.e. at the limit of good dispersion. Upon increasing the Mo loading, the formation of bulk MoO3 in the XA1H oxidic precursor induces the formation of large MoS2 crystallites and provokes an important decrease of the activity per gram of Mo. In the case of the XA1T series, the activity decrease is limited because MoS2 remains as small nanocrystallites well distributed in alumina even for very high Mo loading. The method of preparation with the ATTM presented in this work allows, at high Mo loading, the formation of MoS2 nanocrystallites well distributed in the alumina, their size is smaller than those present in bulk MoS2. The evolution of DBT transformation is therefore in agreement with the Rim-Edge model proposed by Daage and Chianelli [7] on divided bulk MoS2. The morphology (stacking and length) of the MoS2 crystal]ites as well as the absence of interaction with alumina, which permit to avoid the presence of some steric effect for the approach of the DBT molecule, allow to
112
increase the activity in conversion of DBT. The HDS and HYD activities are therefore higher than the ones obtained with the XA1H or bulk disulfide. As no variation in the stacking is observed with the Mo loading the HDS/HYD ratio is not changed. As a conclusion of this work, it appears that the synthesis of MoS2A12Os catalysts by direct incorporation of ATTM in the alumina gel generates an active phase which allows to increase the DBT conversion. These catalysts are therefore good candidates for deep HDS. 4. ACKNOWLEDGEMENTS C. M. thanks the Exxon Company for the permission to use their catalytic set up to perform measurements on DBT conversion 5. R E F E R E N C E S
1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16
H. Topsoe, B. S. Clausen and F. E. Massoth, Hydrotreating Catalysis, Springer-Verlag, Berlin (1996) and references therein. J. Grimblot, Catal. Today, 41 (1998) 111. R. Candia, O. Sorensen, J. Villadsen, N. Y. Topsoe, B. S. Clausen and H. Topsoe, Bull. Soc.Chim. Belg., 93 (1984) 763. G.L. Schrader and C. P. Cheng, J. Catal., 80 (1983) 369. T. Hayden and J. A. Dumesic, J. Catal., 103 (1987) 366. S. Kasztelan, H. Toulhoat, J. Grimblot and J. P. Bonnelle, Appl. Catal., 13 (1984) 127. M. Daage and R. R. Chianelli, J. Catal., 149 (1994) 414. L. Le Bihan, C. Mauchauss~, L. Duhamel, J. Grimblot and E. Payen, J. Sol-Gel Sci. Tech., 2 (1994) 837. L. Le Bihan, Doctoral Thesis, Lille, France (1997). P.T. Vasudevan and Fan Zhang, Appl. Catal.A, 112 (1994)161. Fan Zhang and P. T. Vasudevan, J. Catal., 157 (1995) 536. A. Muller, E. Diemann, A. Branding, F. W. Baumann, M. Breysse and M. Vrinat, Appl. Catal., 62 (1990) L13. Y. Bensimon, Doctoral Thesis, Montpellier, France (1989). E. Payen, S. Kasztelan, S. Houssenbay, R. Szymanski and J. Grimblot, J. Phys. Chem., 93 (1989) 6501. A. Galtayries, S. Wisniewski and J. Grimblot, J. Electron Spectrosc., 87 (1997) 31. J. C. Muijsers, T. Weber, R. M. van Hardeveld, H. W. Zandbergen and J. W. Niemantsverdriet, J. Catal., 157 (1995) 698.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
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Influence of sulphidation and fluoridation on the HDN of o-toluidine over tungsten catalysts ex ammonium tetrathiotungstate Mingyong Sun and Roel Prins Laboratory for Technical Chemistry, Federal Institute of Technology (ETH), 8092 Zurich, Switzerland Abstract Alumina-supported tungsten catalysts, prepared by decomposition of supported ammonium tetrathiotungstate, had a much higher hydrodenitrogenation activity than classically prepared tungsten catalysts, because they were much better sulphided. Fluoride addition substantially increased the activity for the hydrogenation of the aromatic ring and the direct cleavage of the C(sp~)-N bond in o-toluidine, while it hardly changed the activity for C(spS)-N bond breaking in methylcyclohexylamine and olefin hydrogenation.
1. INTRODUCTION To process heavy and low quality feedstocks, which are rich in highly refractory nitrogen and sulphur-containing compounds and multi-ring aromatics, catalysts with a high activity for hydrogenation of aromatic rings in hydrodenitrogenation (HDN) and deep hydrodesulphurization (HDS) are required. Therefore improved methods of preparing catalysts with a high hydrogenation activity are sought. In this respect, Ni-W/AI~O3 seems to be a promising option [1]. However, it is known that tungsten supported on alumina is very difficult to sulphide because of the strong interaction between the oxidic W phase and alumina [2,3]. This relatively low sulphidability provides a large potential for improving the performance of W-based catalysts. There are two ways to improve the sulphidation of W catalysts. One is to impregnate the support with a sulphidic tungsten compound instead of with an oxidic compound. Vasudevan et al. prepared molybdenum sulphide catalysts by decomposition of supported ammonium tetrathiomolybdate ((NH4)~MoS4, ATTM) [4, 5] and tungsten sulphide catalysts by decomposition of supported ammonium tetrathiotungstate ((NH4)2WS4, ATT) [6]. They conducted the HDS of thiophene and the hydrogenation (HYD) of propylene to test the hydrogenation and C-S cleavage activities of catalysts prepared with ATTM and ATT. They also investigated the effect of pre-treatment on the performance of their catalysts. They found that catalysts prepared by decomposition of supported ATTM and ATT had much higher HYD and HDS activities than conventional catalysts prepared by sulphidation of oxidic precursors. A study of the HDN activity of catalysts prepared with ATT and ATTM has not yet been published.
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Another possibility is to add compounds that improve the sulphidability of the catalyst. Additives such as fluoride or phosphate are often incorporated into hydrotreating catalysts. Benitez et al. investigated the influence of fluoride on the structure and hydrotreating activity of W/A1203 catalysts. They found t h a t the incorporation of fluoride in the alumina support led to a moderate increase in the HDS activity for gas oil and to a greater increase in the HDN activity for pyridine [7]. Because several factors are involved in the effect of fluoride on W/A1203 catalysts, it is difficult to determine the exact role of fluoride in the catalysts. Fluoride may influence not only the interaction of the oxidic W phase with the alumina, and thus change the sulphidability and the WS2morphology , but it may also influence the surface acidity. Classically prepared sulphided tungsten-based catalysts are actually mixtures of oxides and sulphides. The improvement in sulphidability of the W phase, induced by additives (F or P), is considered to be one of the main reasons for the improved catalytic performance of additivecontaining tungsten-based catalysts [7,8]. Most studies on the effect of fluoride have concentrated on the overall HDS and HDN. Few studies considered the elementary reactions involved in HDS and HDN processes such as the hydrogenation of the aromatic ring, the cleavage of the C-S or C-N bond, and olefin saturation, which are important for understanding catalyst behaviour. Substituted anilines are important intermediates in the HDN of nitrogen-containing heterocycles, and the conversion of these intermediates is an essential step in the HDN process. In the present study, o-toluidine (Tol) and methylcyclohexylamine (MCHA) were chosen as model reactants. O-toluidine is the simplest substituted aniline. Its HDN process includes the hydrogenation of the phenyl ring, direct C(sp~)-N bond cleavage, and the hydrogenation of the methylcyclohexene (MCHE) intermediate to methylcyclohexane (MCH). Thus, by conducting the HDN of o-toluidine we will obtain important information as to how the different catalysts catalyze these reactions.
2. EXPERIMENTAL 2.1. Catalyst Preparation ATT was synthesized according to the procedure described by R a m a n a t h a n and Weller [9]. ATT/A1203 (ATT) and ATT/A1203F (ATTF) catalysts were prepared by means of the incipient wetness impregnation method. The original 7-A1203 pellets (Condea, surface area 228 m2/g, pore volume 0.49 ml/g) were milled and sieved to get a particle size of 0.25-0.45 mm. A1203F was prepared by impregnating the A1203 with an aqueous solution of ammonium fluoride followed by drying at 120~ for 4 h and calcination at 500~ for 4 h; the resulting fluoride content was 1 wt%. Dried A1203 and A1203F were impregnated with a solution of ATT in N,N-dimethylformamide under argon. After impregnation, ATT/A1203 and ATT/A120~F were dried in a vacuum desiccator. WO]A1203 (WO) and WO3/A1203F (WOF) catalysts were prepared by impregnating the A1203 and A1203F with an
115
aqueous solution of ammonium(meta)tungstate followed by drying at 120~ for 4 h and calcination at 500~ for 4 h. The tungsten loading of all catalysts was 16.5 wt% in WO3. 2.2. A c t i v i t y t e s t The catalytic reactions were carried out in a continuous-flow micro-reactor filled with 0.2 to 0.8 g catalyst diluted with 8 g SiC. All catalysts were pre-treated in situ with a mixture of 10% (mol) H2S/H 2 at 1.5 MPa by heating the catalyst at a rate of 5~ until it reached 400~ at which temperature it was kept for 4 h. After pre-treatment, the pressure was increased to 3.0 MPa, and the liquid reactant was fed to the reactor by means of a high-pressure pump, with n-octane as the solvent. The composition of gas-phase reactant is given in Table 1. H2S was produced by adding a certain amount of dimethyldisulphide to the feed, and n-heptane was used as an internal standard.
Table1 G a s - p h a s e composition of t h e feed Components
Partial Pressure, kPa
H2 Octane O-toluidine Heptane H2S CH 4
2882 85 9 12 6 6
3. R E S U L T S AND D I S C U S S I O N 3.1. O v e r a l l H D N c o n v e r s i o n of t o l u i d i n e According to the Langmuir-Hinshelwood mechanism, the rate of HDN can be written as dPTo/dX =kTo~Kro,PTo,/(I+Kro~PTo,+KNn3PNH3). At low conversion, this equation can be simplified by making the toluidine partial pressure in the denominator equal to the initial pressure and ignoring ammonia adsorption. This gives dP~o/dx = k'~o~P~o,,where k'To~is the pseudo first order reaction apparent rate constant, k'= kwolgwo/(l+ KwolP~ Thus we have: ln(1-Xro~)= -k'x, where X~o~is the overall conversion of toluidine. Figure 1 shows plots of-ln(1-Xro 1) versus space time for the four different catalysts. The R 2 values are larger t h a n 0.99 for all lines. The good linear behaviour of the curves indicates that the simplified equation is justified.
116 The apparent first-order rate 400 constants for the different /X ATTF catalysts, the slopes of the lines, X are given in the second column / ATT of Table 2. They show that the ~x~ 200 two ATT catalysts have much WO higher activities than the WO catalysts. The addition of ~' 0 I fluoride improves the activity of 100 200 the ATT as well as of the WO catalyst. The activity of fluoridespacetime [min.g/mol] containing catalysts compared Figure 1. Activity comparison for Tol to F-free catalysts increased by HDN (370~ 3.0 MPa) a factor of about 1.4 for the WO catalyst and by about 1.3 for the ATT catalyst. The ATT method increased the activity by a factor of about 5.6 and the combination of these two methods increased the activity by a factor of about 7.2. r
Table 2 The apparent rate constants in the HDN of Tol and hydrogenation of CHE in 103mol/min.g at 370~ and 3.0 MPa Catalysts
k' Tol,
k' T,
k ~..
k'CHE
WO WOF
0.62 0.89
0.06 0.11
0.56 0.78
14 15
ATT ATTF
3.5 4.5
0.23 0.28
3.2 4.2
26 26
The higher sulphidability is one of the reasons why the activity of a WOF catalyst is 40% higher than that of its non-fluoride counterpart [7]. Upon fluoride addition the W phase goes from strong interaction with the support, low reducibility and low WS 2 crystallite stacking for the fluoride-free catalyst to a more reducible and sulphidable form and with more stacked WS 2 crystallites for the fluoride-containing catalyst [7]. It is well known that the addition of fluoride to an alumina support increases its surface acidity [10,11]. F modifies catalytic properties in two ways: On the one hand, it diminishes the interaction of the Wphase and the alumina support, thus improves the sulphidability of the W phase and changes the morphology of WS 2. On the other hand, it can bring about higher surface acidity. ATT catalysts are already (almost) fully sulphided, and their sulphidability can hardly be improved by adding fluoride. However, the weakened interaction between the W phase and the alumina support still favours
117
the formation of larger WS 2 crystals and higher stacking during the decomposition of ATT. In the case of WO, all factors contribute to the promotional effect, while in the case of ATT, the contribution made by improving the sulphidability can be neglected. This explains why the promotional effect of fluoride is greater for the WO catalyst than for the ATT catalyst. 3.2. E l e m e n t a r y r e a c t i o n s in t h e H D N of t o l u i d i n e Figure 2 shows the HDN reaction network of o-toluidine. The removal of CH3 the nitrogen atom can proceed via two reaction pathways: by direct C-N bond HE kT cleavage of o-toluidine to toluene and by hydrogenation of o-toluidine to MCHA, which very quickly undergoes elimination CH3 l kn of NH 3 to MCHE. Direct C(sp3)-N cleavage of MCHA producing MCH is also a possible path, because the CH3 7 selectivity for MCH in the HDN of o~ 7 NH2 toluidine as well as of MCHA did not go CH3 to zero at low conversion as it would have done if all the MCH had come from MCHE (Figs 3 and 5). MCHA was not detected in the product mixtures of any of the catalysts under Figure 2. HDN reaction Network of Tol our experimental conditions, which means t h a t the rate of ammonia elimination is so fast that MCHA is converted before it leaves the catalyst beds. Therefore, the hydrogenation of the aromatic Tol ring is the rate-determining step. The increase in the conversion of o-toluidine must be due to the increase in the activity for the hydrogenation of the aromatic ring and/or C(sp2)-N direct cleavage. Two experiments, one starting with MCHE and the other with toluene, and both with 2-ethylaniline in the feed, proved that less t h a n 0.5% of MCHE was converted to toluene and only about 0.5% of toluene to MCHE and MCH even at very high space-time. Thus, the hydrogenation of toluene to MCHE and MCH and the dehydrogenation of MCHE to toluene can be neglected under our experimental conditions, 3.0 MPa and 330 - 370~ as long as Tol is present. Therefore, the hydrogenolysis of o-toluidine to toluene and the hydrogenation of o-toluidine to MCHA can be considered to be two independent parallel paths under the present conditions. By plotting -ln(1-X r) and -ln(1-X H) versus spacetime, the apparent rate constants k' T and k' H for the two paths can be calculated, where Xr is the conversion of o-toluidine to toluene and X His the conversion of o-toluidine to all products obtained by hydrogenation of the aromatic ring. The very good linearity (R2>0.99) of the curves indicates that the two reactions, the direct cleavage of C-N bond and the hydrogenation of the aromatic Tol ring can
d)
d
118
be treated as two parallel pseudo-first-order reactions. The results are summarized in Table 2. The data in Table 2 show that adding fluoride to the WO catalyst increases the activity for the C-N bond breaking by a factor of 1.8 while it increases the activity of aromatic ring hydrogenation by a factor of 1.4. However, for ATT catalysts, the fluoride effect is less pronounced, the factor being 1.2 for C-N bond breaking and 1.3 for aromatic ring hydrogenation. On the other hand, the ATT method favours aromatic ring hydrogenation to a greater extent than C-N bond cleavage. The promotional factor of the ATT method for aromatic ring hydrogenation is 5.7 and is only 3.8 for C-N bond breaking. The enhanced sulphidability is essential to both aromatic hydrogenation and hydrogenolysis, but it favours hydrogenation more than hydrogenolysis. The different effect of the ATT method and fluoride addition on aromatic ring hydrogenation and C(sp~)-N bond cleavage confirm that such reactions take place on different active sites. The distribution of the products of the aromatic ring hydrogenation 1oo path is also catalyst-dependent. WO For the same amount of -~ 8O WOF TT hydrogenation product, the ATT ATTF and F-containing catalysts produce less MCH and more MCHE than 4o the corresponding F-free and WO catalysts (Fig. 3). On the other hand, we have seen that ATT and F-containing catalysts have a 0 10 20 30 higher activity for the Conv. of To] to MCHE and MCH, % hydrogenation of Tol. In order to investigate this further, we carried Figure 3. Distribution of hydroout a simultaneous reaction of genation products cyclohexene (CHE) and Tol. The (370~ 3.0 MPa) results (Table 2, last column) show that the hydrogenation activity for CHE was hardly changed by the addition of fluoride. From the reaction network, we know that MCHE is formed by hydrogenation of Tol, followed by a fast elimination of ammonia, and removed by further hydrogenation. The ATT method and the addition of fluoride do not improve the catalytic activity for olefin hydrogenation as much as they improve aromatic ring saturation. As a consequence, there is an accumulation of MCHE in the products of the ATT and F-containing catalysts. . . . . . . . . . . .
3.3. H D N of M C H A
I
.......
I
.........
I
The HDN of MCHA was carried out to study the influence of catalyst preparation methods on the C(sp3)-N breaking in MCHA. ATT catalysts have a higher activity for conversion of MCHA (Fig. 4) and a lower selectivity for MCH (Fig. 5) than WO catalysts. The addition of fluoride did not change the rate of
119
disappearance of MCHA (Fig. 4) and decreased the selectivity for MCH in the ATT and WO catalysts (Fig. 5). In the HDN of MCHA, the 0.8 nitrogen atom can be removed by elimination of NH~ to MCHE (which can be hydrogenated to MCH) and by direct C(sp3)-N o~ 0.4 ] ~ ~ breaking, thus leading to MCH. At low conversion of MCHA, little MCHE can be 0 50 100 hydrogenated to MCH because space time [min.g/mol] of the strong inhibition effect of MCHA. The difference in Figure 4. Conv. of MCHA vs space. selectivity for MCH is mainly time (320~ 3.0 MPa) due to the difference in the ratio of the rates for the two paths. Fluoride addition ~, 40 favours the formation of MCHE via elimination and disfavours the formation of MCH via 2o direct C(sp3)-N breaking. -~"~'~'" ATTF Considering that the Fcontaining and F-free catalysts O have the same activity for 0 20 40 60 80 MCHA conversion, it seems conversion of MCHA, % that, upon fluoride addition, the sites that catalyze C(sp3)-N Figure 5. Selectivity of MCH breaking are modified to sites (320~ 3.0 MPa) that catalyze ammonia elimination, while the total number of sites does not change. ATT catalysts have a higher activity for conversion of MCHA and a lower selectivity for saturated product than WO catalysts. That is because the ATT method favours the formation of MCHE via elimination more than the formation of MCH via C(sp3)-N breaking. The ATT method may increase the number of active sites for the two mechanisms to a different extent, or it may change the intrinsic activity of the active sites for the different mechanisms in different ways.
/
0
0
-,-
u ..............................................
.e..4
!
4. C O N C L U S I O N S Sulphided tungsten catalysts prepared by decomposition of supported ATT have a much higher activity in the HDN of o-toluidine and methylcyclohexylamine than catalysts prepared from tungsten oxide. The preimpregnation of the alumina support with fluoride has a promotional effect on
120
the HDN of o-toluidine over both catalysts. The effect on the tungsten oxide catalyst is larger, because the sulphidability of ATT catalysts can hardly be increased by adding fluoride. The ATT method and the fluoridation have different effects on C(sp2)-N and C(sp3)-N bond cleavage, NH~ elimination, olefin hydrogenation, and aromatic ring hydrogenation. The ATT method favours aromatic ring hydrogenation more than C(sp2)-N bond cleavage. Fluoridation improves the C(sp2)-N bond cleavage more over oxide catalysts, and improves aromatic ring hydrogenation more over ATT catalysts, while it has little effect on the conversion of MCHA and the hydrogenation of CHE.
5. R E F E R E N C E S
1. S.P. Ahuja, M.L. Derrien and J.F. Le Page, Ind. Eng. Chem. Prod. Res. Dev., 9 (1970) 272. 2. B. Scheffer, P.J. Mangnus and J.A. Moulijn, J. Catal., 121 (1990) 18. 3. P.J. Mangnus, A. Bos and J.A. Moulijn, J. Catal., 146 (1994) 437. 4. P.T. Vasudevan and S.W. Weller, J. Catal., 99 (1986) 235. 5. F. Zhang and P.T. Vasudevan, J. Catal., 157 (1995) 536. 6. K. Wilkinson, M.D. Merchan, and P.T. Vasudevan, J. Catal., 171 (1997) 325. 7. A. Benitez, J. Ramirez, A. Vazquez, D. Acosta and A. Lopez Agudo, Appl. Catal., 133 (1995) 103. 8. J. Cruz Reyes, M. Avalos-Borja, R. Lopez Cordero and A. Lopez Agudo, Appl. Catal., 120 (1994) 147. 9. K. Ramanathan and S.W. Weller, J. Catal., 95 (1985) 249. 10. F.P.J.M. Kerkhof, J.C. Oudejans, J.A. Moulijn and E.R.A. Matulewicz, J. Colloid Interface Sci., 77 (1980) 120. 11. A. Corma, V. Fornes and E. Ortega, J. Catal., 92 (1985) 284.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
121
Modeling Molybdenum Carbide-Based Hydrodesulfurization (HDS) Catalysts Using Carbon-Modified Mo(110) Surfaces Charles L. Roe and Kirk H. Schulz Michigan Technological University, Department of Chemical Engineering, 1400 Townsend Drive, Houghton, MI 49931
Abstract Transition metal carbides, such as Mo2C, have been proposed as substitutes for group VIII metal catalysts, since they exhibit similar catalytic properties in some applications. Mo2C catalysts have shown potential for commercial use in hydrodesulfurization (HDS) processes and tend to resist sulfur poisoning better than platinum group metals. Although these molybdenum carbide catalysts look encouraging as replacements for MoS2-based catalysts, questions remain regarding the fundamental surface chemistry associated with the HDS of organosulfur molecules on carbided and sulfided molybdenum catalyst surfaces. To further investigate the suitability of Mo2C for HDS applications, the interaction of sulfur-containing molecules with molybdenum surfaces was examined by utilizing carbonmodified Mo(110) single crystals as model catalysts. Specifically, the reactivity of ethanethiol and 1,2-ethanedithiol were studied on the clean Mo(110), defective p(4x4)-C/Mo(110), and p(4x4)-C/Mo(110) surfaces using temperature programmed desorption (TPD), Auger electron spectroscopy (AES), and low energy electron diffraction (LEED). Ethanethiol and 1,2-ethanedithiol TPD experiments demonstrated that the presence of multiple sulfhydryl (SH) groups influences surface chemistry, given the differences observed in product distribution. Ethanethiol experiments performed on clean Mo(110) surfaces yielded ethane and ethylene as reaction products, while 1,2-ethanedithiol TPD experiments produced acetylene, ethylene, vinyl thiol, and ethanethiol. Ethanethiol TPD experiments showed that no significant differences in reactivity, selectivity, or reaction pathways exist between clean Mo(110) and the defective p(4x4) surfaces. 1,2-Ethanedithiol TPD experiments performed on the clean Mo(110) and p(4x4)-C/Mo(110) surfaces produced similar reaction products, although significant changes were observed in selectivity. On the clean surface, the major desorption products were acetylene, ethylene, vinyl thiol, and ethanethiol. However, the reaction of 1,2-ethanedithiol on the p(4x4)-C/Mo(110) surface produced only acetylene and ethylene. Thus, complete desulfurization of 1,2ethanedithiol occurs on the p(4x4) surface upon decomposition, yielding only hydrocarbon products. 1. INTRODUCTION Metal sulfide materials, such as MoS2, are commonly used as industrial hydrodesulfurization (HDS) catalysts to remove sulfur from the organosulfur compounds found in petroleum feedstocks. However, transition metal carbides have recently shown potential for use in commercial HDS processes. Indeed, a 13-M02C catalyst supported on alumina was shown to have a higher activity for thiophene desulfurization than similar sulfided Mo/A1203 catalysts with the same molybdenum loading [1 ]. Although hydrodesulfurization reactions have received significant attention, the catalytic role of surface carbon in desulfurization pathways is not well
122
understood [2,3]. There are indications that carbidic surfaces may be involved in the reaction pathways associated with hydrodesulfurization [4]. In addition, surface intermediates have been proposed that advocate bonding to adsorbed surface carbon during the decomposition of organosulfur compounds [5]. Carbon has also been suggested as a support material for MoS2based HDS catalysts because carbon can easily be burned away, such that the molybdenum can be recovered and reused [6]. In fact, MoSz-based catalysts supported on carbon are reported to be more catalytically active than those supported on alumina [7,8]. Although molybdenum carbide materials look encouraging as replacements for MoSz-based HDS catalysts, questions remain regarding the fundamental surface chemistry associated with the desulfurization of organosulfur molecules on carbon-modified molybdenum surfaces. Reactivity studies performed on single crystal materials can often provide insight into the reaction pathways and kinetics of industrially significant catalytic systems. Previous surface science studies examining HDS reactions have focused on using Mo(110) and Mo(100) single crystal surfaces as model catalysts, but have not been extended to include models of the newer molybdenum carbide-based materials. In particular, desulfurization reactions have been extensively studied on Mo(ll0) surfaces [9]. In fact, the well-characterized surface intermediates and reactivity of organosulfur molecules on Mo(110) make this surface ideal to investigate the effects of surface carbon on desulfurization reactions. Recently, there have been several publications on the reactivity of Mo(110) surfaces with ordered carbon overlayers. Frtihberger and Chen have examined the reaction of hydrocarbons such as CO, ethylene, cyclopentene, and cyclohexene on the clean Mo(110) and p(4x4)-C/Mo(110) surfaces [10-12]. They observed significant differences in surface chemistry for the adsorption of these molecules on clean and carbon-modified Mo(110) surfaces, which include different surface intermediates, decomposition pathways, and adsorption sites. Furthermore, Frtihberger and Chen suggest that surface carbon may alter the reactivity of Mo(110) surfaces by electronic modification of the substrate and by acting as a site-blocking agent [10]. Young and Slavin have extensively studied and characterized the structure of the p(4x4)-C/Mo(110) surface, proposing that the adsorbed carbon atoms sit in the 4-fold hollows at the centers of each unit cell of the Mo(110) substrate [13]. Although the p(4x4)-C/Mo(110) surface has been prepared and well characterized, HDS reactions on this surface have not been investigated. To further ascertain the suitability of molybdenum carbides for HDS applications, the interaction of ethanethiol and 1,2-ethanedithiol with molybdenum surfaces was examined by utilizing carbon-modified Mo(110) single crystals as model catalysts. Specifically, the reactivity of these thiols were studied on the clean Mo(110), defective p(4x4)-C/Mo(110), and p(4x4)C/Mo(ll0) surfaces using temperature programmed desorption (TPD), Auger electron spectroscopy (AES), and low energy electron diffraction (LEED) [14,15]. Although not prominent impurities in crude oil, ethanethiol and 1,2-ethanedithiol were chosen for initial studies in order to determine the effect of carbon on the surface chemistry of relatively simple alkanethiols. The dithiol was specifically incorporated in these studies to probe the differences in reactivity between thiols with singular (ethanethiol) and multiple (1,2-ethanedithiol) sulfhydryl groups 2. EXPERIMENTAL SECTION All experiments were performed in an ion-pumped, stainless steel, single chamber ultrahigh vacuum (UHV) system equipped with a Thermionics TTS-240N ion-pump. The chamber contains an Omicron Vakuumphysik Spectaleed L-A-2000 unit, capable of both low
123
energy electron diffraction (LEED) and Auger electron spectroscopy (AES). For AES measurements, the unit is operated as a retarded field analyzer with an incident electron beam energy of 1.5 kV and a resolution of AE/E = 0.45%. LEED experiments are performed using the unit as a reverse-view, four grid, LEED optics system. Temperature programmed desorption (TPD) experiments were performed using a Leybold Inficon Transpector H200M quadrapole mass spectrometer. During TPD experiments, a linear 2 K/sec. rise in sample temperature was achieved using a Hewlett Packard 6672A DC power supply and a proportional control algorithm. No corrections were made to the TPD spectra for ion gauge or mass spectrometer sensitivity. A 99.9% zone refined Mo(110) single crystal sample was obtained from Alfa Aesar. Sample positioning is accomplished using a Thermionics EC-1275-XYZ manipulator attached to a RNN-150/FA differentially-pumped rotary platform. The Mo(110) single crystal sample is mounted on a sheet of 0.5-mm thick tantalum foil and mechanically held in contact with the tantalum. This sample mounting arrangement has been previously described in more detail [ 16]. Ethanethiol (Aldrich, 97% pure), 1,2-ethanedithiol (Aldrich, 90+% pure), and ethylene (Matheson Gas, polymer grade, 99.9%) were used in the experimental studies as received. Sample dosing was accomplished by backfilling the vacuum chamber to a desired pressure for a given length of time. All reactant exposures were measured in Langmuirs (1 L = 1 0 -6 Yorr.seconds). 3. RESULTS AND DISCUSSION 3.1. Clean and Carbided M o ( l l 0 ) Surface Structures Ethanethiol and 1,2-ethanedithiol TPD experiments were performed on the clean Mo(110) surface, the defective p(4x4)-C/Mo(110) surface, and the p(4x4)-C/Mo(110) surface [14,15]. As shown in Figure l(a), the Mo(110) surface has a distinctive (lxl) hexagonal LEED pattern, which is characteristic of the clean single crystal surface.
Figure 1. LEED patterns obtained for the (a) clean Mo(110), (b) defective (4x4)-C/Mo(110), and (c) p(4x4)-C/Mo(110) surfaces. A p(4x4)-C/Mo(110) surface was prepared by dosing 300 L of ethylene onto a clean Mo(110) surface held at 1200 K. The Mo(110) sample was further annealed at 1200 K for 25-30 rain, which produced an ordered carbon overlayer. The p(4x4)-C/Mo(110) surface has a characteristic AES peak-to-peak C/Mo ratio of 0.30 and a corresponding LEED pattern depicted in Figure l(c). As noted previously, Young and Slavin have characterized the structure of this surface [13]. Frfihberger and Chen have extensively studied the reactivity of this surface towards
124 hydrocarbons and have noted that the p(4x4) surface has properties similar to molybdenum carbides [3,10-12]. St. Clair et. al. have studied and characterized an ot-Mo2C (0001) surface using LEED and XPS [18]. The LEED pattern of the o>MozC (0001) surface has (4x4) hexagonal periodicity when this single crystal is carbon terminated, indicating structural similarity between the p(4x4)C/Mo(110) and a-MozC (0001) surfaces. The defective p(4x4)-C/Mo(110) surface was prepared by dosing 180 L of ethylene onto a clean Mo(110) surface held at 1200 K, followed by annealing the sample at 1200 K for 25-30 min. This procedure yielded a surface with an ordered carbon overlayer, which is characterized by a AES peak-to-peak C/Mo ratio of 0.20 and a distinct LEED pattern shown in Figure l(b). The LEED pattern for the defective surface resembles that of the p(4x4) surface, but does not have perfect p(4x4) periodicity due to missing spots in the diffraction pattern. Although the exact structure of this surface is not currently known, the defective p(4x4)-C/Mo(110) surface can be reproducibly formed and has not been previously reported to exist. Both the defective and p(4x4) surfaces exhibit AES carbon peak shapes that correspond to carbidic surface carbon and are not indicative of graphitic carbon. Overall, the amount of carbon coverage is the key difference between the defective (AES C/Mo = 0.2) and p(4x4) surfaces (AES C/Mo - 0.3). 3.2. Thiol TPD Experiments on the Mo(ll0) Surface On the Mo(110) surface, ethanethiol thermally decomposes to ethane and ethylene as reaction products as shown in Figure 2. Although hydrogen was not detected during TPD experiments, the background pressure of hydrogen in our UHV chamber was sufficiently high enough that desorbing H2 may have gone undetected. For the reaction of 1,2-ethanedithiol on the clean surface, both hydrocarbon and thiol products were observed, which included acetylene, ethylene, vinyl thiol (CH2=CHSH), and ethanethiol. Figure 3 depicts the TPD spectra resulting from 1,2-ethanedithiol experiments on the Mo(110) surface. Thus, the difference in surface chemistry between ethanethiol and 1,2-ethanedithiol on clean Mo(110) is attributed solely to the difference in molecular structure between these thiols.
x15
c 19
c 19
m
-~ c~
~
_.__.__
(o
.-~.
x2__2. . . .
v=_~nyJTh~o_'.~
~
O~,~
OOv '
I
100 200
'
I
"
300
I
400
'
I
500
'
I
600
'
I
700
'
I
800
Ethanethiol
I
'
900
100 200
'
Temperature (K) Figure 2. TPD spectra for the reaction of 1L of ethanethiol on clean Mo(110).
I
'
I
'
300
I
400
'
I
500
'
I
600
'
I
700
'
I
800
'
900
Temperature (K) Figure 3. TPD spectra for the reaction of 1L of 1,2-ethanedithiol on clean Mo(110).
Roberts and Friend performed TPD experiments for ethanethiol on Mo(110) and report the same reaction products at similar temperatures [17]. Ethanethiol is proposed to decompose to both ethane and ethylene via a thiolate surface intermediate, which is consistent with the
125 observations and results obtained during TPD studies. However, 1,2-ethanedithiol contains two sulfhydryl (SH) function groups and produces some novel bidentate/cyclic surface intermediates [15]. These intermediates are bound to the clean surface through either one or both sulfur atoms of the dithiol. AES measurements taken after TPD studies of ethanethiol and 1,2-ethanedithiol on Mo(110) indicated the presence of surface sulfur, which further supports the argument that these thiols bond to the surface through sulfur atoms. 3.3. Thiol TPD Experiments on Carbided Mo(ll0) Surfaces Ethanethiol and 1,2-ethanedithiol TPD experiments were carried out on a defective p(4x4)-C/Mo(110) surface in a similar manner to the experiments performed on the clean Mo(110) surface. Figure 4 shows TPD spectra for the reaction of 1L of ethanethiol on the defective p(4x4)-C/Mo(110) surface. Despite the presence of an ordered carbon overlayer, the spectra for ethylene and ethane are virtually identical to the corresponding spectra taken during a 1L ethanethiol TPD experiment on the clean Mo(110) surface. As with experiments on clean Mo(110), residual surface sulfur was left behind during TPD studies of ethanethiol on the defective p(4x4) surface.
>,~. m
c
c 03
c~ m m
k,..
c
t,~
-~,..t ~ m
Ethane.
.
x5
'
I
'
I
9
I
300 400
9
I
500
'
I
600
'
I
700
9
I
800
L,.,,t~.n
__. ,,
Acetylene Ethylene
.-~:
Hydrogen
100 200
j~J~'~~,,..,a~
O ? v
'
100 200
900
Figure 4. TPD spectra for the reaction of 1L of ethanethiol on the defective p(4x4)-C/Mo(110) surface.
300 400
500 600
700 800
900
Temperature (K)
Temperature (K)
Figure 5. TPD spectra for the reaction of 1L of 1,2-ethanedithiol on the p(4x4)-C/Mo(110) surface.
As shown in Figure 5, 1,2-ethanedithiol TPD experiments were performed on the p(4x4)C/Mo(110) surface. In contrast to the clean Mo(110) surface, only acetylene and ethylene were observed as reaction products. Small amounts of 1,2-ethanedithiol were also found to desorb at low temperatures. Again, sulfur was left behind on the p(4x4)-C/Mo(110) surface following TPD experiments. No ethanethiol TPD experiments were performed on the p(4x4) surface. 4. RESULTS AND DISCUSSION 4.1. Mo(ll0) Surface Comparatively, there are notable differences between the TPD spectra for ethanethiol and 1,2ethanedithiol on the Mo(ll0) surface. These differences in product distribution are evidently due to the differences in molecular structure between
/CH3 H2C\
> Ethane > Ethylene + H2
Thiolate Intermediate Figure 6. Reaction Pathways for Ethanethiol on Mo(110). 15
126
these thiols. As depicted by Figure 6, ethanethiol is thought to form a thiolate surface intermediate that further decomposes to hydrogenation (ethane) and [3-dehydrogenation (ethylene) products. These proposed reaction pathways are consistent with TPD experimental results and with similar studies found in the literature [9,17]. During 1,2-ethanedithiol TPD experiments, acetylene, ethylene, ethanethiol, and vinyl thiol were detected as reaction products. As shown in Figure 7, we propose that 1,2ethanedithiol decomposes via two distinct types of CH3CH2SH CH2=CHSH surface intermediates. Thus, EthanetiiolProductHS\/CH2 HS\~cH VinylTiiol.Product the dithiol is bound to the Mo(110) surface though H2C\ .~ H C \ either one (monodentate, Intermediates I and III) or both (bidentate, IntermediateI Intermediate III Intermediates II and IV) HSCH2CH2SH1,2-Ethanedithiol~ ~ [sulfurlLeft Surface Behind ] sulfur atoms. As with the /CHcCH2\ /CH=CH\ ethanethiol TPD studies, 1,2-ethanedithiol yields both IntermediateII Intermediate IV hydrogenation (ethanethiol) + + and 13-dehydrogenation CH2= C H 2 CH- CH (vinyl thiol) products via the EthyleneProduct AcetyleneProduct monodentate surface Figure 7. Reaction Pathways for 1,2-Ethanedithiol on Mo(110). 14 intermediates, while the bidentate species produce acetylene and ethylene. These proposed surface intermediates are similar to those formed on Mo(110) by ethylene glycol (HOCH2CHzOH), which is the oxygen analogue to 1,2-ethanedithiol [19]. ,,,._
4.2. Carbided Mo(110) Surfaces Based on TPD results from the reaction of ethanethiol and 1,2-ethanedithiol on the defective p(4x4)-C/Mo(ll0) surface, the adsorbed surface carbon of the defective surface appears to have no effect on reaction pathways, surface intermediates, or reactivity of Mo(110). In terms of peak position and line-shape, the ethanethiol TPD spectra obtained on the defective surface are almost identical to spectra taken on clean Mo(110). Preliminary data for the reaction of 1,2-ethanedithiol on the defective p(4x4) surface also indicates that the defective surface behaves like the clean surface. Frfihberger and Chen suggest that the nature of the adsorbed carbon on Mo(110) has a strong effect on catalytic properties [20]. From their studies of hydrocarbons on carbon-modified Mo(110) surfaces, they observe that mere graphitic surface carbon does not appear to affect hydrocarbon surface chemistry other than to reduce reactivity. These experimental results further demonstrate that the relative amount of surface carbon is also a factor that influences the catalytic properties of the Mo(110) surface, not just the chemical state of the surface carbon. Clear differences exist between the reaction of 1,2-ethanedithiol on the clean Mo(110) surface and p(4x4)-C/Mo(110) surface, due to a shift in the reaction pathways. TPD experiments demonstrate that complete desulfurization of the dithiol occurs on the p(4x4) surface, since ethanethiol and vinyl thiol are absent from the product distribution. It appears as though the
127
cyclic surface species (Intermediates II and IV) are selectively favored as intermediates, which give rise to ethylene and acetylene as products. Given that acetylene and ethylene come off of the p(4x4) surface at lower temperatures relative to the clean Mo(110) surface, the reaction of the dithiol on the p(4x4) surface requires less activation energy to proceed. Thus, relative to the clean Mo(110) and defective p(4x4)-C/Mo(110) surfaces, the p(4x4) surface appears to contain enough surface carbon in the proper chemical state to induce subtle changes in the selectivity and activation energy of the dithiol reaction pathways. 5. CONCLUSIONS To probe the effects of surface carbon on Mo(110) desulfurization chemistry, the reaction of ethanethiol and 1,2-ethanedithiol were investigated on the clean Mo(110), defective p(4x4)C/Mo(110), and p(4x4)-C/Mo(110) surfaces. Due to the structural similarity of these probe molecules, both ethanethiol and 1,2-ethanedithiol form thiolate surface intermediates, which further decompose to C2 reaction products during TPD experiments. However, due to the presence of two sulfhydryl functional groups, 1,2-ethanedithiol also forms multiply bonded species to both clean and carbided Mo(110) surfaces. Thus, the differences in product distribution that are observed between ethanethiol and 1,2-ethanedithiol on each surface are the result of differences in molecular structure. Despite the presence of an ordered carbon overlayer, the defective p(4x4)-C/Mo(110) surface was found to catalytically behave like clean Mo(110) during TPD experiments. The carbon overlayer associated with the defective surface does not appear to affect thiol surface chemistry. Relative to the clean Mo(110) surface, 1,2-ethanedithiol TPD studies performed on the p(4x4)-C/Mo(110) surface showed a dramatic change in selectivity and resulted in complete desulfurization. The decomposition of 1,2-ethanedithiol on the clean surface produced both thiols and hydrocarbons, while the reaction of the dithiol on the p(4x4) surface yielded only hydrocarbons as products. The carbided Mo(110) surfaces differed only in surface structure and the extent of carbon coverage. Surface chemistry differences between the clean and carbided Mo(110) surfaces appear to be primarily surface structure related and are not exclusively related to the presence of carbidic surface carbon. 6. ACKNOWLEDGEMENTS The authors gratefully acknowledge the National Science Foundation (CTS-9523936) and the Michigan Research Excellence Fund for support of this work. Additionally, C. L. Roe would like to acknowledge support from a Michigan Technological University Challenge Fellowship. 7. REFERENCES
1. P.A. Aegerter, W.W.C. Quigley, G.J. Simpson, D.D. Ziegler, J.W. Logan, K.R. McCrea, S. Glazier, and M.E. Bussell, J. Catal., 164, 109 (1996). 2. R.R. Chianelli, M. Daage, M.J. Ledoux, Adv. in Catalysis, 40, 177 (1994). 3. J.G. Chen, Chem. Rev., 96, 1477 (1996). 4. R.L. Seiver and R.R. Chianelli, U.S. Patent #4431747 (1984).
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5. R. Ramachandran and F.E. Massoth, Can. J. Chem. Engr., 60, 17 (1982). 6. J.L. Schmitt and G.A. Castellion, U.S. Patent #3997473 (1976). 7. J.P.R. Vissers, B. Scheffer, V.H.J. DeBeer, J.A. Moulijn and R. Prins, J. Catal., 105,277 (1987). 8. J.C. Duchet, E.M. van Oers, V.H.J. DeBeer, and R. Prins, J. Catal., 80, 386 (1983). 9. B.C. Wiegand and C.M. Friend, Chem. Rev., 92, 491 (1992). 10. B. Frfihberger and J.G. Chen, Surf. Sci., 342, 38 (1995). 11. J.G. Chen and B. F~hberger, Surf. Sci., 367, L 102 (1996). 12. B. Frfihberger and J.G. Chen, J. Am. Chem. Soc., 118, 11599 (1996). 13. M.B. Young and A.J. Slavin, Surf. Sci., 245, 56 (1991). 14. C.L. Roe and K.H. Schulz, "Reaction of 1,2-Ethanedithiol on Clean, Sulfur-Modified, and Carbon-Modified Mo(110) Surfaces", J. Vac. Sci. Technol. A, 16, 1066 (1998). 15. C.L. Roe and K.H. Schulz, "Reaction of Ethanethiol on Clean and Carbon-Modified Mo(110) Surfaces as a Function of Sulfur Coverage", J. Phys. Chem., (IN PREPARATION). 16. S.L. Peterson, K.H. Schulz, C.A. Schulz, and J.M. Vohs, Rev. Sci. Instrum., 66, 3048 (1995). 17. J.T. Roberts and C.M. Friend, J. Phys. Chem., 92, 5205 (1988). 18. T.P. St. Clair, S.T. Oyama, and D.F. Cox, "Surface Characterization of c~-Mo2C (0001)", Sure Sci., (IN PREPARATION). 19. K.T. Queeney, C.R. Arumainayagam, M.K. Weldon, C.M. Friend, and M.Q. Blumberg, J. Am. Chem. Soc., 118, 3896 (1996). 20. B. Frthhberger, J.G. Chen, J. Eng, Jr., and B.E. Bent, J. Vac. Sci. Technol. A, 14, 1475 (1996).
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
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TiO2-coated on A1203 supportprepared by CVD method for HDS catalysts Kohichi Segawa* and Shinobu Satoh Department of Chemistry, Faculty of Science and Technology, Sophia University 7-1 Kioi-cho, Chiyoda-ku, Tokyo 102-8554, Japan. Phone: +81-3-3238-3452, Telefax: +81-3-3238-4350 Email:
[email protected]
Abstract
Composite TiO2-A1203 supports have been prepared by chemical vapor deposition (CVD) over 7-A1203 substrate, using TiC14 as the precursor. The specific surface area of each TiO2-A1203 support is comparable to that of 7-A1203. High dispersion of TiO 2 on the A1203 surface has been obtained, and no cluster formation has been detected. Mo/TiO2-A1203 catalysts have been synthesized by impregnation method. The catalytic behavior of Mo supported on A1203, TiO 2 and TiO2-A1203 has been investigated for the deep hydrodesulfurization (HDS) of dibenzothiophene (DBT) and methyl substituted DBT derivatives. The conversion over the TiO2-A1203 supported catalysts, in particular for the HDS of 4-methyl-DBT (4-MDBT) and 4,6-dimethyl-DBT (4,6-DMDBT) is much higher than the conversion obtained over A1203 supported materials. The ratio of the corresponding cyclohexylbenzene (CHB)/biphenyl (BP) derivatives is increased over the composite support, indicating that the prehydrogenation of an aromatic ring is important for the HDS of DBT derivatives over TiO2-A1203 supported catalysts.
1. INTRODUCTION Recently, the interest in the deep HDS of gas oil has been renewed due to the stricter regulations concerning the sulfur content in Diesel oil. In order to protect the global environment, the Japanese government has lowered the limit of the sulfur content to 0.05 wt% in 1997, and will have lowered to 0.005 wt% in the future. To meet this assignment, the development of a suitable catalyst for the deep HDS is a very important subject. In the petroleum refining industry, 7-alumina supported molybdenum oxide catalysts promoted with cobalt or nickel have been widely used for the HDS of sulfur compounds. Recently, TiO2 supported molybdena catalysts have attracted increasing attention, because of their higher reducibility to a lower valence state of molybdenum and their higher catalytic activity for HDS [1-3], as well as for hydrocracking, compared to alumina supported materials [4]. Various studies focused on the characterization of molybdenum species supported on TiO2, using several experimental techniques such as Raman spectroscopy, XPS and temperature programmed reduction (TPR) [5-10]. The formation of the active sites when MoO3 changes to MoS2 by sulfurization has been elucidated by Arnoldy et al. [ 11 ]. Nishijima et al. [ 12] and Shimada et al. [ 13] have investigated molybdenum catalysts supported on A1203, TiO2, SiO 2 and MgO by means of extended X-ray absorption fine structure (EXAFS), XPS and Raman spectroscopy. They reported the presence of different molybdenum species on the surface of the supports investigated. However, pure TiO2 supports have very small specific surface areas compared to alumina, and it is difficult to make pellets. Furthermore, the active anatase structure possesses only low thermal stability. This makes TiO2 support alone unsuitable for industrial applications. However, coating of the surface of commercially available 7-alumina
130
with TiO2 may lead to TiOE-A1203 composite supports which could help to overcome the disadvantages of pure titania supports[ 14-17].
S CH3
S CH 3
CH 3
- H2S CH 3
CH 3
CH 3
Hydrodesulfurization route
Hydrogenolysis route CH3
CH3
Scheme 1. Reaction pathways in HDS of 4.6-DMDBT. 4-MDBT and 4,6-DMDBT are key sulfur compounds in the gas oil fraction and make Arabian light middle distillates difficult to desulfurize [18-21]. Nevertheless, in order to achieve deep HDS, it is essential to desulfurize these molecules. Two reaction pathways for the HDS of alkyl substituted DBT are proposed; they are illustrated in Scheme 1. The direct sulfur abstraction (hydrogenolyis route) leads to BP derivatives, whereas prehydro-genation is forming hexahydrodibenzothiophene as an intermediate (hydrodesulfurization route), which is desulfurized to CHB derivatives. Furthermore, it is known that hydro-genation of neighboring phenyl groups reduces the steric hindrance caused by the methyl groups. To elucidate the influence of the positions of methyl-groups on the conversion rate, we studied the HDS of a mixture of DBT, 2-methyl-DBT (2-MDBT) and 2,8-dimethyl-DBY (2,8-DMDBT), as well as of a mixture of DBT, 4-MDBT and 4,6-DMDBT over Mo- catalysts on the different supports. We studied how to prepare composite types of TiOE-AI203 support by chemical vapor deposition (CVD) method of TIC14 on the substrate. Then we have studied deep HDS of dibezothiophene (DBT) derivatives as a model reaction. Since DBT derivatives are key sulfur compounds in the gas oil fraction.
2.EXPERIMENTAL
2.1 Preparation Titania (TiOE) was supplied by Degussa (P-25, 51mEg-l) and the alumina (T-A1203) support was provided by NIKKI (N611-N, 186 mEg-l). The composite types of TiOE-A1EO3 supports were prepared by CVD method of TiC1 4 on 7-alumina [22], according to the following procedure: 2 g of A1203 (12-24 mesh) substrate was placed in a quartz tubular reactor and pretreated for 2 h at 773 K in oxygen flow. The substrate was then exposed to TIC14 (WAKO) vapor (0.43 kPa) at 473 K, when TIC14 was mixed with N2 as carrier gas. The decomposition time was varied between 0.5 h and 20 h, in order to obtain different loadings of TiOz. Afterwards, the sample was hydrolyzed by water vapor (2.30 kPa) with N2 as carrier gas at 473 K for 2 h. Calcination was carried out under 02 flow for 2 h at 773 K. The supported Mo catalysts used in this study were prepared by an
131
impregnation method. Impregnation was performed at 323 K for 100 h, using an aqueous solution of ammonium heptamolybdate (0.004M) over support. After impregnation, all materials were calcined at 773K for 10 h in air. 2.2 HDS Reaction The HDS reaction was carried out in a fixed bed high pressure flow reactor, consisting of a 0.5 inch stainless steel tube packed with 250 mg catalyst mixed with quartz sand. Before the reaction, the catalysts were dried at 773 K for 5 h under oxygen stream and presulfided with a mixture of HES (5 %) and H2 for 2 h at 573 K under atmospheric pressure. After purging of excess HES by nitrogen stream at 573 K for 0.5 h, H2 and the reaction mixture were supplied to the reactor. The used methyl-substituted dibenzothiophene derivatives were synthesized as described in the literature [23]. All reaction mixtures used in this study were diluted in n-dodecane. The initial content of sulfur was 0.05 wt% and the reaction temperature was 573K (H2 pressure: 3 MPa, HE flow rate: 0.2 dmamin-1, LHSV: 12.08-16.07 h-l). The liquid products collected from a gas-liquid separator were analyzed by GC (HITACHI G-3000) and GC-MS (GC: HEWLETT PACKARD 5890 SERIES II, MS: NIHONDENSHI JMS-SX 102A). The activity of the catalysts under investigation was estimated by the conversion of the DBT derivatives and by the ratio of the products after reaching steady state. 2.3 Characterization of the catalysts Nitrogen adsorption measurements were performed at 77 K on a BEL Japan BELSORP 28SA, to determine the specific surface area. As pretreatment, 200 mg of support were placed in a quartz tube and evacuated for 6 h at 573 K. The pore volume and the pore size distribution was obtained by the Dollimore-Heal method [24]. XPS measurements were carried out at room temperature to investigate the state of molybdenum on the surface of the different supports. Data were collected before and after sulfiding for 2h at 573K, with 5%H2S diluted in H2 as carrier gas. The data were taken on a Surface Sciense Laboratory SSX-100 spectrometer, using monochromized A1 Ko~-radiation (1486.6 eV) and the C (Is) binding energy (285.5 eV) as reference. Before measuring, all samples have been evacuated and pretreated at 573 K.
3. RESULTS AND DISCUSSION 3.1 Characterization of catalysts Figure 1 shows the pore size distribution obtained for 7-A1203 (N61 l-N: 186 m 2g-1), TiO2 (P-25:55 m2g-1) and TiO2-A1203 composite supports (prepared by CVD method over 7-A1203:10.2 wt% TiO 2). It is obvious that pure TiO2 possesses no pore system, they are composed with fine samll particles and its crystal form shows anatase form by XRD powder pattern. With regard to 7-A1203 and TiO2-A1203, the average pore diameter is only slightly shifted to lower values, whereas the distribution of the pore volume is not significantly affected by the coverage of alumina with TiO2. Thus, TiO2 can be assumed to be situated on the surface of the A1203 support, and the incorporation of significant amounts of titanium into the 7-alumina matrix can, within the sensitivity of the investigation methods, be excluded. These results reflect a homogeneous dispersion of TiO2 on the surface of the A1203 support. The resulting composite TiO2-A1203 support remains mesoporous, with a more or less monomodal pore size distribution and an average pore diameter of about 3.8 nm. Significant discrepancies in the pore size distribution, caused by the different loadings of the two investigated TiO2-A1203 support, have not been detected. Judging by the results described above, we can assume that TiO2 covers the surface of the 7-alumina support mainly without forming precipitations. The IR spectra described above
132
strongly support this assumption, even though the vibrational bands of A1-OH groups are not completely eliminated at loadings of 11 wt% TiO2. This might be due to the fact that the alumina surface is not completely covered. With regard to a closed packing of TiO2 and under the assumptions of a titania monolayer, the specific area of the 7-alumina support would be completely covered at loadings of about 15 wt% titania. Furthermore, it is known that in the bulk phase of poorly crystalline alumina there are A1-OH groups present, even after calcination. These hydroxyl groups can not be substituted by A1-O-Ti groups, and the corresponding vibrational bands in the IR spectra remain unchanged. Nevertheless, higher loadings of titania may lead to a more complete elimination of hydroxyl bands on the alumina surface. The comparison of the surface composition and the total amount of TiO2 unambiguously reveals that TiO2 is highly dispersed on the surface of the 7-alumina support. According to these results, CVD allows the coverage of the surface of alumina supports by TiO2, using TiC14 as precursor without forming large aggregates, under variation of the TiO2 loading over a wide range. 200
I
I
I
I
150 AI203 >
<1 100
........ 0
.......
.... 0 - - -
TiO 2
TiO2-AI203
50
0
2
4
6
8
10
Pore Radius/nm
Figure 1. Pore size distribution of different supports. Molybdenum XPS investigations have been performed in order to elucidate the state of molybdenum sites on the TiOz-A1203 support. The results obtained by the Mo 3d XPS spectra for calcined and sulfided Mo/A1203, Mo/TiO2, Mo/TiO 2-A1203, crystalline MoO3 and crystalline MoS2 powder are illustrated in Figure 2. The peaks obtained for Mo/A1203 (B) and Mo/TiO2-A1203 (D) are much broader than those obtained for MoO 3 (A) and Mo/TiO2 (C), whereas the binding energy distribution detected for Mo/A1203 (B) is broader than that of Mo/TiOz-AI203 (D). The broader peaks of Mo 3d were stronly dependent on the support interactions. Nevertheless, the state of the molybdenum species on the TiOz-A1203 composite support can be regarded as a transition state between the Mo/A1203 and Mo/TiO2. The broadening of the 3d XPS peaks might be caused by strong interactions between Mo and the support, as well as by structural distortions. TiO2 loadings of 10.2 wt% provide Mo species which are more similar to the Mo species present on pure TiO2 supports than to molybdenum supported on A1203. The fact that TiO 2 loadings above 10.2 wt% lead to larger line broadening of the Mo XPS peaks might be caused by saturation effects while loading the A1203 surface with TiO 2 and the formation of TiO2 precipitations. However, within the
133
sensitivity of the experimental techniques used, no TiO2 clusters have been detected. After sulfurization, the maxima due to Mo(VI)-species have strongly decreased. The binding energies for crystalline MoS2 (H) are 233.0 eV (3d3/2) and 229.9 eV (3d 5/2). The spectra of sulfided MoO3 on the different supports can also be assigned to MoS2 species, whereas Mo (VI) species are still present. As already obtained for the calcined materials, sulfided Mo/A1203 (E) and Mo/TiOE-A1203 (G) reveal broader binding energy distributions than MoS2 and Mo/TiO 2 (F). This might be caused by the same reasons as mentioned above. However, the binding energies (3d5/2) of the sulfided catalysts increase as follows: Mo/A1203 < Mo/TiO2-A1203 = Mo/TiO 2. Nevertheless, in agreement with the calcined materials, the spectra of MoS2 and of sulfided Mo/TiO2 are very similar with regard to the half width of the peaks, indicating the structure to be very similar. As already stated above, and in particular in the case of the A1203 support, there are still Mo(VI) species present and the sensitivity of the XPS technique used, is not high enough to make any conclusion on the exact structural state of Mo on the surface of the different supports.
I '
Mo 6+ Mo 4+ r'--I
! !
I
S(2s) I !
~~'J
(H)
,.,
; (F) (E)
(D)
I i I
I I
tl
I
,,
(c)
!
(B)
I I I
i
224 Binding Energy/eV
(A) I
240
Figure 2. Mo 3d XPS spectra of the MoO3 (A), Mo/A1203 (MOO3; 6.0 wt%) oxidic state (B), Mo/TiO2 (MOO3; 6.0 wt%) oxidic state (C), Mo/TiO2-A1203 (MOO3; 6.0 wt%, TiO 2; 10.2 wt%) oxidic state (D), Mo/A1203 (MOO3; 6.0 wt%) sulfided state (E), Mo/TiO 2 (MOO3; 6.0 wt%) sulfided state (F), Mo/TiO2-A1203 (MoO 3; 6.0 wt%, TiO2; 10.2 wt%) sulfided state (G), and crystalline MoS2 (H).
134
3.2 H D S reactions
As already pointed out, the HDS of DBT derivatives leads to the corresponding substituted BP and CHB compounds (see Scheme 1). The product selectivity mainly depends on the reaction pathway. Prehydrogenation of neighboring phenyl groups leads to an intermediate, and then CHB derivatives are produced (hydrodesulfurization route). In contrast to that, the desulfurization without prehydrogenation (hydrogenolyis route) leads to the corresponding BP products. Table 1 Effect of substituent groups on dibenzothiophenes (DBT) for HDS over Mo/A1203 catalyst
sC%-2 9
1
6
4
7
3
Substituent position
DBT
4-methyl-DBT
4,6-dimethyl-DBT
Relative activity
1.00
0.43
0.17
Substituent position
DBT
2-methyl-DBT
2,8-dimethyl-DBT
Relative activity
1.00
1.50
1.97
_ , .
Reaction conditions: temperature; 573 K, pressure 3 Mpa, LHSV; 12.08-16.07 h -l, H 2 flow rate; 200 cm3 min-1,MOO3; 6 wt% To elucidate the influence of the positions of methyl groups on the conversion rate in a mixture of different DBT derivatives, we studied the HDS of a mixture of DBT, 4-MDBT and 4,6-DMDBT (molar ratio" 1/1/1) over the Mo/A1203 catalysts (MOO3" 6 wt%). For comparison, the HDS of a mixture of DBT, 2-MDBT and 2,8-DMDBT (molar ratio: 1/1/1) has been performed under the same reaction conditions. The relative activity obtained at 573 K and a H2 pressure of 3 MPa, for the HDS reaction of DBT derivatives over the Mo/A1203 catalysts under investigation is depicted in Table 1. As expected, the conversion for the different sulfur compounds increases as follows: DBT > 4-MDBT > 4,6-DMDBT. This leads to the assumption of a competitive mechanism between these sulfur compounds. The electron density at the sulfur atom is enhanced in the case of all methyl-substituted DBT derivatives. Nevertheless, concerning the hydrogenolyis pathway, the steric hindrance retarding the C-S bond occurs only when the methyl groups are at 4 or 4,6 position. However, the relative activity of 2-MDBT and 2,8-DMDBT in a mixture with DBT are as follows: DBT < 2-MDBT < 2,8-DMDBT. Due to the higher electron density located at the sulfur, and the less steric hindrance retarding the C-S, the highest conversion over all catalysts under investigation was found for 2,8-DMDBT. The results suggest that the methyl migration from 4- and/or 6- positions of 4,6-DMDBT may enhance the HDS activity. To increase the isomerization activity and hydrogenation activity over the HDS catalyst would be preferable for future developments of HDS catalyst. With regard to DBT, the conversion rate over the different catalysts increases as follows: Mo/A1203 < Mo/TiO2 < Mo/TiOz-AI203 (10.2 wt% TiO2), that are shown in Figure 3. Concerning the conversions of 4-MDBT and 4,6-DMDBT, all catalysts revealed lower conversion rates than that of DBT, because of the lower reactivities of these compounds. However, with regard to 4-MDBT and 4,6-DMDBT, the ratios of the methyl-substituted
135
CHB/BP ratios detected over the investigated catalysts are increased, whereas observed conversions are significantly higher over Mo/TiO2and Mo/TiO2-A1203 than that of Mo/A1203. However, most remarkable is the very high catalytic activity of the TiO2-A1203 supported molybdenum, reaching conversions of 50 % for the HDS of 4,6-DMDBT. Over Mo/A1203, 4,6-DMDBT conversions of only 34 % are obtained. The high conversions for the HDS of 4,6-DMDBT over Mo/TiO2-A1203 can mainly be attributed to the high ratios of corresponding CHB/BP ratios, which become higher than those obtained over Mo/TiO2 and Mo/A1203. This implies the hydrodesulfurization route to be more important for the HDS of 4,6-DMDBT over Mo/TiO2-A1203 compared to the A1 203 and TiO2 supported Mo catalysts. As already pointed out above, prehydrogenation of an aromatic ring reduces the steric hindrance by methyl groups during the C-S bond scission and leads to the corresponding CHB derivatives.
Figure 3. HDS activities for DBT, 4-MDBT, and 4,6-DMDBT over Mo catalyst on various supports: Reaction conditions" temperature; 573 K, pressure 3 Mpa, LHSV; 12.08-16.07 h-l, H2 flow rate; 200 cm3 min-1, MOO3; 20 wt% Judged by the detected conversions discussed above, over TiO2-A1203 and TiO 2 supported Mo catalysts, the hydrodesulfurization route is much more important than over the material supported on A1203. According to Mochida et al. [32], the prehydrogenation of an aromatic ring leads to higher reactivities of 4-MDBT and 4,6-DMDBT. This results in comparably high conversions for the HDS of 4-MDBT and 4,6-DMDBT over the Mo catalysts supported on the TiO2-A1203 composite material. In agreement with Kabe et al. [33], the reactivity of the sulfur compounds under investigation decreases in the order : 2,8-DMDBT > 2-MDBT > DBT > 4-MDBT > 4,6-DMDBT. This is true for all investigated catalysts. 4. CONCLUSIONS We prepared TiO2-A1203 composite supports by CVD method. The HDS reactions of DBT, 4-MDBT, and 4,6-DMDBT were carried out over molybdenum catalysts. The conversion rates obtained over Mo/TiOz-AI203 catalyst were much heigher than that obtained over Mo/A1203. With regard to the HDS of 4,6-DMDBT, this catalyst revealed higher conversion than Mo/TiO2. According to the higher CHBs/BPs ratios obtained over Mo/TiOz-AI203, the hydrodesulfurization route was promoted (see Scheme 1). XPS
136
investigations of catalysts before and after sulfiding suggeat that the interaction between Mo and y-A1203 is stronger than that between Mo and TiO 2-A1203, and also suggest that the reducibility from oxidic to sulfided Mo species on the TiO2-A1203 is higher than that on the y-A1203 support. Therfore, the number of active sites for HDS has increased on the surface of TiO2-A1203 supports. 5. R E F E R E N C E S
1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11. 12. 13. 14. 15. 16. 17. 18. 19. 20. 21. 22. 23. 24.
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
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Sulfur uptake, exchange and HDS activity of NiMoOx/Al203 catalysts Tamfis Koltai, Maria Dobrovolszky, Phi Trtrnyi Institute of Isotope and Surface Chemistry, Chemical Research Center, Hungarian Academy of Sciences H-1525 Budapest P. O. Box 77 Hungary Abstract The sulfur uptake and exchange of sulfur bonded to catalyst with gas phase H2S has been investigated by isotope tracer with two NiMoOx/AlzO3 catalysts of different alumina support. It was found that the higher maximum extent of sulfur uptake was paired with higher sulfur exchange and HDS activity. 1. INTRODUCTION The correlation between sulfur- catalyst interaction and hydrodesulfurization (HDS) activity is very much discussed in the literature. Chianelli and his colleagues found optimum catalytic activity at medium metal-sulfur bond strength of supported metal catalysts [1] and explained the HDS activity on the basis of the covalent ratio of that bond [2]. Norskov et. al. also found a correlation between HDS activity and sulfur- metal bond strength [3-5]; Kogan et. al attribute a decisive role to the mobility of sulfur in the activity of catalyst in HDS reactions [6]. In some of our previous communications we reported that higher uptake of sulfur was paired with higher HDS activity of alumina supported NiMoOx in comparison with alumina and amorphous silica - alumina supported NiWOx catalyst with lower HDS activity and sulfur uptake [7, 8]. No definite correlation has been found however between sulfur uptake and catalytic activity for a greater number of Mo-based metal promoted catalysts [9]. These studies were performed with slightly sulfided catalysts, and only small differences in uptake of sulfur were observed between the catalysts. With the aim of clarifying the correlation between the propensity to take up sulfur, the mobility of irreversibly bonded sulfur, and the catalysts' behavior (activity and selectivity) we determined the maximum amount of sulfur uptake, studied the exchange between gas phase H2S and sulfur bonded irreversibly to the catalyst, and the HDS by two different alumina supported NiMoOx catalysts. 2. EXPERIMENTAL 2.1. Catalysts Sample No. 1 was prepared in the Eindhoven Institute of Technology by impregnation of Ketjen 0.01 - 1.5 E y - A1203 (324 m2/g specific surface) with an aqueous solution of (NH4)2Mo7024,of Ni(NO3)2, and Nitrilo Triacetic Acid. The sample was dried for 8 hours at 393 K and calcined in air for 2 hours at 723 K [ 10]. Sample No. 2 was prepared in the Chemicals and Fuels Engineering Dept., University of Utah. A commercial 7 - A1203 (Ketjen) of 180 m2/g was impregnated consecutively with a solution of ammonium paramolybdate and nickel nitrate. The wet catalyst was dried, then calcined in air for 8-10 hours at 773 K [11]. The Mo and Ni content of both samples was determined by Prompt Gamma Activation Analysis (PGAA) at the Budapest Neutron Center. The characteristic physical data are collected in Table 1.
138
Table 1. Characteristic data of the NiMOOx/AI203 samples nM(1017/m8) Mo Ni Sample No. 1 4.77 1.60 Sample No. 2 5.08 . 3.03
H/(Ni+Mo) [-] 0.63 0.40
Surface area
m2/8 287 1 1 7 ...............
It is well known that at low (< 20 m%) concentration MoOx forms a monomolecular layer on alumina. From H/YM values we calculate the total number of oxygen ions on the surface [12, 131. In order to decide, whether metallic nickel was present on the surface of alumina the activity of the catalysts was checked in the dehydrogenation of cyclohexane. It was found that Sample No. 1 was not active in this reaction below 673 K, whereas Sample No. 2 catalyzed the cyclohexane-benzene conversion at 573 K. This indicated in agreement with literature data [ 14, 15] that metallic nickel was present on the surface of Sample No. 2 only. The activity of this sample was equal (TOF: 1.29 s1) with that of Ni/Al203 catalyst of 2.3 atom% ofNi (TOF: 1.16 sl), whereas a Ni/A1203 catalyst of 1.66 at% of Ni was not active at all. This can explain why Sample No. 1 was not active in this reaction.
2. 2. Sulfur uptake and exchange measurements A glass microanalytic pulse apparatus described elsewhere [9, 16] was utilized for mild sulfidation. The sulfur uptake from 6 radioactive H235S pulses (0.1 ml, 4.8 91017 S atoms) was measured. The amount of H2S retained by the catalysts was determined from the balance of radioactivity. Total sulfur uptake and exchange were studied by sulfidation of catalyst samples (24 mg each) with a H2/H235S mixture circulating with a flow rate of 9.68 NTP ml/s through the sample. The gases were mixed in a 145 ml vessel at NTP. The total pressure of the mixture was ~ 400 ton-, whereas the partial pressure of H2S 30 and 200 torr. Special measurements indicated that the total sulfur uptake became constant at ~ 30 ton.. The number of H2S molecules in the vessel was in the range of(1.25 - 8.50). 102~ The amount of sulfur taken up by the catalyst was followed by a decrease in the radioactivity in the gas phase measured at 3 minute intervals. The count rate in the gas phase reached a constant value at 60 minutes; however, the circulation in every sulfidation experiment was continued for 120 minutes. The amount of total sulfur uptake (Stug) has been calculated by expression t
rev
irr
I st
Sug -=S~g + S . =-~0 -m.~s
(1)
Here: L0 and I~t- the counting rates in the gas phase before and at the end of sulfidation respectively m.~s- number of H2S molecules in the gas phase with I~0 counting rate S~, the reversibly retained sulfur was then removed by treatment in vacuum and the sulfidation was repeated up to constant I'st in the gas phase. The amount of irreversibly bonded sulfur has been calculated as Irr t f st Su~ -Sag - ]-~-0.m.~s (2) Repeated treatment in vacuum was followed by treatment with a non radioactive Hz/H2S mixture, and the increase of gas phase radioactivity was measured. The sulfur exchange has been calculated by the expression:
139
$2 ~ - ~I- - 2-- - ~I.om H . Is
s
(3)
where: I~ and I0 are respectively the maximum and initial (background) counting rates in the gas phase; Is: the counting rate related to a definite rams. (Publication with details of this method is currently in preparation) Special measurements indicated that no radioactive sulfur could be removed with H2 of 400 ton pressure at 673 K. Another series indicated that the S~c values were independent of H2S partial pressure in the region pros -> 30 torr [17].
2. 3. HDS experiments The conversion of thiophene was studied in a pulse system, described in detail elsewhere [9]. The sample mass was 10-2 g; the amount of thiophene pulses injected into the H2- stream (flow rate: 30 NTP ml/min) was 5 - 10-4 ml - 3.82 9 10 TM molecules. HDS conversion of thiophene was calculated from the ratio of non-converted thiophene measured by GC v~ith a 5 m long 18 % squalane column on Chromosorb P at 383 K. The HDS acti,~iw of the samples was determined in mildly reduced, (by six H2S pulses) and strongly (treated for two hours in H2/H_~S stream of 0.1 v/,,% H_~S) sulfided form. Special measurements indicated [17] that the sulfur uptake by Sample No. 1 was in the order of 10 TM S/rag for 55 minutes of treatment in these conditions, i. e. -- 1.1 S/(Mo+Ni). It will be seen that this was - 60 ot; of the total uptake of sulfur, consequently the catalysts submitted to this procedure for two hours could be regarded as strongly sulfided ones. Also measured were the thiophene conversion values of totally sulfided samples in the circulation system. Data on the degree of conversion were not comparable with those measured in the pulse system. However, the selectivity values were of interest. 3. RESULTS AND DISCUSSION
3. 1. Sulfur uptake Data on sulfur uptake by Samples Nos. 1 and 2 are collected in Table 2. S - uptake by alumina supports was 3 and 11% of the total uptake by Samples Nos. 1 and 2 respectively, but the uptake decreased after the second pulse. Table 2. Sulfur uptake data (10 I7 molecules/mg) From 6 HzS pulses Maximal 623 K 673 K 673 K Si~/(Ni+Mo) Irrev. Rev. Irrev. Rev. Irrev. Rev. [-] Sample No. 1 1.2 0.1 1.5 0.2 8.51'2 3.1 1.3 Sample No. 2 1.6 0.5 2.1 0.3 14.01"2 10.0 1.7 1 pi~s = 30 torr; (at pros = 200 torr for Sample No. 1 S~ =8.8) 2 in molecules/nm2:3.0 and 12.0 respectively Table 2 shows that the irreversible uptake from 6 pulses reaches only 8 - 10 % of the maximal one. Fig~are 1 indicates the pulse by pulse uptake by Sample No. 2.
140
5 o
14~-
~-"
12 ill
i
=~=6-
I
II
9Cumulated total
, '
I Cumulated irrev.
~o4WE S 2-
o
o 0
1
2
3
4
5
6
Pulse Number
Figure 1. H2S uptake by Sample No. 2 at 673 K It is seen that the irreversible uptake decreases with increasing number of pulses whereas the amount of reversible uptake increases. This is a general tendency of uptake dynamics observed with other catalysts too [16]. The data in Table 2 indicate that the total uptake by Sample No. 2 was substantially higher than that by Sample No. 1. It is noteworthy, that both the amount of reversible sulfur uptake and its ratio to the irreversible one is substantially higher for Sample No. 2 than that for Sample No. 1 in spite of the much higher BET surface of Sample No. 1. This indicates that the reversibly bonded sulfur can not be regarded as a purely physisorbed one. 3. 2. Sulfur exchange Sulfur exchange data, presented in Table 3. indicate a substantially higher extent of exchange
Table 3 Sulfur exchange data T = 673 K: pros = 30 torr S~ S~~/(Ni+Mo) Sample No. 1 Sample No. 2
[ 1017molecules/mg] 1.98 3.43
0.31 0.42
sL
sE~ , s~L
[mol/nm 2] 0.69 2.93
[%] 23.3 24.4
by Sample No. 2 than that by Sample No 1, comparison with data in Table 2 indicate that the higher sulfur uptake is paired with higher extent of exchange. It is remarkable however, that the ratio of exchange to irreversible uptake are practically equal for both catalysts. This indicates the nearly equal stren~h of sulfur catalyst interaction and, presumably, the identity of structure of NiMoS species. It follows from this, that both irreversible sulfur uptake and exchange are determined by extensive properties of the samples. This is in agreement with the identical temperature dependence of the irreversible uptake: increase of 25 and 3 1 % for a 50 K temperature increase (Table 2). 3. 3. Conversion of thiophene The extent of converted thiophene has been related to catalyst mass, to the total number of Ni + Mo (TOF), and to the "active" part of the surface (Table 1, Column 6). Data on HDS are presented in Table 4.
141
The extent of converted thiophene has been related to catalyst mass, to the total number of Ni + Mo (TOF), and to the "active" part of the surface (Table 1, Column 6). Data on HDS are presented in Table 4. Table 4. Conversion of thiophene in pulse system T= 673 K, mass of catalyst 10-2 g State Reduced Sulfided (6 H2S) Sulfided (2 hours) Units [ 1017 TOF [mol/ [ 1017 TOF [mol/ [ 1017 TOF [mol/ mol/mgs] [s"1] nm2s] mol/mgs] [s"1] nm2s] mol/mgs] [sq] nm2s] Sample 1 1.05 0.17 0.37 1.66 0.30 0.56 1.26 0.199 0.44 Sample2 2.70 0.33 2.31 3.22 0.40 2.75 2.75 0.340 2.37 From the data in Table 4 it can be seen that the HDS activity of Sample No. 2 is substantially higher than that of Sample No. 1, such as its higher extent of sulfur uptake and exchange. Slight sulfidation enhances the activity of both samples; it is decreased, however, by further sulfidation. It follows from this, that the higher-in comparison with Sample 1- HDS activity is not a consequence of higher actual sulfur content: the positive correlation between sulfur uptake, exchange and catalytic activity indicates that the propensity of the catalyst to take up sulfur determines its activity in sulfur exchange and HDS, and this is presumably connected with the same -or similar- property of the catalyst. From the mechanism of sulfur uptake [18] and exchange [19] it follows that higher concentration of anionic vacancies is paired with higher extent of sulfur uptake and with that of exchange. Positive correlation between HDS activity and surface anion vacancies concentration was found in a number of investigations [20-22], in agreement with Topsoe's taodel of thiophene adsorption on the vacancy by sulfur atom [23]. The increased activity of Hz/HzS treated samples in comparison with H2-treated ones is presumably a result of direct exchange of sulfur with catalyst oxygen [ 18] followed by vacancy formation at the expense of surface sulfur atoms removed as H2S. With higher extent of sulfur uptake, however, the number of anionic vacancies can be reduced due to their occupation by sulfur. This explains the decreased lIDS activity in the case of strong sulfur treatment. There are two possible explanations of the higher sulfur uptake, the extent of exchange, and the HDS activity of Sample No. 2 in comparison with Sample No. 1 1. the higher Ni:Mo ratio in Sample No. 2 induces higher amount of anion vacancies [24] 2. the presence of Ni(O) on the surface. This possibly leads to higher extent of Ni-Mo-S surface phase resulting in higher HDS activity [25]. This is in agreement with the generally accepted view that the catalytic activity does not depend on the total amount of the promoter: it depends on the extent of Ni-Mo-S structures [26, 27]. In Table 5, data are collected on the distribution of hydrocarbon products of thiophene conversion
142
Table 5. Distribution of hydrocarbon products (molecular percentage) of thiophene conversion Treatment THT Bu BulBu2'(C4 Bu2-/BUl555 12.8 2.53 H2 2.9 6.8 21.9 4.9 ,,.'~76 6 H2S 1.8 8.3 22.6 62 4 Sample 1 2 h H2/H2S 5.2 65 1 1.2 2.72 4.6 23.9 1.8 2.74 Circ. 2' 0 7.0 24.4 66.8 46.4 2.3 2.36 Circ. 5' 0 31.6 19.7 47.2 4.2 6.65 Circ. 20" 0 41.4 7.1 Circ. 80' 1.5 43.5 11.1 35.5 8.4 3.20 H2 0.4 7.9 20.1 56.3 15.4 2.80 6 H2S 0.7 10.3 21.2 60.4 7.3 2.85 6.6 23.4 65.5 3.8 2.80 Sample 2 2 h H2/H2S 0.7 Circ. 2' 0.3 8.7 23.8 63.4 3.8 2.66 Circ. 20' 0.1 15.1 21.0 53.3 8.6 2.54 It should be added to the data in Table 5 that in the circulation system the pattern of ~(C4 products indicates a deficit in carbon balance: the number of C1- hydrocarbons is lower than that of C3-hydrocarbons, and the deficit increases with increasing circulation time. This indicates the formation of carbon deposits on the catalysts -this was another reason for not comparing HDS conversion degree data obtained in the circulation system with those obtained in the pulse one. It should to be noticed, that conditions applied at industrial hydrotreatment processes are totally different (flow system, high pressure) from those, applied here: circulation system, 53 kPa total, 5-6 kPa partial pressure of thiophene. The data in Table 5 indicate no significant differences between the two NiMo/AI203 samples with respect to pattern of products. This supports the earlier-expressed concept: the differences in HDS-activity of the two samples are caused by differences in their extensive properties; the strengths of catalyst-sulfur interaction are near to equal and the surface structures are probably- identical. The high Buf/BUl: ratio indicates a low ratio of threefold coordinately unsaturated (3M) sites, responsible for but-1-ene formation [28-30]. 4. CONCLUSIONS There are substantial differences between the two Ni]V[o/AI203samples with respect to the extent of their sulfur uptake, their sulfur exchange, and their catalytic activity in HDS. The higher propensity to take up sulfur -expressed by the extent of maximum uptake- is paired with the higher extent of exchange and with higher catalytic activity. The existence of this correlation is explained from the viewpoint of Massoth's mechanism for sulfur uptake and exchange: S-uptake, exchange, and HDS activity depend on the concentration of anion vacancies of the catalyst.
143
ACKNOWLEDGMENTS
The study was supported in part by the Commission of the EU in the framework of the JOU2CT93-0409 and by the Hungarian National Res. Sci. Fund, (OTKAT 017051). The authors are grateful to Professors Frank Massoth (University of Utah) and V. H. J. de Beer (Eindhoven Institute of Technology) for kindly supplying the catalytic samples. The authors are indebted to Mr. V. Galsb,n for performing the cyclohexane experiments, and to Mr. K. Matusek for performing BET and TPR measurements. 5. REFERENCES
[1] [2]
[3]
[4]
[5] [6] [7]
[8] [9] [10]
[11] [12] [13] [14]
[15]
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[18]
[19]
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[21] [22] [23] [24] [25] [26] [27] [28]
T. A. Pecoraro and R. R. Chianelli, J. Catal. 67 (1981) 430 S. Harris and R. R. Chianelli, J. Catal. 86 (1984) 400 J. K. Norskov, B. S. Clausen and H. Topsoe, Catal. Lett. i 3 (1992) 1 H. Topsoe, B. S. Clausen, N. Y. Topsoe, J. Hyldost and J. K. Norskov 206th National Meeting, American Chemical Soc. Chicago IL, August 22-27, 1993 T. Braun, M. Holmgard, C. V. Ovesen, C. J. H. Jacobsen, I. V. Nilsen, E. TOrnqvist, B. S. Clausen and H. Topsoe, 6th Nordic Symposium on Catalysis, Hornbaek, Denmark, 1994. Book of Abstracts, Session 3A V. M. Kogan, A. A. Greish and G. V. Isagulyants, Catal. Lett. 6 (1990) 157 M. Dobrovolszky, T. I. Korhnyi, K. Matusek, Z. PaLl and P. T6t6nyi, 11th Int. Congr. Catalysis, programme Po. 370 Baltimore 1996 T. I. Korhnyi, M. Dobrovolszky, T. Koltai, K. Matusek, Z. Pab,l and P. T6t6nyi, Fuel Processing Technology, Special Issue (in press) M. Dobrovolszky, K. Matusek, Z. Paal and P. T6t6nyi, J. Chem. Soc. Faraday Trans. 89 (1993) 3137 W. R. A. M. Robinson, J. A. R. van Veen, V. H. J. de Beer, R. A. van Santen, Fuel Processing Technology, Special Issue (in press) F. E. Massoth, Personal communication Y. C. Xie and Y. Q. Tang, Adv. Catal. 37 (1990) 1 L. Tamaska, A. Redey, P. T6t6nyi, RKC Letters Vol. 65, No. 2, (1998) 285 A. Balandin, Adv. Catalysis 10 (1958) 96 P. T6t6nyi, Surface and Defect Properties of Solids, v. 5. p. 81, Chem. Soc. Burlington House, London 1976 M. Dobrovolszky, Z. Pahl and P. T6t6nyi, Appl. Catal. 142 (1996) 159 T. Koltai and P. T6t6nyi, Radiochim. and Radioanal. Acta (in preparation) F. E. Massoth, J. Catal. 36 (1975) 164 F. E. Massoth and P. Zeuthen, J. Catal. 145 (1994) 216 S. J. Tauster, T. A. Pecoraro and R. R. Chianelli, J. Catal. 63 (1980) 515 S. J. Tauster and K. L. Riley, J. Catal. 67 (1981) 250 T. A. Bodrero, C. H. Bartholomew and K. C. Pratt, J. Catal. 78 (1982) 253 H. Topsoe, in J. P. Bonnelle, B. Delmon and E. Derouane (Eds.), Surface Properties and Catalysis by Non Metals. Reidel, 1983, p. 329 P. Ratnasamy, A. V. Ramaswamy and S. Sikasauber, J. Catal. 68 (1981)433 A. Catafat, J. Laine, A-Lopez-Agudo and J. M. Palacios, J. Catal. 162 (1996) 20 H. Topsoe, B. S. Clausen, Catal. Rev. Sci. Eng. 26 (1984) 395 R. J. A. van Veen, E. Gerkema, A. M. van der Kraan, P. A. J. M. Hendriksand, H. J. Beens, J. Catal. 133 (1992) 112 K. C. Campbell, M. L. lVlirza, S. J. Thomson and G. Webb, J. Chem. Soc. Farad. Trans. I. 80 (1984) 1989
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
S T A B I L I T Y O F CoMo/A120 3 C A T A L Y S T S : CYCLES ON HDS
145
EFFECT
OF
HDO
Viljava, T.-R., Komulainen, S., Selvam, T. and Krause, A.O.I. Helsinki University of Technology, Department of Chemical Technology, P.O. Box 6100, FIN-02015 HUT, Finland E-mail:
[email protected]
Abstract
The effect of alternating hydrodesulfurization (HDS) and hydrodeoxygenation (HDO) cycles on the activity of a commercial presulfided Co-Mo/7A1203 catalyst was studied in a flow reactor at 250 ~ and 1.5 MPa by changing the feed from benzothiophene to phenol and back. For comparison, HDS and HDO reactions of these model compounds were studied separately, both on the presulfided and on the oxide catalyst and with or without the addition of a sulfiding agent, H2S or CS2, to the feed: Both the total conversion and the HDS conversion of benzothiophene decreased after a HDO period. The HDS reaction was affected more than the hydrogenation of benzothiophene to dihydrobenzothiophene. The HDS activity of the catalyst was, however, slowly recovered with time on sulfur-containing feed. A preceding HDS cycle was beneficial for the HDO cycle and the catalyst was more stable against deactivation than without the HDS cycle. Addition of sulfiding agents to the oxygen-containing feed decreased HDO conversion on the presulfided catalyst. The hydrogenolysis route to benzene was more suppressed than the route to cyclohexane and cyclohexene. It was not possible to activate the oxide form of the catalyst for HDO by adding sulfiding agents to the oxygencontaining feed.
1. I N T R O D U C T I O N
Hydrodesulfurization (HDS) has been an essential process in oil refining for over 50 years. Catalysts containing Mo and Co or Ni as promoters on 7-A1203have generally been applied in these processes [1]. The catalysts are typically activated by sulfiding either with H2S or by exposing the catalyst to the sulfurcontaining feed. The development of alternative raw materials for fuel and chemical production has led to increased interest in hydrodeoxygenation (HDO). Oxygen content is especially high in coal liquids and in liquefied biomass, which has
146
potential as a renewable substitute for oil. The direct use of biomass-based liquids as fuel is limited due to their high viscosity and poor storage stability related to their oxygen-containing components [2]. The sulfur content of these liquids is often negligible or very low [3,4]. So far, HDO has not been as thoroughly studied as HDS or hydrodenitrogenation (HDN). Model compound studies with different oxygencontaining reactants have, however, been reported. Most of these studies have been carried out in a batch reactor and on a Ni-Mo/A1203 or a Co-Mo/A1203 catalyst [6-9]. Studies concerning the stability of the presulfided catalyst during HDO are few [6,10]. Some H2S is probably needed to maintain the catalyst in its active sulfided state during HDO, when the feed does not contain sulfur, but the presence of higher H2S concentrations has a negative effect on HDO [1,6]. As a whole, the role of H2S during HDO is not yet clear. In addition, HDO reactions have been observed to be suppressed in the presence of a sulfur-containing functionality, either in the same or in a separate molecule with the oxygen functionality [5,11]. On the other hand, HDS reactions have also turned out to be retarded to some extent in the case of competitive HDS and HDO. If instead of simultaneous hydrotreating of sulfur- and oxygen-containing feeds, oxygencontaining feeds are planned to be used occasionally as substitutes for normal sulfur-containing feeds, information concerning the stability of the hydrotreating catalyst when the feed composition is changed would be valuable. Therefore, we have studied the effect of alternating HDS and HDO periods on the stability of a presulfided Co-Mo/7-A1203 catalyst. We have changed the feed from benzothiophene to phenol and back over 5-8 h periods. The HDO activity of the catalyst in these HDO cycles has also been compared with the activity of the catalyst in the presence of added sulfiding agents, H2S or CS 2.
2. E X P E R I M E N T A L
The catalyst was a commercial HDS catalyst (Ketjenfine 742-1.3Q, Akzo Chemie) containing 4.4 wt-% of CoO and 15 wt-% of MoO 3on 7-A1203. The catalyst was crushed and sieved to a fraction of 0.5-0.75 mm. 0.5 g of the catalyst was packed into a flow reactor (ID= 10 mm, L= 380 mm) between two layers of silicon carbide. The catalyst was calcined at 400 ~ under N 2 for 2 h. After calcination, the catalyst was either used as such in the hydrotreating tests or it was presulfided with 5 % H2S/H 2 (flow rate 2.5 1/h NTP) at 400 ~ for 4 hours. The temperature of the catalyst layer was decreased to the reaction temperature, 250 ~ and the total pressure was adjusted to 1.5 MPa. The liquid feed containing 3 wt-% of phenol (J.T.Baker, lab. grade) or benzothiophene (Fluka, >97%) in m-xylene (Merck, >99 %) was introduced to the reactor at a rate of 10 g/h. The gas feed rate was 2 l(NTP)/h. Liquid sampling was started after 2 h of liquid feed. The samples were taken at 30 min intervals and they were analysed by gas chromatography (HP 6890 A, flame ionisation
147
detector, capillary column DB-1). After the reaction, the catalysts were analysed for sulfur and carbon content using a Leco SC-444 analyser. The cycles of HDS of benzothiophene and HDO of phenol were carried out using different vessels and feed lines for the sulfur-containing and oxygencontaining feeds. The feed was changed at 5-8 h intervals. For comparison, HDS and HDO reactions of benzothiophene and phenol were studied separately, both on the presulfided and the oxide catalyst and with or without the addition of a sulfiding agent, either H2S or CS~, to the feed.
3. R E S U L T S A N D D I S C U S S I O N 3.1. H D S a n d H D O r e a c t i o n s o n f r e s h s u l f i d e c a t a l y s t
Ethylbenzene and dihydrobenzothiophene were the reaction products of benzothiophene, and benzene, cyclohexane and cyclohexene were the reaction products of phenol under the reaction conditions studied. The simplified networks for the compounds are presented in Fig. 1. The reaction pathways of the Hydrogenation (HYD) dihydrobenzothiophene I Hydrogenolysis (Ha) benzothiophene Hydrogenationhydrogenolysis (HYD-HG) Direct hydrogenolysis OH
(HG)
6
@ aromatics: benzene
phenol Hydrogenationhydrogenolysis (HYD-HG) Figure
ethylbenzene
O+0
non-aromatics: cyclohexane and-hexene
1. Simplified reaction networks of benzothiophene
and phenol.
148
reactants have been discussed in more detail elsewhere [1,5,6,11]. Ethylbenzene was the only desulfurized product of benzothiophene. It can be formed directly from benzothiophene and via the hydrogenated sulfur-containing intermediate, dihydrobenzothiophene. HDO of phenol proceeds via two separate routes: direct hydrogenolysis of the carbon-oxygen bond giving the aromatic reaction product (HG route) and the combined hydrogenation-hydrogenolysis route giving the nonaromatic products (HYD-HG route). Hydrogenation of benzene to cyclohexane and cyclohexene is negligible under the conditions studied. Both the activity and the selectivity of the presulfided catalyst stayed near constant during the hydrotreating of benzothiophene for 6-10 h. In contrast, a significant decrease in HDO conversion with on-stream time was observed in the reactions of phenol. HDO conversion decreased from about 35% to 26% in 10 h when no sulfiding agent was added to the reactor feed. The selectivity of the reaction pathways, however, stayed near constant: the HG route of HDO was preferred over the HYD-HG route and the average selectivity towards benzene was 91%. It has been proposed [6-9,12,13] that there are at least two kinds of active sites on a hydrotreating catalyst, one of them is responsible for hydrogenolysis and the other for hydrogenation reactions. The selectivities of the HDO routes, which do not depend on the time on stream, suggest t h a t if two
Table 1. Effect of sulfidation of the catalyst and addition of sulfiding agent on the HDO activity of the CoMo catalyst
Presulfided catalyst
Conversion of HG, %a 30.9
Conversion of HYD-HG, %a 3.0
Catalystb S, % C, % 7.2 4.3
Oxide catalyst
0.1
0
Oxide catalyst, 5 % H2S in the gas feed
1.1
1.2
7.8
6.4
Oxide catalyst, 0.13 % CS 2 in the liquid feed
0.8
0.4
3.7
5.1
Presulfided catalyst, 0.25 % H2S in the gas feed
5.9
2.1
6.0
5.3
Presulfided catalyst, 0.13 % CS 2 in the liquid feed
5.9
2.5
7.7
6.0
a Average for the samples with on-stream time between 2-4 h. b Catalyst analysed after the reaction. Calculated sulfur content 8.2 %.
149
kinds of active sites exist on the catalyst, these sites are deactivated at equal rates during HDO in the absence of sulfur. Due to the low reaction temperature used, no formation of heavier reaction products was detected and the coke content of the catalyst was low (See Table 1). Moreover, the sulfur content of the catalyst after an HDO period of 24 h was still about 87% of the calculated sulfur content of a properly sulfidated catalyst. The addition of sulfiding agents - H2S or CS 2 - to the oxygen-containing feed decreased the HDO conversion on the presulfided catalyst by about 75% (See Table 1). The HG route to benzene was more effectively suppressed t h a n the HYD-HG route, as observed also in our earlier study in a batch reactor [5]. These results give support to the dual site theory in which the active sites of the catalyst differ in their sensitivity towards sulfur.
3.2. A l t e r n a t i n g c y c l e s of H D S and HDO on sulfide c a t a l y s t When HDS of benzothiophene was carried out in alternating periods of 4-8 h with HDO of phenol, both the total conversion and the HDS conversion of benzothiophene decreased after each HDO cycle (See Fig. 2). The HDS reaction was more affected by HDO cycles than the hydrogenation of benzothiophene to dihydrobenzothiophene. The HDS activity of the catalyst, however, slowly recovered with time on sulfur-containing feed. From the slope of the curves of the 3:rd HDS cycle in Fig. 2, it could be predicted that the HDS activity is restored to its former level in about 20 h. This increase in HDS activity during hydrotreating of a benzothiophene-containing feed may be caused by slow resulfidation of the catalyst.
Figure 2. Effect of HDO cycles on HDS.
150
Figure 3. Effect of HDS cycles on HDO. The effect of HDS cycles on HDO was more complicated. When the experiment was started with an HDS cycle, the HDO conversion of phenol in the first HDO cycle was slightly higher than in a separate run with phenol in the absence of sulfur-containing compounds (See Fig. 3). Furthermore, no deactivation of the catalyst was detected during the l:st HDO cycle of 8 hours. In the subsequent HDO cycles, the total HDO conversion was about 4 percentage units lower after each HDS cycle. The selectivities were, however, unaffected. Moreover, the catalyst did not deactivate within the HDO cycles. The sulfur content of the catalyst after 3 cycles of HDS and HDO was the same, and the carbon content slightly higher, 6.8%, than in a separate run of 24 h with phenol. If the first cycle was HDO, the HDO conversion was diminished by about 7 percentage units and the selectivity towards benzene by about 4 percentage units after each HDS cycle. In summary, it seems that a preceding cycle of HDS is beneficial for the following HDO cycle because it stabilises the catalyst. HDO cycles decrease the HDS activity of the catalyst, but the activity recovers at least partially with time on sulfur-containing feed in the absence of the oxygen compound. 3.3. H D S a n d H D O r e a c t i o n s o n o x i d e c a t a l y s t The HDS activity of the oxide form of the catalyst was about 20% of the activity of the sulfided form of the catalyst (See Table 2). During HDS, as expected, the activity of the oxide form of the catalyst increased slightly with time on stream. The sulfur content of the catalyst after the HDS experiment of 12 h was about 65% of the calculated sulfur content for proper sulfidation. The efficiency of benzothiophene as a sulfiding agent would, however, probably be much higher at normal sulfidation temperatures. Addition of H2S to the reactor feed increased the total conversion of benzothiophene on the oxide catalyst, but
151
Table 2. Effect of sulfidation of the catalyst and addition of sulfiding agent on the HDS activity of the CoMo catalyst
Presulfided catalyst
Total conversion, % 87.8
HDS conversion, % 78.5
Catalyst a S, % C, % 8.0 6.3
Oxide catalyst
31.9 b
15.9 b
5.3
Oxide catalyst, 5 % H2S in the gas feed 49.6 b
15.0 b
4.4
Catalyst analysed after the reaction. Calculated sulfur content 8.2 %. b Average for the samples with on-stream time between 2-4 h. a
the HDS activity was of the same order of magnitude as in the absence of the extra sulfiding agent. The HDO activity of the oxide form of the catalyst was negligible (See Table 1). The addition of H~S or CS 2 to the reactor feed as a sulfiding agent increased the HDO activity of the oxide catalyst only very slightly, and the initial activity of the catalyst decreased gradually with the on-stream time. However, the sulfur content of the oxide catalyst markedly increased due to the presence of the sulfiding agent: In 12 h on the H~S-containing stream, it reached 95% of the calculated sulfur content of a properly sulfidated catalyst. Thus, it seems that the oxide catalyst can not be sulfided to the active form for HDO by exposing it on sulfiding agents during HDO, even though the sulfur content of the catalyst markedly increases. 3.4. C a t a l y s t d e a c t i v a t i o n As a cause for the deactivation of the sulfided catalysts during HDO, at least three factors have been suggested: water formed in the HDO reactions, coke or high molecular weight reaction products formed from the oxygen-containing molecules, or replacement of the catalytic sulfur with oxygen [14]. Non-selective blocking of the active sites by water or coke could partly explain the similar deactivation of the both reaction routes in HDO of phenol on the sulfided catalyst in the absence of H2S. This deactivation should also be seen during the HDO cycles carried out after the HDS cycles. In these cycles deactivation was, however, negligible indicating that other deactivation mechanisms must affect the activity, too. Replacement of the catalytic sulfur with oxygen is seen in the sulfur contents of the catalysts after the reaction: during HDO of phenol the sulfur content of the presulfided catalyst decreased by 12% when no sulfur was added to the feed. However, resulfidation of the catalyst during HDS cycles recovers most of the activity lost by this mechanism. It has been suggested [14] that the oxygen-containing compounds adsorbed on the catalyst in the presence of diminished hydrogen surface concentrations
152
would polymerise to high molecular weight species. These species may be the reason for the decrease in the HDO conversion detected between the HDO cycles before and after an HDS cycle (See Fig. 3): During a short period after changing the sulfur-containing feed to the oxygen-containing feed, HDS and HDO are competing. Due to the higher reactivity of the sulfur compound there is not enough active hydrogen available for HDO, and the adsorbed oxygen compounds can polymerise decreasing the activity of the catalyst. This type of deactivation can probably not be recovered even with longer periods of HDS.
4. C O N C L U S I O N S If an HDS process should treat sulfur-containing and oxygen-containing feeds sequentially, the efficiency of HDS can be expected to decrease after a preceding HDO period, but the HDS activity of the catalyst is slowly recovered with time on sulfur-containing oxygen-free feed. The catalyst for HDO should be presulfided, because it is not possible to activate the oxide catalyst during HDO by adding sulfiding agents to the oxygen-containing feed. A preceding HDS cycle is beneficial for the stability of the catalyst during HDO. Addition of low molecular weight sulfiding agents to the feed during HDO does not increase the stability of the presulfided catalyst, but such agents could be used to adjust the reaction selectivities in favour of the formation of the non-aromatic reaction products.
REFERENCES
1. H. Topsoe, B.S. Clausen and F.E. Massoth, Hydrotreating catalysis. Science and Technology, Springer-Verlag Berlin Heidelberg, 1996. 2. R. Maggi and B. Delmon, Stud. Surf. Sci. Catal., 106 (1997) 99. 3. B. Gevert, Upgrading of Directly Liquefied Biomass to Transportation Fuels, PhD Dissertation, Chalmers University of Technology, 1987. 4. V.K. Bathia, K.V. Padmaja, S. Kamra, J. Singh and R.P. Badoni, Fuel, 72 (1993) 101. 5. T.-R. Viljava and A.O.I. Krause, Stud. Surf. Sci. Catal., 106 (1997) 343. 6. M.J. Girgis and B.C. Gates, Ind. Eng. Chem. Res., 30 (1991) 2021. 7. E. Laurent and B. Delmon, Ind. Eng. Chem. Res., 32 (1993) 2516. 8. E. Laurent and B. Delmon, Appl. Catal. A, 109 (1994) 77. 9. E. Laurent and B. Delmon, Appl. Catal. A, 109 (1994) 97. 10.V. LaVopa, Catalytic hydrodeoxygenation of benzofuran in a trickle bed reactor: Kinetics, poisoning, and phase distribution effects, PhD Dissertation, Massachusetts Institute of Technology, 1987. ll.T.-R. Viljava and A.O.I. Krause, Appl. Catal. A: General, 145 (1996) 237. 12.B. Delmon and G.F. Froment, Catal. R e v . - Sci. Eng., 38 (1996) 69. 13.B. Delmon, Bull. Soc. Chim. Belg., 104 (1995) 173. 14. E. Furimsky, Catal. R e v . - Sci. Eng., 25 (1983) 421.
HydrotreatmentandHydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
153
CoMo/A120~ a n d CoMo/TiO2-AhO~ c a t a l y s t s in h y d r o d e s u l f u r i z a t i o n : r e l a t i o n s h i p b e t w e e n t h e p r o m o t i n g e f f e c t of c o b a l t a n d t h e n a t u r e o f the support.
M.Vrinat*, D. Letourneur, R. Bacaud, V. Harl6, B. J o u g u e t and C. Leclercq
Institut de Recherches sur la Catalyse, 2 Av. A. Einstein, 69626, Villeurbanne Cedex,.Fr.
Abstract Molybdenum and cobalt-molybdenum catalysts have been prepared over A1203 and a TiO2-A1203 mixed oxide in order to get more insight on the relation between the nature of the support and the promoting effect of cobalt. These catalysts were characterized by UV-visible diffuse reflectance spectroscopy (DRS), analyzed by energy dispersive spectroscopy (EDS) and evaluated in dibenzothiophene hydrodesulfurization. For the unpromoted catalysts, the Mo/TiO2-A1203 sample presented a very high activity as compared to the Mo/A1203 catalyst. However, over the titania based support a limited activity enhancement by Co promotion was observed. Such a drawback of the TiO2-A1203 support was correlated to a large amount of cobalt involved in the formation of CoMoO4; moreover, variations observed in the stability of the sulfide phase support these explanations.
1. I N T R O D U C T I O N
Sulfided CoMo catalysts are used for a long time for hydrotreating processes and these catalysts have generally been prepared over alumina. However, recent environmental specifications regarding the diminution of sulfur and aromatics in transportation fuels have spurred active development of investigation to improve activity and selectivity of such industrial catalysts. As reported in reviews, severals studies suggested that changing the support could improve catalyst activity [1,2]. In this way, a lot of work has been done on TiO2 since it has been shown that molybdenum disulfide supported over such oxide present an intrinsic activity three up to four times higher than over alumina [3,6]. Such interest in the use of TiO2 has been also demonstrated using mixed oxide supports with a
154
rich TiO2 content [7-9]. In the case of MoS2 supported over TiO2-A12035%, a support with the composition TiO2(95%)A1203(5%), the very high activity as compared to MoS2 supported over AleO3 could not be correlated simply with variations in the morphology (length and stacking) of the MoS2 crystallites. Large variations in the reducibility of the supported sulfide were observed, the reducibility of the MoSe phase being higher when supported over the mixed oxide [8]. Although such oxides appeared promising when used as supports of MoS2, the promoting effect induced by cobalt or nickel is generally lower t h a n t h a t observed over A1203, which fact is still to be understood. The present contribution deals with the comprehension of the level of this synergetic effect by comparison of Mo and CoMo catalysts prepared over a commercial A1203 and over a TiO2-A12035% mixed oxide prepared by coprecipitation of aluminium and titanium isopropoxydes. These catalysts were characterized by UV-visible spectroscopy, analyzed by energy dispersive spectroscopy and evaluated in the hydrodesulfurization of dibenzothiophene (DBT).
2. EXPERIMENTAL 2.1. Supports and catalysts preparation Preparation of the TiO2-A1203 support with the molar ratio AleOd(TiO2+A1203)=0.05 has been described in detail previously [9]. Briefly, the method involved the coprecipitation of aluminium and titanium isopropoxydes dissolved in n-propanol by adding acidified water in a large excess. The precipitate was left under slow agitation for 24 h, filtered, washed with demineralized water, dried at 383 K during 24 h and then calcined for 4 h at 823 K. The resulting solid was named TiOe-A12035% and presented a surface area of 160 m2/g as determined by nitrogen physisorption using an automated BET apparatus. Molybdenum was deposited on the carrier by the pore volume method using an aqueous solution of ammonium heptamolybdate of appropriate concentration in order to obtain a solid with 2.8 atoms of molybdenum per square nanometer of support. The impregnated sample was then dried at 393 K and later calcined at 673 K for 4 h (heating rate 2 K/min). In the case of CoMo catalyst cobalt was introduced by coimpregnation using cobalt nitrate, and the molar promoter to molybdenum ratio r =Co/[Co+Mo] was equal to 0.3. The Mo/AleO3 catalyst was prepared according to the same procedure using an industrial 7AleO3 support (surface area 232 me/g) and the CoMo/A1203 catalyst was an industrial sample (MOO3=14 wt%, CoO 3 wt%) having the same A1203 support.
2.2. Catalysts characterization Diffuse reflectance spectra (UV-vis DRS) of the catalysts in their oxidic form were obtained in Perkin-Elmer Lambda 9 spectrophotometer using BaSO4 as a reference.
155
EDS analyses were performed on a J E O L 2010 FEG electron microscope equipped with a Link Isis microanalysis system. The H2-H2S sulfided catalyst was dispersed in an ethanol solution and a droplet of the suspension was deposited over a holey carbon film supported on a copper grid. The K lines of S, Co, O, A1 and the L lines of Mo were selected for quantitative analyses. 2.3. C a t a l y t i c a c t i v i t i e s The HDS of dibenzothiophene was carried out in a continuous flow high pressure microreactor working under a total pressure of 30x10 ~ Pa, with a partial pressure of DBT of 4.80x102 Pa and H2S added to the feed (252x10 ~ Pa). Experiments were performed between 533 and 588 K. The HDS activity was expressed by the pseudo first order rate constant calculated after 16 h time on stream at the pseudo stationary state, by the following equation : k = - ~ ~0 x ln(1- x), mxC 0 where k is the rate constant of the reaction (1.g-l.s-l), Fo the molar flow rate of DBT (molls) at the inlet, x the conversion of reactant, m the weight of catalyst (g) and Co the concentration of DBT (moll1). Prior to catalytic tests the samples were sulfided ex situ under H2-H2S (10%) mixture for 4 h at 673 K.
3. R E S U L T S AND D I S C U S S I O N 3.1. Activities
9
Results given in Figure support leads to a catalyst with the CoMo samples, we observed when the TiO~-A12035% support
1 indicate t h a t in the Mo series, the TiO2 rich a high activity as compared to the Mo/A1203. For a dramatic decrease of the promoting effect of Co is used.
Figure.1. HDS activities compared at 573 K .
156
These results confirm previous ones obtained for NiMo catalysts prepared over the same supports and evaluated in aromatic ring hydrogenation (8). Such a low synergetic effect could be explained by a difficult formation of the CoMoS phase during the sulfidation. To answer that question the catalysts have been characterized by UV diffuse reflectance spectroscopy and EDS analysis. 3.2. Diffuse reflectance s p e c t r o s c o p y : The DRS spectra of catalysts in their oxidic state, supported on TiO2A12035 % and A1203, are presented in Figure 2. The intense adsorption band at about 240-340 nm, recorded for all samples, could be in first analysis attributed to the ligand-metal charge transfer 02. -~ Mo ~+ [10, 11]. For the aluminasupported sample, this band would correspond to tetrahedral (250-280 nm) and octahedral (290-330 nm) forms of Mo(IV). However, in the TiO2-A1203 supported catalyst, the adsorption at 365 nm is due to the absorption edge of the 02. --->Ti 4+ charge transfer band [9], and it is not possible to clearly determine the changes occurring in the molybdenum species, which appear in the 250-330 nm wavelengths range. For the CoMo/A12Oa catalyst a broad triple band in the 500-700 nm region is observed. This band could be attributed to ligand field transition of tetrahedral Co(II) [12, 13]. In this catalyst this band has been assigned to tetrahedral Co(II) in COA1204. For the TiO2-A12035% support, the probability to incorporate cobalt into the A1203 to form COA1204 is strongly reduced due to the low content of this oxide. This could explain why this band is very weak for the titania based support. In the region 1100-1800 nm, the bands appearing correspond to both octahedral Co(III) and tetrahedral Co(II) in oxygen surrounding. For the CoMo/A1203 catalyst, this octahedral Co(III) could not be attributed to the presence of oxide compounds like CoO, Co304 or CoMoO4 [14 - 17]. This point is confirmed by the blue colour of the catalyst. This octahedral Co(III) would be located near the surface in strong interaction with the molybdates. That interaction is not well defined and Topsoe and Topsoe [15] proposed t h a t the octahedral Co(III) would be present as CoO6 octahedra associated with the octahedrically coordinated Mo atoms located in one dimensional chains on the alumina surface. The spectrum of CoMo/TiO2-A12035% is very different from the one of CoMo/A1203. A very weak band is observed between 1100 and 1800 nm and only a little shoulder appears in the region 400-700 nm. It has been proposed that the low promoting effect of cobalt on CoMo/TiO2 could be explained by a migration of cobalt in the titania lattice, to form an inactive surface phase, probably anamorphous titanate [3]. To check t h a t possibility a CoTiO3 titanate was prepared as a reference. The spectrum of CoTiO3 is given on Figure 2 and, comparison with that of CoMo/TiO2-A12035% clearly indicates that cobalt atoms do not migrate in the titania structure. A Co/TiO2-A12035% sample with the same amount of cobalt t h a n the CoMo/TiO2-A12035% was also prepared and its spectrum given in Figure 2 appears very different. In the region 1100-1800 nm, bands prove the presence of tetrahedral and octahedral cobalt confirmed by the large band near 700 nm and the shoulder near 400 nm [7].
157
CoTiO 3
A CoMo/Co/TiO2-A1203 5%
Co/TiO2-A1203 5%
CoMo/TiOz-A1203 5%
CoMo/A1203
TiO2-A1203 5%
200
I
I
I
400
600
800
I
I
I
I
I
1000 1200 1400 1600 1800 2000
(nm) Figure 2. DRS spectra of various catalysts supported on A1203 or TiO2-A12035% This spectrum confirms that cobalt is not incorporated in the titania lattice but that cobalt oxide Co304 is formed. This point is proved by the black colour of this catalyst. When Co and Mo are coimpregnated on Co/TiO2-A12035% only small
158
changes in the 400-800 nm region are observed. The same compound as in CoMo/TiO2-A12035% is formed. It could be CoMoO4 species in which cobalt has an octahedral structure (low intensity). This compound is known to be difficult to sulfide and to give low activity catalysts [16, 18-19]. This result could explained the low promoting effect for the catalysts supported on TiO2-A12035% by a difficult formation of the CoMoS phase due to the presence of CoMoO4 in the oxidic precursor. This oxide is easily detected by its light green colour.
3.3. EDS analysis of sulfided catalysts. Our DRS experiments have clearly demonstrated that in calcined CoMo precursors the cobalt atoms located at the surface of the support change with the nature of this oxide. Therefore, variations in the nature of the actives species of the sulfided catalysts are expected. To check that point our samples have been analyzed by EDS. The use of probe size of 2nm allow to focus the beam over the MoS2 slabs. Results given in Figure 3 indicate the Co/(Co+Mo) ratios obtained for numerous analyses over the two samples. It is clear that the cobalt is not so homogeneously dispersed over the TiO2-A12035% support than over the alumina, suggesting that on the former support the formation of the CoMo surface phase precursor of the active CoMoS phase is more difficult than over alumina.
~" + O ~" o o
0,8 0,7 -~ 0,60,5 0,4 ' 0,3 ~ 0,2
CoMo/AI203
o,1 i ~ , r 1 6 2 0
~
!
i
1
,0'1
3
~
5
0,8 0,7 "S" 0,6 0,5 -,+ o 0,4-i
r
!
I
7
e
I
,
i
oo -
11
9
- ~ - = - I
13
~ -~
15
I
17
0,1'
'~ 1
~,
1
i
19
I
1
oeo
I
i
21
i
!
23
! ---4--------4---
25
27
I
AA 5
7
9
29
CoMo/TiO2-AI203 5%
' 3
r
~
11
13
15
17
19
21
23
25
~A 27
29
EDS spectrum number
Figure.3. EDS analyses of CoMo sulfided catalysts.
Moreover, it was noticed that over decreases after the first analysis upon beam used in these studies. On the contrary, over considerably less pronounced. That fact,
A1203 the S/(Co+Mo) atomic ratio exposure under the small probe size TiO2-A12035% the phenomenon was which has never been previously
159
reported, could be related to the energy required to form sulfur vacancies which are believed to be the actives sites for many hydrotreating catalysts. Indeed, several works have been done to try to correlate the HDS activity of sulfides with the metal-sulfur bond energy [20,21]. For promoted catalysts, Byskov et al [22] have recently reported that the presence of cobalt (or nickel) atoms at the edge of MoS2 leads to a significant lowering of the metal sulfur binding energy, and therefore to an increase in the concentration of the active sites for the reaction (promoting effect). Our DRS results and the EDS analysis appear therefore in good agreement since the lower stability of the sulfur phase observed over A1203 (leading to more uncoordinated molybdenum active sites) is associated with a high dispersion of cobalt in close contact with molybdenum in the oxidic state.
4. C O N C L U S I O N The aim of the present work was to understand the effect of the nature of the support of CoMo hydrotreating catalysts over the level of the promotion induced by cobalt. From the UV-vis data presented above, it is demonstrated than in the case of the CoMo/TiO2-A12035% catalyst cobalt is not lost into the titania matrix, but mainly involved in the formation of CoMoO4 (hardly sulfided compound). The formation of a CoMo interaction in the oxidic state, assumed to be the precursor of the active CoMoS phase is therefore reduced, in agreement with the lower synergetic effect observed. Moreover, during EDS analysis of the sulfided samples a lower stability of the CoMo/A12Oa was noted. This fact is in line with recent proposals on the effect of cobalt on the lowering of the metalsulfur bond energy, leading therefore to an increase in the concentration of active sites.
Acknowledgements The present work was carried within the framework of the program (( HDS of Gasoils )) supported by ELF, IFP, TOTAL and CNRS-ECODEV. 5. R E F E R E N C E S
1- M. Breysse, J.L.Portefaix and M. Vrinat, Catal. Today, 10(1991)489. 2- F. Luck, Bull.Soc.Chim.Belg., 100(1991)781. 3- Y.S. Ng and E. Gulari, J.Catal., 95(1985)33. 4- H. Shimada, T.Sato, Y. Yoshimura, J. Haraishi, A. and Nishijima, J. Catal. ,110(1988)275. 5- J. Ramirez, S.Fuentes, G.Diaz, M.Breysse, M.Lacroix and M. Vrinat, Appl. Catal., 52(1989)211. 6- C. Pratt, J.V. Sanders, V. Cristov, J. catal., 124(1990)416. 7- E. Olguin, M. Vrinat, L. Cedeno, J. Ramirez, M. Borque and A Lopez-Agudo, Appl. Catal., 165(1997)1.
160
8- V. Harl~, M. Breysse, J. Ramirez and M. Vrinat, Actas XIV Simposio Iberoamericano de Catalisis, Sociedad Chilena de Quimica edit., Vol. 3, 1994, p.1357. 9- J. Ramirez, L. Ruiz-Ramirez, L. Cedeno, V. Harle, M. Breysse and M. Vrinat, Appl. Catal., A93(1993)163. 10- Y. Moro-Oka, S. Tan and A. Ozaki, J. Catal., 12 (1968) 291. 11- H. Praliaud, J. Less. Common. Metals, 54 (1977) 387. 12- H. Ashley, P.C.H. Mitchell, J. Chem. Soc A, (1968) 2821. 13- P. Gajardo, P. Grange, B. Delmon, J. Catal., 63 (1980) 201. 14- C. Wivel, B.S. Clausen, R. Candia, S. Morup, H. Topsoe, J. Catal., 87 (1984) 497. 15- N. Y. Topsoe and H. Topsoe, J. Catal., 75 (1982) 354. 16-.J.A.R. van Veen, E. Gerkema, A.M. van der Kraan, P.A.J.M. Hendriks, H. Beens, J. Catal., 133 (1992) 112. 17-.X. Gao, Q. Xin, Catal. Lett., 18 (1993) 409. 18- C. P. Cheng, G.L. Schrader, J. Catal., 60 (1979) 276. 19- J. Medena, C. Van Stam, V.H.J. de Beer, A.J.A. Konings, D.C. Koningsberger, J. Catal., 53 (1978) 386. 20- P. Raybaud, G. Kresse, J. Hafner, H. Toulhoat, J. Phys., Condens., Matter, 2(1997)11085. 21- J. K. Burdett and J.T. Chung, Surf.Sci., 236(1990)L353. 22- L.N. Byskov, B. Hammer, J.K. Norskov, B.S. Clausen and H. Topsoe, Catal. Letters 47(1997)177.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
161
Effect of Chelating Agents on HDS and Aromatic Hydrogenation over CoMo- and NiW/AI20 3 Yukie Ohta, Takehiro Shimizu, Takehide Honma, and Muneyoshi Yamada Department of Applied Chemistry, Graduate School of Engineering, Tohoku University Aoba, Aramaki, Aoba-ku, Sendai 980-8579, JAPAN
Abstract Hydrotreatment catalysts (CoMo-, NiMo- and NiW/A1203) were prepared by an impregnation method with solutions containing a chelating agent (nitrilotriacetic acid (NTA), ethylenediaminetetraacetic acid (EDTA) or cyclohexanediaminetetraacetic acid (CyDTA)), and were subjected to some test reactions: hydrodesulfurization (HDS) ofbenzothiophene and dibenzothiophene, and hydrogenation (HGN) of o-xylene and 1-methylnaphthalene. Chelating agent modified CoMo- and NiW/A1203 showed higher activities in both HDS and HGN than the catalysts without the modification. The chelating agents had little effect on the activity of NiMo/A1203. CyDTA was the most effective for HDS activity of CoMo/A1203 and HGN activity of NiW/A1203. The chelating agents had no effect on the activity of each component catalysts (i.e., Co-, Ni-, Mo-, and W/A1203). The activity of the CyDTA-modified catalysts depended strongly on sulfiding temperature. CyDTA exhibited the improving effect at higher sulfiding temperatures, while working as an inhibitor at lower temperatures. Our previous study had indicated that the pre-formation of the MoS2-1ike structure was necessary to induce the intrinsic promoting effect of Co on the activity of Mo/A1203. The chelating agent, forming Co complex which decomposes at a rate depending on sulfiding temperature, was considered to adjust timing when Co ions interact with the MoS2-1ike structure.
1. I N T R O D U C T I O N From the recent environmental point of view, hydrotreatment of petroleum fractions to produce clean and high quality transportation fuels is becoming more and more important. In particular, HDS and aromatic HGN of diesel fuel are very important, and there is a growing need for improvements in catalyst performance of CoMo/A1203 and NiW/A1203, which are widely used for HDS and HGN treatments of petroleum fractions.
162
The catalysts mentioned above are conventionally prepared by impregnation followed by drying, calcination, and sulfiding for activation.
In order to obtain a high
performance catalyst, every preparation step should be optimized.
Since one of the most
important characteristics of these catalysts is the activity-promoting effect of Co or Ni, the preparation steps should be optimized to induce the intrinsic promoting effect of Co or Ni. Van Veen group reported that CoMo catalysts prepared by impregnating the supports (carbon, silica, or A1203 ) with a solution containing Mo, Co and NTA showed higher activity for HDS of thiophene at atmospheric pressure. They have assigned the higher HDS activity to the formation of a complex of NTA with Mo and Co on the supports. Being stimulated by their work (HDS of thiophene at atmospheric pressure), we have started to investigate the applicability of other chelating agents in other reactions. In the present work, impregnating solutions containing NTA, EDTA or CyDTA were used to prepare modified CoMo- and NiW/A1203. EDTA forms complexes with both Mo and Co as well as NTA does, while CyDTA only with Co.
2. E X P E R I M E N T A L Every catalyst examined here was prepared by an incipient wetness method as follows [2, 3] : -A1203 was impregnated with an aqueous solution containing a chelating agent, cobalt nitrate (or nickel nitrate) and ammonium paramolybdate (or ammonium metatungstate), then dried at 393 K in air. The molar ratio of a chelating agent to Mo (or W) was 1.2 for NTA, and 0.6 for EDTA or CyDTA. These catalysts are abbreviated hereinafter as "chelating agent"-"combination of metals", e.g., NTA-CoMo. The catalysts thus obtained were used without calcination. The catalysts were packed in a conventional fixed bed flow reactor and sulfided insitu in the stream of 5% HzS/H 2 under 1.1 MPa at 673 K. Immediately after the sulfiding,
activity tests were started by feeding the reactant into the reactor. The reaction conditions of HDS were as follows: 5 tool% benzothiophene (BT) in n-dodecane (or 2.5 mol% dibenzothiophene (DBT) in decalin), 543 (or 573) K, 5.1 MPa, LHSV 300 h -1, in H2 stream (300 ml/min), Hz/reactant feed 1,000 vol/vol.
The reaction conditions of HGN were as
follows: o-xylene (or 1-methylnaphthalene (1-MN)), 573 K (603 K), 5.1 MPa, LHSV 7 (or 75) h -1, in H2 stream (300 ml/min), HJreactant feed 1,000 vol/vol. Products were analyzed with GC (FID) and/or GC-MS. Details of the apparatus and the procedure were described in our preceding papers [4, 5].
163
3. R E S U L T S
AND DISCUSSION
3.1. H D S R e a c t i o n s In
HDS
reaction
of
BT,
ethylbenzene
(EB)
and
a
little
amount
of
dihydrobenzothiophene (DHBT) were produced. In order to compare the catalyst activities easier, the conversion level of BT was controlled below 30-40% by adjusting the catalyst loading
(a)
EDTA-CoMo NTA-CoMo CyDTA-CoMo CoMo Commercial CoMo 0
EDTA-CoMo NTA-CoMo CyDTA-CoMo CoMo Commercial CoMo
0.3 0.6 0.9 1.2 EB Yield/metal/%/gmol
EDTA-NiW NTA-NiW CyDTA-NiW NiW Commercial NiW
(b)
biph~yl ~ ]
cyclohexylbenzene,,,1 ,6 (a)
0 0.2 0.4 0.6 0.8 1 BP and CHB Yields/metal/%/gmol
EDTA-NiW NTA-NiW
(b)
CyDTA-NiW NiW 0
0.3 0.6 0.9 1.2 EB Yield/metal/%/gmol
EDTA-NiMo
0 0.2 0.4 0.6 0.8 1 BP and CHB Yields/metal/%/gmol
CyDTA-NiMo ~
NTA-NiMo CyDTA-NiMo
NiMo
NiMo 0
0.3 0.6 0.9 1.2 EB Yield/metal/%/gmol
Figure 1. Effects of chelating agents on benzothiophene HDS activity, Reaction conditions : 543 K, 5.1 MPa.
(c) I
I
I
I
0 0.2 0.4 0.6 0.8 1 BP and CHB Yields/metal/%/gmol
Figure 2. Effects of chelating agents on dibenzothiophene HDS activity. Reaction conditions : 573 K, 5.1 MPa.
in the reactor. The EB yield was used as an index of HDS activity of the catalysts. The
catalyst activities thus obtained, were compared based on the %EB yield/mol-metal.
The
activities of both CoMo and NiW were improved by the addition of the chelating agents in the following order: C y D T A > E D T A > N T A > none (Figures 1(a) and 1(b)). CyDTA-CoMo
and C y D T A - N i W
attained about 70%
and 65%
activities than the corresponding unmodified catalysts, respectively.
higher HDS
On the other hand,
164
activity of NiMo was not affected by the chelating agents (Figure 1(c)). Effects of these chelating agents on the catalyst activity were further examined in HDS of DBT. In this reaction, biphenyl (BP) was mainly produced and cyclohexylbenzene (CHB) was also produced. The sum of BP and CHB yields was regarded as HDS activity of each catalyst. The activities of CoMo and NiW were also improved by the addition of the chelating agents in the following order: CyDTA > EDTA > NTA, none (Figures 2(a) and 2(b)). The activity of CoMo was increased ca. 25% by the addition of CyDTA, but remained unchanged by NTA.
Van Veen et al. have reported negative results for HDS of DBT on
NTA-modified CoMo/A1203 [6].
Reasons for this contradiction are not yet clear.
HDS
activity for DBT was much more promoted in NiW catalyst than in CoMo catalyst by the addition of the chelating agent. HDS activity of the NiW catalyst might be more susceptible to the preparation method than that of the CoMo catalyst.
3.2. HGN Reactions For HDS of DBT, it has often been said that HGN activity of the catalyst is important. So we have examined HGN activities of the catalysts prepared with the chelating agents
in HGN
of o-xylene as a test reaction.
Reaction products
were
1,2-
dimethylcyclohexane, some dimethylcyclohexanes, and m-, p-xylene. HGN activity of NiW for o-xylene (the sum of yields of 1,2-dimethylcyclohexane and other cycloaliphatic compounds) was improved three-fold by the addition of CyDTA (Figure 3). 5-methyltetralin 1-methyltetralin EDTA-NiW
CyDTA-CoMo
NTA-NiW
CoMo
CyDTA-NiW
CyDTA-NiW
NiW
NiW
0
0
5 10 15 20 HGN Yield/metal/%/mmol
Figure 3. HGN activity of CyDTA-modified CoMo- and NiW/A1203 for o-xylene. Reaction conditions 9573 K, 5.1 MPa.
0.1
0.2
0.3
0.4
0.5
HGN Yield/metal/%/gmol Figure 4. HGN activities of chelating agentmodified NiW/A1203 for 1-methylnaphthalene. Reaction conditions" 603 K, 5.1 MPa.
1-MN was also hydrogenated to examine the effect of the chelating agents on the HGN activity of the catalysts. 1-methyltetralin and 5-methyltetralin were mainly produced. The sum of the yields of 1-methyltetralin and 5-methyltetralin was used as an index of HGN activity. HGN activity of NiW was improved about 40% by the addition of CyDTA (Figure
4). The activity of each component system, i.e., Co-, Mo-, Ni- or W/A1203 was not
165
improved by the addition of the chelating agents. The chelating agents are considered to improve synergy between Co and Mo or between Ni and W, leading to induce the intrinsic promoting effects of Co for CoMo and Ni for NiW. The chelating agents may have a role to improve the formation of specific active phase.
3.3. Complex Formation Constants An important development has recently been made in research of the Co-Mo-S structure by van Veen et al. [ 1]. In the study expecting to understand "a real support effect", they found that Co-Mo-S phase was selectively formed by using NTA in impregnating solution by means of M6ssbauer emission spectroscopy. The preparation method, originally invented by a researcher of Shell for S i O 2 supported hydrotreatment catalysts [7], has been applied to the study using the extended X-ray absorption fine structure (EXAFS) measurement, and it was suggested that Co was located at the edge site o f M o S 2 structure and was coordinated with five or six sulfur atoms [8, 9].
Table 1 Complex formation constants of literature 10) Co
Ni
Mo
W
EDTA
16.31
18.62
18.76
19.67
NTA CyDTA
10.38 18.92
11.54 19.40
18.60 -*
19.03 -*
* We confirmed with NMR that no complexes were formed.
Van Veen reported that the improving effect of NTA on the activity of C o M o / S i O 2 (or CoMo/active-C) is due to its ability to form complex with Mo and Co at the same time. In the present work, we have examined van Veen's proposition under different conditions. Table 1 shows literature values of complex formation constant of the chelating agents with the related metal ions.
NTA forms complexes with all the metal ions.
CyDTA, the most
effective chelating agent, however, forms complexes with Co or Ni ions, but not with Mo or W ions. Considering CyDTA was more effective than NTA, the ability to make a complex of chelating agent and promoter (e.g., Co, Ni) is rather important to improve the synergy between promoter and Mo or W ions.
3.4. Sulfiding Temperature Dependence In order to investigate the mechanism in which chelating agents improve synergy between Mo and Co or between W and Ni, effects of sulfiding temperature on the catalyst
166
activity were examined. Figure 5(a) shows the sulfiding temperature dependence of HDS activities of CyDTA-CoMo and CoMo. Figure 5(b) shows the dependence of HGN activities of CyDTANiW and NiW. In these Figures, the following two points are noticed with respect to the improving effect of CyDTA. First, the activity of the CyDTA-modified catalysts depends more strongly on sulfiding temperature than that of the unmodified catalysts. Secondly, the order of the catalytic activities of the CyDTA-modified and unmodified catalysts are inverted at a lower sulfiding temperature.
At higher sulfiding temperatures, the CyDTA modified
catalysts show higher activities in HDS and HGN reactions than the unmodified catalysts. At a lower sulfiding temperature, however, the activities of CyDTA-modified catalysts are lower than those of the unmodified catalysts. That is, CyDTA exhibits the improving effect at higher sulfiding temperatures, while working as an inhibitor at lower temperatures. 1.2
(a)
9
CyDTA-CoMo O
~0.8 .,.a
~0.6 E 0.4
.,..~
~0.2
_$
0 450
0
~10 -
CoMo
N5 -i~
o
o
-8
I
(b) CyDTA-NiWO
-~15 -
9
m
~20
I
550 650 750 Sulfiding temperature/K
Z
:=0
Q 450
I
0 NiW I
550 650 750 Sulfiding temperature/K
Figure 5. The sulfiding temperature dependence of catalytic activities of modified and unmodified catalysts. (a) Benzothiophene HDS Reaction : 543 K, 5.1 MPa. (b) o-Xylene HGN Reaction : 573 K, 5.1 MPa. 3.5. R o l e of C h e l a t i n g A g e n t s The sulfiding temperature dependency of the effect of CyDTA is considered to be caused by the strong interaction between CyDTA and Co (or Ni) ion as shown in Table 1. At relatively lower sulfiding temperatures, CyDTA interacts strongly with Co (or Ni) ion, resulting in inhibiting Co (or Ni) ion from interacting with Mo (or W) or A1203. At higher sulfiding temperatures, however, the complex between CyDTA and Co (or Ni) ion decomposes, resulting in the interaction between Co and Mo (or Ni and W). Our concept is depicted in Figure 6. In the preceding paper [11], we reported the effects of various pretreatments (including sulfiding and reducing) on the activity and structure of CoMo/A1203. In the report, the intrinsic high activity of CoMo was found to be induced by proper sulfiding pretreatment. From the results of activity test and Mo K-edge EXAFS analysis, it was concluded that the
167
appearance of the intrinsic promoting effect of Co was closely connected with the formation of MoSz-like structure by proper sulfiding pretreatment. The pre-formation of the MoSz-like structure was necessary to induce the intrinsic promoting effect of Co on the activity of Mo/AI203. That is, the intrinsic promoting effect of Co was induced on the surface of the MoS2-1ike structure.
Figure 6. Scheme of the fornation of active sites in CyDTA-CoMo/A1203.
The present results support our previous proposition. As shown in Figure 6, without chelating agents, Co ions can react with A1203 or be sulfided to form less active COA1204 or Co9S8, respectively. Co ions also interact freely with Mo to interfere with the formation of the MoSz-like structure. In the presence of a chelating agent such as CyDTA, Co ions are so strongly complexed with the chelating agent that sulfiding of the Co ions or interaction of the Co ions with Mo or A1203 will be inhibited, while Mo ions are sulfided to form the MoSz-like structure. At lower sulfiding temperatures, the chelated Co ions remain undecomposed. Accordingly, highly active sites resulted from interaction between Co and the MoS~-like
168
structure will not be formed. A chelating agent works as an inhibitor. At higher sulfiding temperatures, the chelated Co ions decompose with time. In other words, the chelated Co ions decompose after Mo ions are sulfided to some extent to form the MoSz-like structure.
The Co ions thus formed, eventually interacts with the MoS2-1ike
structure to form highly active sites. Thus, the role of chelating agents is considered to adjust timing when Co ions interact with the MoS2-1ike structure, leading to induce the intrinsic synergy between Mo and Co. 4. A C K N O W L E D G E M E N T A part of this work has been carried out as a research project of the Japan Petroleum Institute commissioned by the Petroleum Energy Center with the subsidy of the Ministry of International Trade and Industry.
5. R E F E R E N C E S J. A. R. van Veen, E. Gerkema, A. M. van der Kraan, and A. Knoester, J. Chem. Soc.,Chem. Commun., (1987) 1684. T. Shimizu, S. Kasahara, T. Kiyohara, K. Kawahara, and M. Yamada, Sekiyu Gakkaishi, 38 (1995) 384. K. Hiroshima, T. Mochizuki, T. Honma, T. Shimizu, and M. Yamada, Appl. Surf. Sci., 121/122 (1997) 433. M. Yamada, A. Saito, T. Wakatsuki, T. Obara, J.-W. Yan, and A. Amano, Sekiyu Gakkaishi, 30 (1987) 412. M. Yamada, Y.-L. Shi, T. Obara, and K. Sakaguchi, Sekiyu Gakkaishi, 33 (1990) 227. J. A. R. van Veen, H. A. Colijn, P. A. J. M. Hendriks, and A.J. van Welsenes, Fuel Processing Technology, 35 (1993) 137. 7
M. S. Thompson, Eur. Pat. Appl., EP 181035 (1986).
8
S. M. A. M. Bouwens, J. A. R. van Veen, D. C. Koningsberger, V. H. J. de Beer, and
R.
Prins, J. Phys. Chem., 95 (1991) 123.
9
S. M. A. M. Bouwens, F. B. M. van Zon, M. P. van Dijk, A. M. van der Kraan, V. H. J. de Beer, J. A. R. van Veen, and D. C. Koningsberger, J. Catal., 146 (1994) 375
10
L.G.Silen and A.E.Martell, Stability constants of metal-complexes vol.2, Chemical Society, London, 1964.
11
S. Kasahara, Y. Udagawa, and M. Yamada, Appl.Catal.B. Environmental, 12, (1997) 225.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
169
G e n e s i s , S t r u c t u r a l , a n d C a t a l y t i c P r o p e r t i e s of N i - M o - P - A l u m i n a b a s e d H y d r o t r e a t i n g C a t a l y s t s P r e p a r e d by a Sol-Gel M e t h o d Ryuichiro Iwamoto a and J e a n Grimblot b
a Petroleum Refining Technology Center, Idemitsu Kosan Co., Ltd., 1280 Kami-izumi, Sodegaura, Chiba, 299-0293, Japan. b Laboratoire de catalyse h6t6rog~ne et homog6ne, URA CNRS 402, Universit6 des Sciences et Technologies de Lille, 59655 Villeneuve d'Ascq C6dex, France.
Abstract Ni oxide - Mo oxide - P oxide - Alumina with wide range of P loading (from 1 to 10 wt% P) were prepared by a sol-gel method to elucidate the role of P on the genesis, structural, and catalytic properties of Ni-Mo based hydrotreating catalysts. Specific surface area of catalysts decreased gradually in proportion to the P loading from -500 to 260 m2/g. X-ray powder diffraction revealed that both small and large amounts of P within the alumina framework provoke aggregation of Mo related species. 27A1-NMR indicated that a part of octahedral aluminium sites is highly distorted in the presence of Ni, Mo and P. From 31P-NMR measurements, predominant formation of polymeric P oxospecies and AIPO4 was observed after the calcination step. Thiophene HDS activity gave a maximum at 2 wt% of P due to an increase in Mo dispersion. However, large amount of P has a negative effect on HDS activity due to the formation of bulk MOO3. Butane formation during thiophene HDS decreases with P addition which may indicate the segregation of Ni from MoS2. 1. INTRODUCTION The effects of P introduced in some formulations of hydrotreating catalysts are still matter of debate as extensively discussed in a recent review [1]. Indeed, P may be incorporated in the alumina framework to form
170
amorphous or crystalline aluminium phosphate(s) or it may also induce formation of undesirable phase like MoO3 badly dispersed on the alumina support. P also modifies the textural characteristic and acidity of the catalysts. When present in moderate loadings, it may have a beneficial influence on some hydrotreating reactions. In the previous work on the unpromoted Mo-PAlumina catalysts prepared by a sol-gel method [2], we have shown that the thiophene HDS activity was not promoted by P loading while large amounts of P decreases the catalytic performances, probably by formation of bulk MoO3 in the oxide precursor. Such Mo-P-Alumina catalysts have been also largely described by solid state NMR [3]. Their acidity and hydrogenation properties have also been studied [4]. The role of P on a promoted Mo-Alumina based hydrotreating catalysts is also quite interesting as it is the system for conventional commercial hydrotreating catalysts. Muralidhar et al. [5] reported that thiophene HDS activity over Co-Mo-P-Alumina does not change at 0.5 wt% P loading while it decreases at 5 wt% P loading. Eijsbouts et al. [6,7] concluded that thiophene HDS over Ni-Mo Alumina is not considerably promoted by P addition. On the other hand, Atanasova and Halachev [8] found that P gives maximum activity for thiophene HDS a t - 2 wt% P905 over Ni-Mo-P-Alumina. Walendziewski [9] also observed a small maximum for thiophene HDS over a Co-Mo-P-Alumina at 1.3 wt% P. Chadwick et al. [10] also reported that thiophene HDS over Ni-MoAlumina shows a broad maximum at -1 wt% P. Lewis et al. [11] observed positive effects of P on gas oil HDS at -1 wt% P over Ni-Mo-P-Alumina. Kemp et al. [12] reported that Ni-Mo-P-Alumina and Co-Mo-P-Alumina prepared by hydrogel method shows the higher HDS activity for cracked heavy gas oil than commercial catalyst. Jones et al. [ 13] found that gas oil HDS activity increases up to 3 wt% P. Chen et al. [14] also reported that HDS of atmospheric residue over Co-Mo-P-Alumina shows maximum activity a t - 5 wt% P. Therefore, the effect of P on the HDS activity over promoted Mo-P-Alumina has been not well understood yet. In this work, the influence of P on the genesis and structural properties of Ni-Mo-P-Alumina catalysts prepared by a sol-gel method was investigated. The obtained catalysts have been also tested in thiophene HDS.
2. EXPERIMENTAL
The Ni-Mo-P-Alumina catalysts were prepared by a sol-gel method [2]. The oxide precursor was obtained by the hydrolysis of A1 sec-butylate in the presence of 99% of HsPO4, (NH4)6MoTO24, and Ni(NOs)2-6H20. Ni/(Ni+Mo) atomic ratio was adjusted to 0.33 since it is considered to give the highest promoting effect on conventional HDS catalysts [15]. The solid obtained was
171
dried at 100 ~ and calcined at 500 ~ for 3 h. The catalysts thus prepared are noted as MPD(Y-Z), NPC(X-Z), NMPD(X-Y-Z), NMPC(X-Y-Z) where MP, NP or NMP mean Mo-P, Ni-P, or Ni-Mo-P-Alumina catalysts ; D or C means dried or calcined samples; X, Y, or Z means the expected loadings in wt% of the elements Ni, Mo, or P, respectively. The chemical compositions were provided by the "Service Central d' analyses du CNRS" (Vernaison, France). The calcined samples were characterized by BET specific surface area (QUANTASORB Jr., Quantachrome) after preheating at 200 ~ for 30 min. X-ray diffraction (Siemens D5000 Diffractometer equipped with a goniometer, a monochrometor, and a Cu X-ray tube), 27A1-NMR (ASX400 BRUKER ; resonance frequency 104.26 MHz, recycling time 3 sec., pulse time 1 ~sec., spinning frequency 15 kHz, and Al(H20)6 S§ as a reference) and 31p-NMR (ASX100 BRUKER; resonance frequency 40.53 MHz, recychng time 40 sec., pulse time 2 psec., spinning frequency 7 kHz, and H~PO4 as a reference) were also used to characterize the catalysts. Thiophene HDS was carried out at 300 ~ with the procedure already described [2]. 3. RESULTS AND DISCUSSION 3.1. Chemical composition and specific surface area (SSA) Table 1 shows chemical composition and SSA of calcined samples. The obtained amount ofNi, Mo, and P is close to that expected from the preparation procedure. The amount of carbon remaining in the calcined samples is less than 0.5 wt% for all the catalysts. Table 1 Composition and SSA of calcined Ni-M0-P-Alumina sol-gel catalysts. Catalysts (~) Mo Ni P Carbon SSA (wt%) (wt%) (wt%) (wt%) (m2/g) MPC(0-0) 0 0 0 0.5 503 MPC(0-11) 0 0 10.6 0.3 474 NMPC(6-20-0) 20.8 6.5 0 0.3 609 NMPC(6-20-1) 19.3 7.2 1.3 0.5 526 NMPC(6-20-2) 21.6 6.8 2.4 0.4 461 NMPC(6-20-4) 20.4 5.8 3.8 0.3 451 NMPC(6-20-6) 19.5 6.3 6.3 0.5 400 NMPC(6-20-10) 20.3 5.8 9.7 0.3 263 (a)X, Y, and Z in MPC(Y-Z) and NMPC(X-Y-Z) mean the expected loading of Ni, Mo and P, respectively.
172
SSA decreases in proportion to the P content, although it remains as high as 400 m2/g even with high metal loadings and high P content [sample NMPC(6-20-6)]. 3.2. X-ray powder diffraction (XRD) Figure 1 shows the XRD patterns of the Ni-Mo-P-Alumina catalysts. In the P-free NMPC(6-20-0) catalyst, the formation of bulk NiMoO4 is identified (Fig. 2a). The addition of 2 to 4 wt% P into catalyst formulation prevents the formation of NiMoO4 as its relevant diffraction peaks are hardly detected (Fig. 2b and 2c). However, bulk MoO3 appears again above 6 wt% P (Fig. 2d and 2e).
9 MoO3
e) d}
N~
0
i
i
i
20
40
60
2 0
80
/ degree
Figure 1. XRD patterns of Ni-Mo-P-Alumina catalysts, a)NMPC(6-20-0), b)NMPC(6-20-2), c)NMPC(6-20-4), d)NMPC(6-20-6), and e)NMPC(6-20-10) This result indicates that a moderate amount of P increases the dispersion of Mo but smaller or larger amounts of P are detrimental for optimal Mo dispersion. The presence of P may help to increase the stability of Ni-Mo complexes which could be formed in the preparation solution, though larger amount of P then impedes the interaction between the Mo oxo-species and alumina as already seen in the Mo-P-Alumina system [2]. Since no bulk NiO
173
n o r NiA1204 is observed, Ni may be predominantly associated or in close interaction with Mo species rather than with the alumina surface.
3.3. Solid state 27A1-NMR
Table 2 shows top peak value of 27A1-NMR spectra of the dried and calcined Ni-Mo-P-Alumina catalysts. In the dried state, bare alumina [MPD (0-0)] shows a single broad signal assigned to octahedral aluminium site (Alocta) at ~ 7 ppm. With addition of Ni and Mo [NMPD(6-20-0)], a peak of weak intensity attributed to tetrahedral aluminium sites (Alte~a) appears at ~60 ppm. With addition of larger amounts ofP [NMPD(6-20-10)], the formation of A1PO4 (more generally, Altetra-O-P sites) are also observed at 40 ppm. In addition, broadening of Alocta peak toward lower frequency suggests the formation of Alocta-O-P sites as considered from previous 2D 27A1-MQMAS NMR studies [3]. In the calcined bare alumina [MPC(0-0)], signals attributed to Aloct~, Altetra, and 5-fold coordinated aluminium sites are observed at ~ 7, 66, and 30 ppm, respectively. P-free Ni-Mo-Alumina catalyst [NMPC(6-20-0)] shows a spectrum similar to those of the bare alumina and the signal for Ni aluminate is not pronouncedly observed at ~25 ppm in contrast with Mo-free Ni-Alumina catalyst [NPC(6-0)]. This indicates again that Ni species are preferably associated with Mo species rather than with the alumina in the Ni-Mo-Alumina catalyst. This interaction between Ni and Mo in oxide form may transform into a commonly accepted Ni-Mo-S phase after sulfidation [16-18]. The intensity of 5-fold coordinated alumina at 30 ppm increases significantly in NMPC(6-20-4). Furthermore, the top peak position of Alocta at 5 ppm shifts toward lower values and tends to be broader with P and Mo loading. This suggests that distortion of A1 becomes more pronounced in the presence of P, Ni, and Mo. Table 2 Results of 27A1NMR of Ni-Mo-P-Alumina based sol-gel catalysts. Catalysts MP(0-0) NP(6-0) MP(20-0) NMP(6-20-0) NMP(6-20-4) NMP(6-20-10)
Before calcination (ppm) 58s 60~ 61~ 40m
71 71 71 61 51 21
After calcination (ppm) 66m 64m 62m 62m 62m
33s
71
251 30~ 30m 301 361
61 51 61 21 -51
s,m,1 refer to the intensity of spectra, s:small, m:medium, and l:large.
174
In NMPC(6-20-10), the characteristic spectra for A1PO4 are observed at -36 ppm. In this catalyst, the formation of A12(MoO4)3 is not observed while bulk MoO3 is detected by XRD. This is not the same as the case of Mo-P-Alumina catalysts [2,3]. As Ni seems to be preferably associated with Mo species, it makes weaker the interaction between alumina and the Mo-oxo species to conduct after calcination to aluminium molybdate. 3.4. Solid state 31P-NMR Table 3 shows top peak value of 31P-NMR spectra of the dried and calcined Ni-Mo-P-Alumina catalysts. In dried samples, P-Alumina catalyst [MPD(0-11)] gives monomeric and polymeric P oxo-species at -10 and-21 ppm, respectively. On the other hand, the Ni, Mo, and P containing catalysts such as NMPD(6-20-4) and NMPD(6-20-10) give another characteristic signal at -15 ppm as well as that of Mo-P-Alumina catalyst [2,3]. This signal could be assigned to the formation of less polymerized P oxo-species or to the formation of some Ni-Mo-P mixed oxo-species. In calcined samples, the effect of P addition on the 3~P-NMR spectra is not so pronounced. All the calcined catalysts show polymeric P oxo-species and A1PO4 at about -18 and -25 ppm, respectively.
Table 3 Result of s1p NMR of Ni-Mo-P-Alumina based sol-gel catalysts. Catalysts MP(0-11) NMP(6-20-4) NMP(6-20-10)
Before calcination (ppm) - 10 - 10 -10
- 15 -15
-21 -21 -21
After calcination (ppm) - 18 - 18 -18
-25 -25 -25
3.5. Thiophene HDS activity Figure 2 shows the thiophene HDS evolution of the Ni-Mo-P-Alumina catalysts as a function of the P content. For a comparison basis, the results of Mo-P-Alumina [MPC(20-7)] and Ni-P-Alumina [NPC(18-9)] measured at same reaction conditions are also indicated. Concerning the trend for the Ni-Mo-PAlumina catalysts, HDS activity increases up to around 2 wt% P and then smoothly decreases with further addition of P. The initial HDS activity increase up to 2 wt% P is probably due to a better Mo dispersion as revealed by XRD measurements (Fig. 1). In the same way, the activity decrease at higher P loading could be correlated to the formation of bulk MOO3. The HDS conversion level is considerably higher than the sum of the activity of Mo-P-Alumina and Ni-P-Alumina catalysts. Clearly, a large part of Mo in the sulfided state is
175
promoted by the Ni species like in the classical Ni-Mo catalysts. In this way, Fig. 2 gives a clear indication of the promotor hierarchy for thiophene HDS : Ni is the effective promotor of the active MoS2 phase while the further addition of 2 to 4 wt% of P makes the catalysts still more efficient. The combination effects could be due to the contribution of P to stabilize the Ni-Mo complexes during preparation procedure. The Ni-Mo-Alumina catalyst has higher hydrogenation selectivity for saturated C4 production (butane) (26 %) than that of Mo-Alumina catalysts (20 %). This means that the presence of Ni also promotes HYD activity as in the case of thiophene conversion. However, the hydrogenation selectivity decreases quite linearly with the P content from 26 to 15 %. This suggests that a part of Ni species dissociates from MoS2 and Mo is not totally promoted in the presence of P. If this hypothesis is correct, higher Ni/Mo ratio could more improve the HDS activity in the presence of P. 50 45
Ni-Mo-P-AI
40 ---
35
e c o
25
o0 a-r" r
.2 r l--
30
20
15 10
Mo-P-AI
. . . .
0
I
2
i
i
i
I
I
4
I
i
i
i
I
6
I
9 I
I
Ni-P-AI I
J
8
I
A I
I
I
J
10
I
'
'
'
12
P content (wt%)
Figure 2. Thiophene HDS activity of Ni-Mo-P-Alumina catalysts. Mo-PAlumina and Ni-P-Alumina are also indicated for comparison. 4. CONCLUSION
Ni-Mo-P-Alumina catalysts with a wide range of P loading were prepared by a sol-gel method to elucidate the role of P on the textural, structural, and catalytic properties of Ni-Mo based hydrotreating catalysts. The
176
amount of P affects significantly on the physicochemical and catalytic properties of catalysts. P decreases the specific surface area while a moderate amount of P increases the Mo dispersion due to the stabilizing Ni-Mo complex which could be formed during the preparation procedure. The HDS activity is considerably promoted by Ni and P addition up to 2 wt% due to Mo dispersion increase. However, segregation of Ni from MoS2 may occur by the addition of P considering the hydrogenation selectivity decrease. 4. REFERENCES 1 2 3 4 5 6
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I
INDUSTRIAL PROCESS ASPECTS
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
179
Hydroprocessing Kinetics for Oil Fractions Teh C. Ho Corporate Research Laboratories Exxon Research and Engineering Company Annandale, NJ 08801, USA
Abstract The kinetics of hydroprocessing reactions (HDS, HDN, HDA) for complex mixtures such as petroleum distillates exhibit many "peculiarities." Some examples: (1) the overall HDS or HDN reaction order for oil fractions is higher than that for individual organosulfur or organonitrogen species; (2) high-activity catalysts show lower overall order than low-activity catalysts; (3) tough feeds show higher overall order than easy feeds; (4) the overall HDS order decreases with increasing temperature; and (5) the overall order depends on reactor type (e.g., plug flow vs. stirred tank reactor). This paper discusses a theoretical means to explain these observations.
1. INTRODUCTION Oil fractions by their nature contain an astronomically large number of organosulfur, organonitrogen and aromatic species. The concentration and reactivity of these species vary widely from feedstock to feedstock. In practice, it is the aggregate behavior, not the individual behavior, of these species that matters. For instance, in hydrodesulfurization (HDS), one cares only about the reduction of total sulfur, not of the individual sulfur-bearing species. The purpose of this survey paper is to briefly discuss a theory of kinetic aggregation, which predicts the "peculiarities" described in the Abstract. The analysis is based on the approximation that the number of reacting species in an oil fraction is so large that the mixture can be treated as a continuum. The idea of continuous mixtures is not new [1]; petroleum properties are often measured as a continuous function of boiling point. The problem addressed here can be broadly stated as follows. Let c(k,t) be the concentration of the individual reactant (e.g., sulfur species) with rate constant k and C(t) be the total concentration of all reactants at time t. Each reactant disappears at the rate r(c). The aim is to predict the dependence of C(t) on oil properties and reactor type. It is also of interest to know if an overall aggregated kinetics R(C) can be found for the mixture as a whole. To do so, it is necessary to have complete information on the oil composition and reactivity spectra a priori; that is, to be working with a fully characterized mixture. Unfortunately, oil fractions can seldom be fully characterized. The question then is whether one can make some general statements about C(t) and/or R(C) with minimum information, that is, for a mixture that has only been partially characterized. Fortunately, the answer to this question is affirmative if we restrict ourselves to the high-conversion (or near equilibrium) regime. This is comforting, since in practice it is the high-conversion regime that matters; in HDS one wishes to remove all of the sulfur compounds.
180
We will show how R(C) can be obtained for some specific cases. This is followed by the development of the high-conversion asymptotic kinetics Ro(C) for partially characterized feeds. Specific topics include pore diffusion effects, reactor mixing, and reaction reversibility. We also show how Ro can be used to estimate C(t) for all t. Finally, the limitation of the continuum theory is discussed.
2. AGGREGATED KINETICS Experiments using model compounds have indicated that the individual HDS rate is pseudo-f'trst order [2]; that is, r(c) = kc. In a plug-flow reactor (PFR), c(k,t) = c~(k)exp(-kt), and C(0 can be approximated by
C(t) = ; c f (k)exp(-kt)dk
c(0) = Q
(1)
Here c~k)dk is the concentration of species with rate constant between k and k + dk. For practical purposes, c~k), the distribution of the constituent species in the feed, can be assumed to approach zero rapidly as k - , oo. To portray a wide variety of feed properties with only two parameters, we use the gamma distribution (hereafter called gamma feeds)
c~k) = T~(k/K)~qexp(-Tk/K:)/F(T)
T> 0
(2)
where F is the gamma function. When T = 1, c](k) is an exponential distribution, a feed that comprises predominantly very unreactive species. When T > l, cj(k) is monomodal. A large T means not only an easily treatable but also a relatively homogeneous (a sharper peak) feed. A linear combination of two gamma distributions gives a bimodal distribution. From Eqs.(1) and Eq.(2) one gets C(t) = (1 + ~/T)~, implying that R(C) for the PFR (or equivalent batch reactor) is power law with an overall order n higher than one [ 1], i.e.
dC/dt= R(C)=-KC"
n=l+l/?'>
1
(3)
This says that as time progresses the reactive sulfur species desulfurize rapidly and the mixture becomes progressively more refractory, thus giving rise to a higher overall order. A tough feed with T = 1 gives an overall order of two, while an easy feed with T = 2 gives n = 1.5. This result is consistent with experimental observations [3, 4]. It has been observed that n decreases with increasing temperature [4]. Our interpretation is that increasing temperature would make the feed more reactive, hence a lower n. The overall order n can also be viewed as reflecting the activity of the catalyst for attacking the refractory species [5]: the higher the activity, the lower the order. Experimentally, this is indeed the case [6]. It should be pointed out with gamma feeds, R(C) can be obtained for certain non-firstorder reactions [7-10] in a PFR. When r is of power law, no closed form R(C) can be found. The problem of finding C(t) for nonlinear kinetics is more involved [7, 9, 11 ]. We will not dwell on this subject here, except to mention that zero-order reactions collectively can give rise to an overall order of higher than unity with tough feeds [12]. Hydrodenitrogenation of individual nitrogen species generally is zeroth or fractional order due to strong adsorption [ 13, 14]. For continuous stirred tank reactors (CSTR), R(C) cannot be found for gamma feeds. In what follows we focus on the high-conversion asymptotic kinetics Ro(C) for any arbitrary feed.
181
3. ASYMPTOTIC KINETICS Intuitively, C should be governed by the most refractory species at high conversions (large t). We thus look into the behavior of c/for small k by expanding it into a series near k = 0 cj(k) - k"t'l(c0 + Clk + . . . . .
),
]r > 0
(4)
Consistent with Eq.(2), here a small ,/signifies a tough feed (e.g., ? < 1). 3.1. Plug flow reactor
With Eqs.(1) and (4), one obtains the leading order term C ~ F(?)Co/# at large t. For the PFR, Ro(C) is governed by the most refractory species and has an asymptotic order n' [5, 15]
dC/dt ~ Ro(C)=- ~[r(V)Co]'C'
n ' = 1 + 1/v
(5)
Thus, n' > 1, and the tougher the feeds, the higher the order. Note that here r is first order. For more complicated kinetics (e.g., power law, Langmuir-Hinshelwood, bimolecular reactions), Ro in most cases is also of the power law form [8-12]. Figure 1 shows the asymptotic power law kinetics for deep HDS of a light catalytic cycle oil over three different catalysts [ 16].
-1.6
-2.4 in C -2.$
-3.2
-3.6 -0.6
-0.2
0.2
0.6
in (I/LHSV)
Figure 1" lnC vs. In(1/LHSV) for Different Catalysts 3.2. Continuous stirred tank reactor
Here c(k,t) = cj(k)/(1 + kt) for first-order reactions in a CSTR, and C(t) is of the form
C(t) = ~ c f (k)dk l+kt
(6)
182
At large t, C ~ p / t ~ with p = F(y)F(1-y)Co and u = T when Y < 1, and p = I and u = 1 when T > 1 [17]. The constant I is given by I = ~ c / ( k ) k d k , indicating that the long-time behavior of easy feeds (y > 1) is governed by all species - not just by the refractory species. When y > 1, one deals with feeds comprising primarily very reactive species. For each reactant type, there is an exponential distribution of residence times among all the molecules of that reactant. Even after a long time, C is affected by reactive species because there are lots of such molecules whose residence times can be very long. Table 1 shows the concentration dependencies of Ra(C). The order for the PFR is always higher than that for the CSTR. With the same feed information, in some cases, the performance of the CSTR can be predicted from the PFR data, and vice versa. Table 1. Asymptotic Kinetics for PFR and CSTR
PFR
R a oc
CSTR
7>0
T> 1
T
T= 1
C 1+ 1/T
C
C1/7
C/[InC]
Figure 2 shows the asymptotic order n' as a function of y for both reactors. Although no reaction order can be found for the CSTR when T = 1 (denoted by the open circle), we define an instantaneous order n * ( t ) = dlnRfldlnC, which is a weak decreasing function of conversion, consistent with experiment [18]. At high conversions, n* = 1 + 1/llnC] and therefore is higher than unity. For hydrodemetallization of the feed c / = exp(-k) in a CSTR [ 19], C(t) is described by a 1.5 order kinetics. The same feed gave a second order kinetics when run in a PFR. 6
5
4
n' 3
2
1
0
0
2
4
y.
6
8
Figure 2: Order of Asymptotic Kinetics vs. y ("----" CSTR, no order can be found when y =1, denoted by o)
183
It bears emphasizing that the asymptotic kinetics are both feed and reactor dependent; that is, they are system properties. For the CSTR, 3' plays an especially pivotal role. This is due to the interplay of the wide spreads of reactor residence times, reactant reactivity, and concentration. In a CSTIL the residence times of the molecules of each reactant type are distributed exponentially. While all reactants are slowed down compared to those in a PFIL the fast-reacting ones are hampered more than the slow-reacting ones. As a result, the mixture will appear more homogeneous. An already homogeneous feed [e.g., Eq.(2) for high 3'] will become even more homogeneous. In fact, when 3' > 1, the feed in the CSTR is homogenized to such an extent that the mixture's long-time behavior is qualitatively similar to that of a single reactant. More information on the effect of reactor mixing pattern can be found in [ 17, 20-22]. 3.3. Diffusion Effects When a single species undergoes an nth-order reaction, the presence of severe pore diffusion limitation would shift the order from n to (n + 1)/2. Obviously, the order will not change when n = 1. One may wonder if the overall order of a continuum of first-order reactions would remain intact if all reactions are severely diffusion limited. The question has been addressed [6, 17] by considering the limiting case of severe diffimion limitation, that is, the catalyst effectiveness factor approaches the reciprocal of the Thiele modulus. Denoting nd as the asymptotic order for the diffusion-limited case, one can show that the relationship between na and n' (asymptotic order for kinetics-controlled case) is as follows [ 17]: PFR
CSTR
3' > 0,
na = ( n ' + 1)/2 = 1 + 1/(23')
(7)
3' > 1 ,
na = n ' = 1
(8)
1/2 < 3' < 1,
nd = 3'n' = 1
(9)
0 < 3' < 1/2,
na = n '/2 = 1/(23')
(1 O)
Thus, as far as the apparent order is concerned, the oil in the PFR at high conversions under strong diffusion limitation can be treated as a single reactant. The same cannot be said of the CSTR, however. For the CSTR, diffusion intrusion lowers the asymptotic order only for tough feeds (3' < 1). For easy feeds (3' > 1), the asymptotic order is unity, with or without severe diffusion limitation. This behavior is similar to that of a single first-order reaction. Note that no reaction order can be found for 3' - 1/2 or 3' = 1. The above results can be rationalized as follows. In the diffimion-hindered regime, the disparities among the species become smaller because the fast-reacting ones are hindered more than the slow-reacting ones. Consequently, the mixture becomes more homogeneous. This is why the region of "single-reactant" behavior (i.e., no shift in reaction order) is widened. For further results on the diffusion effects, the reader is referred to [23-25].
3.4. Reversible Reactions Consider the single reactant case r = kc - k ' c ' with an equilibrium constant K
c(Lt) ~ c'(Lt)
(11)
184
where c'j(k) = c'(k,O). Note that r can be replaced by an equivalent irreversible kinetics r(c) = k*(c - G) where k* = k (1+ l/K) is an effective rate constant and G the equilibrium concentration. We will see that this type of equivalence can be carried over to the mixture case. As with cj(k), we let c / ( k ) ~ k~"~(Co'+ c~ 'k + .... ) for small k. The distribution of K is characterized by K(k). For a homologous series of reactions, k and K can often be related by the Polyanyi equation (or the linear free-energy relationship) K(k) = rl k~
(12)
Here both PFR and CSTR exhibit a much wider variety of asymptotic behaviors than in the irreversible case [ 17]. Specifically, the PFR admits 13 possibilities, vs. CSTR's 21 possibilities. Despite this, Ro for both the PFR and CSTR follow power law in most cases. However, there is a fundamental difference between the two reactors. When g > 1, the oil's near-equilibrium behavior in the PFR is governed by species of intermediate reactivities, whereas that in the CSTR is governed by either the most refractory species or all species. For hydrogenation of mononuclear aromatics on metal sulfides, available data indicate that k and K change in the oppose directions; that is, g < 0. In what follows we consider this special case. It can be shown that near equilibrium [ 17], Ro(C) oc ( C - Ce)"
(13)
where C, Ce and n' are the total concentration of c(k,t), total equilibrium concentration and asymptotic order, respectively. Table 2 lists n' for both PFR and CSTR with different feeds. Table 2. Asymptotic Order for Reversible Reactions (g < 0) Reactor
Feed
n'
PFR
Y
7'-g
1+1/7 1 + 1/(T'- g)
CSTR
7_<7'-g< 1 7 > 7' - g, 7' - g < 1 7> 1,7'> 1 +g
1/7 1/(7' - g) 1
4. ONE-PARAMETER MODEL The asymptotic kinetics can allow one to estimate C(t) for all t from a one-parameter model. When t -) 0, C ~ C/. At large t, C follows a power law in most cases for both the PFR and CSTR. Let C ~ o/t ~ be the general form of such power law. Both z and c can be determined from the most refractory fraction of the feed. One can then obtain an approximate expression for C(t) by combining the large and small t asymptotes. The approximate expression, denoted by Cq = C(t)/C# takes the form [26]
185
1
Cq(t) = ( l+tzq/o.q)l/q
q>0
(14)
Equation (14) gives the proper limiting behavior at small and large t. The model parameter q should be determined experimentally at an intermediate time (say, between 45 to 60% conversions). The best result is obtained at t = t* = C/z [26]. Once C(t*) is known, then q =
- tn2/[h~C(t*)/C:]
(15)
One can get a reasonably tight upper bound for C(t) by letting q = 1/z [26]. It should be stressed that the foregoing results also applicable to nonlinear kinetics [26]. The key message here is that, in developing process kinetics, in many cases, it is not necessary to characterize the whole feed. Instead, characterizing only the most refractory fraction and running the reaction at an intermediate conversion may suffice. In practice is should be easier to characterize the refractory species. After all, these are the species that will survive the reaction. 5. LIMITATION OF CONTINUUM APPROACH The continuum approximation will eventually break down after a sufficiently long time, since the number of surviving species is too small to justify the approximation. So time cannot be unconditionally large, even though the asymptotic kinetics are valid for large times. For firstorder reaction mixtures, the condition under which both the continuum theory and its long-time limit are valid has been established [ 16]. The domain of validity is as follows: 1/k* << t << 1/6
(16)
Here/c* is a characteristic rate constant for refractory species and 6 is the difference between the rate constants for two species whose reactivities are very close to each other. For HDS of distillates, we use Houalla et al.'s data obtained at 300~ and 10.5 MPa over a sulfided CoMo/A1203 catalyst [27]. We take k* as the rate constant for HDS of dibenzothiophene (DBT). The value of 6 is estimated by the rate constants for HDS of 4-methyl DBT and 4,6-dimethyl DBT (a better estimate would be the rate constant difference for 4,6-dimethyl DBT and 4,6diethyl DBT). Then the region of validity is 3.87.10 3 << t << 161-10 3 h.g-cat./cm3feed
(17)
which is not very stringent for practical purposes [ 16]. However, if one is interested in the regime of very long times (t > 1/6), a discrete approach may be preferable to a continuum one. 6. CONCLUDING REMARKS In this survey paper, a continuum theory is discussed to gain a better understanding of the diverse kinetic behavior of oil fractions in hydroprocessing. The focus is on the aggregated kinetics of the disappearance of a single total reactant (e.g., total sulfur in the feed). The treatment can be extended to problems where selectivity is of a major concern. For instance, the problem of maximizing the liquid yield in hydrocracking has been treated by introducing a distribution function characterizing the stoichiometry of hydrocracking reactions [28]. The
186
continuum approach has also been applied to modeling of thermal and catalytic processing of coals [29, 30].
7. REFERENCES
1. Aris, R., Arch. Ration. Mech. Anal., 27 (1968) 35. 2. Girgis, M. J. and Gates, B. C., Ind. Eng. Chem. Res., 30 (1991) 2021. 3. Inoue, S., Wada, Y. and Takatsuka, T., ACS Symp., Preprints, 43 No.4 (1998). 4. Ozaki, H., Satomi, Y., and Hisamitsu, T., Proc. 9th World Pet. Cong., 6 PD 18(4) (1976) 97. 5. Ho, T. C., and R. Aris, AIChE J., 33 (1987) 1050. 6. Ho, T. C., R. R. Chianelli, and A. J. Jacobson, Appl. Catal., 114 (1994) 131. 7. Astarita, G., and R. Ocone, AIChE J., 34 (1988) 1299. 8. Astarita, G., AIChE J., 35 (1989) 529. 9. Aris, R., AIChE J., 35 (1989) 539. 10. Chou, M. Y., and T. C. Ho, AIChE J., 35 (1989) 533. 11. Chou, M. Y., and T. C. Ho, AIChE J., 34, (1988) 1519. 12. Ho, T. C., B. S. White, and R. Hu, AIChE J., 36, (1990) 685. 13. Miller, J. T., and Hineman, M. F., J. Catal., 85 (1984) 117. 14. Ho, T. C., Catal. Rev.-Sci. Eng., 30, (1988) 117. 15. Krambeck, F. J., ISCRE 8, Inst. Chem. Eng. Symp. Ser., 87 (1984) 733. 16. Ho, T. C., and B. S. White, AIChE J., 41 (1995) 1513. 17. Ho, T. C., AIChE J., 42 (1996) 214. 18. Gray, M.R., Upgrading Petroleum Residues and Heavy Oils, Marcel Dekker, NY, 1994. 19. van Dongen, R. H. D., Bode, H. van der Eijk, and J. vanKlinken, Ind. Eng. Chem., Proc. Dev., 19 (1980) 630. 20. Luss, D., and Golikeri, AIChE J., 21 (1975) 865. 21. Astarita, G., and A. Nigam, AIChE J., 35 (1989) 1927. 22. Aris, R., Kinetics and Thermodynamic Lumping of Multicomponent Mixtures, G. Astarita and S. I. Sandler, (eds.), Elsevier, Amsterdam 1991. 23. Golikeri, S., and D. Luss, Chem. Eng. Sci., 26 (1971) 237. 24. Ocone, R. and Astarita, A., AIChE J., 39 (1993) 288. 25. Ho. T. C., B. Z. Li, and J. H. Wu, Chem. Eng. Sci., 50 (1995) 2459. 26. Ho, T. C., Chem. Eng. Sci., 46 (1991) 281. 27. Houalla, M., Broderick, D. H., Sapre, A. V., Nag, N. K., deBeer, V. H. J., Gates, B. C., and Kwart, H., J. Catal. 61 (1980) 523. 28. Laxminarasimhan, C. S., Verma, R. P., and Ramachandran, P. A., AIChE J. 42 (1996) 2645. 29. Wang, M., Zhang, C., Smith, J. M., and McCoy, B. J., AIChE. J., 40 (1994) 131. 30. Prasad, G. N., Agnew, J. B., and Sridhar, T., AIChE J., 32 (1986) 1288.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
187
M o l e c u l a r Base A p p r o a c h e s by G C - A E D to H D S of Gas Oil on Sulfide Catalysts I. Mochida a, S. Shin a, K. Sakanishi a, D. Grudoski b and J. Shinnb aInstitute of Advanced Material Study, Kyushu University, Kasugakoen, Kasugasi, Fukuoka 816, Japan bChevron Research and Technology Company, Chevron Way 100, Richmond, Califonia, USA
Abstract GC-AED is a powerful instrument for analyses of every molecular components in the gas oil. Intrinsic and partner dependent reactivities of sulfur species in gas oils are determined in HDS by aid of GC-AED. 1. Introduction Deep refining of gas oil has been requiring more research to fulfil the regulation continuously tightened for the environmental protection. More active catalysts have been looked for these years. So far, CoMo and NiMo sulfides on alumina have solved tasks to keep regulation. However, 50ppm sulfur level of gas oil recently targeted within 5-10 year requires better process design as well as better catalyst. Molecular identification, reactivity, and inhibition of every molecular species in gas oils can be bases to approach the targets. The present communication reports the GC-AED analyses of gas oils to correlate their desulfurization reactivities, taking accounts of molecular aromatic and nitrogen as well as sulfur species. 2. Experiment 2.1. Materials 4 kinds of gas oils, such as Iran light oil(ILO), light cycle oil(LCO), medium cycle oil(MCO), and FSL cut2(FSL), were used in this study. Some properties and elemental analyses of each gas oils are given in Table 1. Table 1. Composition of Various Gas Oils ILO LCO Boiling range, ~ NDA 180-343 Density (15 ~ 0.85 0.89 Carbon, wt%" 85.21 89.90 Hydrogen, wt% a 13.51 9.56 Sulfur, wt% b 1.19 0.42 Nitrogen, ppm a 90 345 *NDA : no data available aElemental analysis, bGC-AED
MCO NDA 0.89 90.19 9.06 0.47 695
FSL 208-275 NDA 85.97 13.05 0.91 295
2.2. Characterization of gas oils Molecular characterization of gas oils was performed with GC-AED (HP 6890/G2350A) and GC-MS (HP 6890/5973). In both instruments, HP-1MS was used as a column and temperature ranges was 40 to 320 ~ by 10~ of heating rate. In
188
the case of GC-AED analysis, C179, $181, and N388 atomic lines were selected for analyzing carbon, sulfur and nitrogen species, respectively. The detail analysis conditions of AED are summarized in Table 2. Table 2. AED Analysis Conditions
AED conditions Transfer line temperature, ~ Cavity temperature, ~ Hydrogen reagent, psi Oxygen reagent, psi Methane reagent, psi Make-up gas, mL/min
Carbon and Sulfur 350 350 43 55 -
100
Nitrogen 350 350 40 80 50 230
The content of aromatic component in gas oil was determined by column separation and HPLC according to the method in reference [1 ]. 2.3. HDS reaction
HDS reaction was carried out by 100ml of autoclave reactor with sampling apparatus. HDS reaction conditions are given in Table 3. Catalysts were commercially available and they were presulfided at 360 ~ in H2 with 10wt% of H2S. Table 3. HDS Reaction Condition of Gas Oils
Catalyst Catalyst/Oil by weight Temperature, ~ H2 pressure, MPa Time, min
CoMo/A1203, NiMo/A1203 0.1 340 5 10,20,30,60,90
3. Results 3.1. Characterization and identification of sulfur and nitrogen species in gas oils
Sulfur species compete with other sulfur molecules, nitrogen and aromatic species in gas oils for the active sites on catalyst. Therefore, the characterization and identification of each species become necessary to explain factors affecting the desulfurization reactivities of sulfur molecules in gas oil. Figure 1 shows the S and N chromatograms of 4 kinds of gas oils measured by GCAED. As seen in S chromatograms, each gas oils showed different distribution of sulfur species depending on their origins and cutting points. LCO contained benzothiophene(BT) and its derivatives as major sulfur components. In MCO, dibenzothiophene(DBT) and its derivatives were major sulfur components. ILO and FSL showed similar distributions of sulfur species, carrying BT and DBT derivatives as major components. As for nitrogen compounds, all of gas oils except for LCO contained carbazole and its derivatives as nitrogen species. In MCO and FSL, carbazole, mono- and dimethylated carbazoles and in ILO, tri-methylated carbazoles were major species. In contrast, LCO carries aniline derivatives, indole and its derivatives as major nitrogen species.
189
Figure 1. S and N chromatograms of various gas oils measured by GC-AED. Table 4 shows the composition of heteroaromatics and aromatics in each gas oils. The order of sulfur content was ILO > FSL > MCO > LCO. While, the order of nitrogen content was MCO > LCO > FSL > ILO. Table 4. Composition of Heteroaromatics and Aromatics in GOs ILO LCO MCO FSL Total S, wt% 1.19 0.42 0.49 0.91 Total N, ppm 90 345 695 295 Basic N, order* 3 1 4 2 PAH, wt% 2.0 1.7 11.3 6.4 * order of quantity determined by basic fraction
3.2. Reactivities of sulfur species in gas oils The relative sulfur content in gas oils was monitored as a function of reaction time and the result is displayed in Figure 2. Total reduction of sulfur species was the order of LCO, ILO, MCO and FSL irrespective of catalysts. Especially, LCO showed dramatical decrease of sulfur species within 10 min of reaction time. CoMo/A1203 was more active for desulfurization than NiMo/A1203, however the difference was trivial.
190 100 .,..,
.
.
.
100m
. ~
ID
"~
80
\
|
60
~
& MCO
4O ._ _~
ILO
,co
-~
e-.1
,4-
,,
if) ID
2O
{1}
n,"
._
0
0
'r~
20
. . . .
ID
I
40
60
n,'
A
80
100
80 [~X~,
El ILO
II\\ \
,
,oi II
\'~
~
~co MCO
9
9
FSL
2o
0
0
20
Reaction Time (min)
40
60
80
100
Reaction Time (min)
Figure 2. Total sulfur reduction of various gas oils as a function of hydrotreatment time(left" CoMo/A1203, right 9NiMo/A1203). To explain the desulfurization reactivities of sulfur species on molecular base, the S chromatograms of above gas oils were investigated as a function of hydrotreatment time. Since the intensity reduction of S chromatograms in HDS reaction were almost the same on both catalysts, results of NiMo/A1203 are illustrated as an example in Figure 3. BT and its derivatives were major contributors to the total reduction of sulfur species regardless of gas oils. In addition, they disappeared at early stage of HDS reaction and almost completely removed within 60 rain. ,.
{'3-BT
:] -0.
0 min
14, , I l,
,~. ,,, ,L~., .~L.....
,
ILO
1
o min
{ '2-B l" {'3-B'I i, I
('I-BT ,o
i[
30min
"2t '2'
] [[
90 min
m~..~
2]
i
)1
t .[J~l,,'~.....
30min 7 2 - = x -
90 min
.~_=_ _,
I ,-
J-
t,,#, . . -. .T. .i
,.!
0 min
""
I ,[I,'3-BT [1
LCO
::1
[11 . . . . . . .
MCO ~ .......-=
_ _
]
"n
4MDB'r 46I)MDBT
l BT
I
~ L ,.
-o,o
,2
,,
.
:![ ,0~
"
I
30 min
"_
_7_2,_. . . . .
] --:~-"'-----J . . . . . . . . . .
___
1
*:~. . . . .
90 min +
Figure 3. Sulfur species in gas oils monitored by GC-AED in a time course of hydrotreatment (NiMo/A1203 catalyst at 340~ 5MPa of H2 atmosphere). Some sulfur molecular species remaining during the HDS reaction were selected to clarify how desulfurization reactivites of sulfur species are influenced by the composition of gas oils. DBT, 4-methyldibenzothiophene(4-MDBT), and 4,6dimethylbenzothiophene(4,6-DMDBT) were selected for this purpose. The
191
desulfurization reactivities of each compounds are illustrated in Figure 4, 5, and 6, respectively.
Figure 6. Relative 4,6-dimethyldibenzothiophene content of various gas oils as a
function of hydrotreatment time(left: CoMo/A1203, right : NiMo/A1203). It is interesting to observe that 4-MDBT and 4,6-DMDBT showed different reactivity trends to that of DBT. The order in reactivities of such sulfur species described by their final conversion was ILO > FSL > LCO > MCO in both catalysts. Unexpectedly, the final conversion achieved by NiMo/A1203 and CoMo/A1203 did not show significant difference, although the conversion of 4,6-DMDBT appeared differently.
192
4. Discussion Intrinsic reactivities of sulfur species were known as the order of BT > alkyl BT > DBT > 4-MDBT > 4,6-DMDBT [2]. However, the desulfurization reactivities of above sulfur species in gas oils were different from their intrinsic reactivities depending on the composition of gas oils. This means that the desulfurization reactivities of sulfur species in gas oil are affected by coexisting partner molecules. The molecular species considered as partners are sulfur, aromatic, and nitrogen species because they and their products compete, inhibit and deactivate the active sites of the catalyst. Therefore, it is necessary to consider the contents of partners to interpret the HDS reactivities of sulfur species in respective gas oil. Total sulfur reduction in gas oils can be explained by the quantity of refractory sulfur species. Examining the major sulfur species in each gas oils, BT was in LCO, 4-MDBT (285ppm) and DBT (238ppm) in ILO, DBT (342ppm), 4-MDBT (337ppm), and 4,6-DMDBT (96ppm) in MCO, and 4-MDBT (321ppm), DBT (276ppm), and 4,6-DMDBT (241ppm) in FSL. Therefore, total sulfur reduction of gas oil was LCO >> ILO > MCO > FSL. Different reactivities of specified species in various gas oils can be explained by the extent of partner inhibition. HDS routes of DBT and 4,6-DMDBT have been reported different, the former is direct elimination of S while the latter is hydrogenation prior to elimination of S [3]. According to figures 4-6, the percentage conversion of above species are summarized as follows
Table 5. Percent Conversion of Specific Sulfur Species in Various GOs CoMo/A1203 NiMo/A1203 DBT 4-MDBT 4,6DMDBT DBT 4-MDBT 4,6DMDBT ILO 81.8 53.7 47.9 79.2 51.3 46.4 LCO 69.0 28.1 20.5 78.1 29.7 20.5 MCO 57.9 19.5 13.7 71.0 19.5 18.1 FSL 67.0 35.6 32.3 58.5 36.2 35.2 The reactivity of DBT over CoMo/A1203 was in the order of ILO > LCO, FSL > MCO. Its reactivity over NiMo/AI203 was in the order of ILO, LCO > MCO > FSL. In contrast, as for 4-MDBT and 4,6-DMDBT, the order of reactivity defined by achieved conversion was ILO > FSL > LCO > MCO. To explain such trends, contribution of respective partner species was necessary. HzS produced in HDS could be counted from the conversion of total sulfur components. In CoMo/A1203 system, the amount of produced HzS within 20min, from ILO, LCO, MCO, and FSL were 0.61, 0.38, 0.27, and 0.38wt%, respectively. In NiMo/AI203 system, they were 0.52, 0.40, 0.25, and 0.30wt%. Considering these values and Table 4, the reactivity of DBT is affected mainly by H2S. Nitrogen species may also influence the reactivity. As for 4-MDBT and 4,6-DMDBT, their reactivities over NiMo/AI203 give some clue to understand the partner inhibition. The order of reactivities defined by initial stage was ILO > LCO > FSL > MCO. This means that the inhibiting partners of 4-MDBT and 4,6-DMDBT are principally nitrogen and aromatic components. Such order was changed to ILO > FSL > LCO > MCO as HDS advanced. This implies that HzS also acts as an inhibition partner in this state. In consequence, the influences of partners on direct elimination of S and hydrogenation prior to elimination of S must be different. Inhibition factors of every component are to be determined for individual species, HzS, aromatic, total and basic nitrogen species. Catalyst deactivation, inhibitor's production and disappearance may influence the final achievement of HDS.
193
References
1. M. Dorbon, I. Ignatiadis, J. Schmitter, P. Arpino,G. Guiochon, H. Toulhoat, and A. Huc, Fuel, 63 (1984) 565. 2. X. Ma, K. Sakanishi, and I. Mochida, Ind. Eng. Chem. Res., 33 (1994) 218. 3. D. Whitehurst, T. Isoda, and I. Mochida, Adv. Catal., 42 (1998) 345.
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
The Nitrided CoMo Hydrodenitrogenation
Catalysts
195
for
Hydrodesulfurization
and
M. Nagai, H. Koyama, S. Sakamoto, and S. Omi Graduate School of Bio-applications and Systems Engineering, Tokyo University of Agriculture and Technology, 2-24 Nakamachi, Koganei, Tokyo 184-8588, Japan
Abstract The activity of CoMo/A1203 catalysts nitrided in a stream of ammonia was tested for dibenzothiophene HDS and carbazole HDN. The surface property of the catalysts was determined using the XPS spectroscopy. The 600~ nitrided CoMo/A1203 catalyst was a very active catalyst for dibenzothiophene HDS with a highly selective C-S bond breakage of dibenzothiophene at 260~, compared with the sulfided catalyst. The 600 C nitrided CoMo/AI~O3 catalyst was also more active than the other nitrided catalysts and hydrogen-reduced catalysts for carbazole HDN at 300~ The ammonia treatment at higher than 500~ increased the C-S hydrogenolysis in dibenzothiophene HDS and the C-N hydrogenolysis in carbazole HDN.
1. INTRODUCTION It is well known that molybdenum compounds catalyze the hydrodesulfurization (HDS) process of petroleum feedstocks by acting as active catalyst. An increasing interest has developed in exploring the catalytic properties of transition metal nitrides such as molybdenum and tungsten nitrides [1-5]. Molybdenum nitride is reported to be an active catalyst to compete with the catalytic properties of noble metals [5,6]. We reported that nitridation of M o / A 1 2 0 3 catalysts enhanced the activity and selectivity for the direct desulfurization of dibenzothiophene to form biphenyl, compared with the sulfided catalyst [1,2]. Hydrodenitrogenation (HDN) has also received considerable attention during the past decades to eliminate nitrogen compounds in refining of petroleum feedstocks by hydroprocessing. Molybdenum nitrides are required as novel catalysts to effectively reduce nitrogen content, hydrogen consumption, and catalyst deactivation [7-10]. Schlatter et al. [9] have shown that unsupported molybdenum nitrides offered an interesting alternative to existing HDN catalysts. It was reported that molybdenum nitride exhibited HDN activity (per one Mo atom) comparable to that of commercial sulfided NiMo/A120~ catalysts but exhibited much higher selectivity for the formation of aromatic products [8,9]. Recently, a few studies on the structure and properties in nitriding of unsupported CoMo catalyst [11-14] have been reported, but alumina-supported CoMo nitrided catalysts are few reported to be used as the catalysts for the HDS and HDN reactions [15]. In this study, the effects of nitriding conditions on the activity and selectivity of the nitrided CoMo/A1203 catalyst for dibenzothiophene HDS and carbazole HDN were studied in a high-pressure flow system. The surface properties of the nitrided CoMo catalyst were also discussed.
196
2.
EXPERIMENTAL
A 4% CoO 12% MoOJA1203 (Ketjenefine 124) was calcined at 450~ for 3 h and nitrided with pure ammonia at 41/h from 400 to 500"~.900~ at 1 ~ held at this temperature for 3 h, and then cooled to room temperature in flowing ammonia [1,10]. The nitrided catalysts were passivated in 1% O2/He for 24 h. For measurement of the catalyst activity, the passivated CoMo/A12Q catalysts were reduced at 500 0(3 for 1 h to remove water before the reaction feed was introduced into the reactor. The catalyst was reduced at 4000(3 for 3 h in carbazole HDN for comparison. The activity measurements for dibenzothiophene HDS at 2600(3 and carbazole HDN at 300~ were carried out using a fixed-bed microreactor at 10.1 MPa total pressure. The feed, consisting of 0.25wt% dibenzothiophene or carbazole dissolved in xylene, was introduced into the reactor at a rate of 20 ml/h with hydrogen flow of 6 1/h. The rates for dibenzothiophene HDS and carbazole HDN were calculated by the formation of the hydrocarbons produced during the reaction. The surface composition and property of the nitrided CoMo/A1203 catalysts were measured using XPS spectroscopy.
3.
R E S U L T S AND D I S C U S S I O N
3.1. The H D S of d i b e n z o t h i o p h e n e
14 12 "7
10
o
~
4
o ~
2
0
1
2
3
4
5
6
7
8
Reaction Time [h] Figure 1. Reaction products in the HDS of dibenzothiophene on the CoMo/A12Q catalyst nitrided at 6000(2 with time on stream at 260 ~ . 9 Biphenyl, A cyclohexylbenzene, and {} dibenzothiophene.
The reaction products in dibenzothiophene HDS on the 600 0(3 -nitrided CoMo/A1203 catalyst with time on stream at 2600(3 and total pressure of 10.1 MPa are shown in Fig. 1. The major product was biphenyl along with a small amount of cyclohexylbenzene. No other hydrogenated products were observed in the reaction at 2600(3. The 6000(3 nitrided catalyst held a high activity (conversion of 75%) at the first hour but sharply dropped to 20% of conversion 7 h after the reaction started. The HDS rates of the nitrided catalysts for dibenzothiophene HDS at 260~ are shown in Fig. 2. The 600 ~ nitrided CoMo/A1203 catalyst (per unit of gram) was more active t h a n the catalysts nitrided at 400 and 800 ~ . The 800 0(3 nitrided catalyst was more active than the 4000(3 nitrided catalyst at the first stage, but the 400 0(3 nitrided catalyst became more active after 3 h of time on stream. This suggested
197
140 ~-~ 120 ~/x ~
100
\,,.~
~o
~ 40 f~o 0
~ ~ = : ~ - - -
/ J
I
0
1
,
I
,
I
2
I
I
3
,
I
4
,
I
5
,
I
6
i
7
8
Reaction Time [h] F i g u r e 2. The HDS of dibenzothiophene on the ( O ) 400, ( ~ ) 600, and (El) 800~ nitrided CoMo/A120~ catalysts as a function of time on s t r e a m at 260~ 100
(D
80
6o O
40
-,,
x,x, '
20
0
.
0
J.
I
1
,
I
2
i
I
3
~
I
4
,,
,I
,
5
I
6
,
I
,
7
Reaction Time [h] Figure 3. The selectivity of the (O) 400, (A) 600, and ([3) 800~ CoMo/A12Oa catalysts for the C-S bond breakage of dibenzothiophene.
8
198
Biphenyl
Cyclohexyibenzene
2I--I~ Dibenzothiophene
1,2,3,4-Tetrahydrodibezothiophene
1,2,3,4,4a,9aHexahydrodibezothiophene
Figure 4. Reaction scheme for the HDS of dibenzothiphene.
that the nitriding treatment at high temperature created high active species but were deactivated easily during the reaction. Furthermore, the selectivity of the 600 ~ nitrided catalyst for the direct C-S hydrogenolysis to hydrogenation of dibenzothiophene was estimated as the molar ratio of biphenyl to cyclohexylbenzene and is shown in Fig. 3. The HDS of dibenzothiophene proceeded in this scheme (Fig. 4). This reaction takes place via the direct C-S hydrogenolysis of dibenzothiophene to form biphenyl or via hydrogenation of dibenzothiophene followed by the direct sulfur removal from dibenzothiophene. The molar ratio of biphenyl to cyclohexylbenzene in the HDS reaction at 7 h is 8.0 and higher on the other catalysts. Although the 8000(; nitrided catalyst held a similar conversion as the 600~ nitrided catalyst at 260~ the HDS rate and sulfur removal selectivity for the 800~ nitrided catalyst were lower than those of the 600~ nitrided catalyst because the hydrogenated compounds such as tetrahydrodibenzothiophene were more produced. Since the molar ratio for the sulfided CoMo/A1203 catalyst was 4.3 in dibenzothiophene HDS at 260~ [16], the nitrided CoMo/A12Q catalysts were more selective than the sulfided CoMo/A12Q catalyst for the C-S bond breakage of dibenzothiophene to produce biphenyl with less hydrogen consumption. 3.2. The H D N of C a r b a z o l e
The product distribution in carbazole HDN on the 600 ~ nitrided CoMo/A1203 catalyst at 300~ and 10.1 MPa total pressure is shown in Fig. 5. The major reaction products were bicyclohexyl and tetrahydrocarbazole with a small amount of 2-ethylbicyclo[4.4.0]decane. No hydrogenated compounds were observed except for tetrahydrocarbazole. The reaction reached a steady state 6 h after the reaction started. The activity of the CoMo/A1203 catalyst decreased initially with time on stream and reached steady state. A plausible reaction scheme for the HDN of carbazole is shown in Fig. 6. Carbazole was successively hydrogenated to perhydrocarbazole on the CoMo catalyst, although the hydrogenated compounds except for tetrahydrocarbazole were not observed [10,17]. Bicyclohexyl was produced during C-N hydrogenolysis of perhydrocarbazole, after the consecutive hydrogenation of carbazole to perhydrocarbazole. Bicyclohexyl was isomerized to 2-ethylbicyclo [4. 4. 0] decane.
199
14 12
10
6 ~ @
0__-.0----0---
4
0 0
1
2
3
4
5
6
7
8
Reaction Time [h] Figure 5. Reaction products in carbazole HDN on the 600~ CoMo/A12Oa catalyst with time on stream at 300~ 89 Tetrahydrocarbazole, O bicyclohexyl, A 2-ethylbicyclo [4.4.0] decane, and 9 carbazole.
H
H
Carbazole
1,2,3,4-Tetrahydrocarbazole
H
Perhydrocarbazole
CBicyclohexyl
\
CIHj
Figure 6. Reaction scheme for the HDN of carbazole.
2- Eth ylbicyclo[4.4.0]decane
Shown in Fig. 7 is the rate of carbazole HDN as a function of nitriding temperature at reaction temperature of 300 ~ The 600 ~ nitrided CoMo/A1203 catalyst showed a high activity for the C-N hydrogenolysis in carbazole HDN but the 400~ nitrided catalyst was extremely less active than the other nitrided catalysts. The HDN rates for the CoMo/A12Q catalysts nitrided at 700 and 9000(2 were higher than those nitrided at 500 and 600~ Although the 900~ nitrided catalyst was expected to shrink and sinter as a result of ammonia treatment at 900 C, the catalyst showed even a high activity for carbazole HDN. Furthermore, the 400~ nitrided catalyst was even less active than the 400~ reduced catalyst (HDN rate' 18 gmol~g) in the C-N hydrogenolysis of carbazole HDN, although the hydrogenated compounds were produced at a similar amount for both catalysts. The 400~ nitrided catalyst was as active as the 500~ reduced catalyst (HDN rate: 27 gmol/hg). This is probably because nitrogen species formed by ammonia decomposition at 400~ [12,20] prohibited the C-N hydrogenolysis in carbazole HDN. However, the
200
70 6O 5o 9
:t 9
40 30
2; C~ 2O 10 I
I
I
I
J
J
400 500 600 700 800 900 Nitriding Temperature [~ Figure 7. The rates for carbazole HDN as a function of nitriding temperature.
CoMo/A1203 catalysts nitrided at 500~ and higher than 500 ~ were extremely more active t h a n the 400~ reduced catalyst. Therefore, the molybdenum nitride species created by ammonia treatment at high temperature is considered to be responsible for the C-N hydrogenolysis in the HDN reaction, compared to the reduced catalyst. Furthermore, the catalyst nitrided at 400 ~ and 5 atm of ammonia was very active at the initial stage of the reaction 3 h after the reaction started, but decreased sharply to the one fifth activity at 7 h. Although the catalyst nitrided at 400 ~ with ammonia at 5 atm was more active than the hydrogenreduced catalyst for C-N hydrogenolysis, both catalysts showed a similar conversion during the reaction. Therefore, the catalyst nitrided at 5 atm of Table 1.
Surface composition (N/Mo) of the nitrided CoMo/A12Q catalyst XPS N lsa/Mo 3d
Catalyst nitrided at
bb
1st layer d 2nd layer
0.31 0.24
400~
ac
0.42 0.41
b 0.42 0.31
500~
a
b
0.63 0.62
0.35 0.28
700~ 0.32 0.27
The N ls spectra envelope overlapped with the Mo 3pa;2 spectra. the ratio, [cps" eV/cps" eV] b before the reaction. c after the reaction. d argon etching time: 0 min, 1st layer; 5 min, 2nd layer.
201
ammonia led to the increase in the HDN activity. The HDN rate and hydrocarbon concentration on the catalyst nitrided at 400~ and 1 atm of ammonia gradually increased 1.3 times higher than the catalyst nitrided with 5 atm at the 7 h-run. 3.3. X P S A n a l y s i s The surface composition (ratio of (N ls and Mo 3p) to Mo 3d) of the nitrided CoMo catalysts was determined using X-ray photoelectron spectroscopy. The ratios of (N ls and Mo 3p)/Mo 3d in the XPS spectra for the CoMo/A12Q nitrided at 400, 500, and 700~ before and after the HDN reaction are shown in Table 1. The ratios of (N ls and Mo 3p)/Mo 3d of the 400 and 500~ nitrided catalysts increased after the reaction but that for the 700~ nitrided catalyst decreased. From the results, although nitrogen was released from the 700~ nitrided catalyst during the reaction, nitrogen was not lost for the 400 and 500~ nitrided catalysts but nitrogen species were accommodated on the surface of the catalysts, and then the activity for C-N hydrogenolysis gradually increased.
4. CONCLUSIONS The 600~ catalyst was more active than the other nitrided catalysts for both dibenzothiophene HDS and carbazole HDN. The 400~ nitrided catalyst was the least active of the nitrided and reduced catalysts for carbazole HDN. The molybdenum nitride catalyst was more selective for the C-S bond breakage of dibenzothiophene to produce biphenyl. Although the CoMo/A120a catalyst nitrided at more than 500~ was more active than the 400~ reduced catalyst. The nitridation of the CoMo/AI~Q catalyst at 5 atm pressure of ammonia remarkably promoted the initial activity compared to that at 1 atm pressure of ammonia for carbazole HDN, but deactivated during the reaction.
5. ACKNOWLEDGEMENTS
This work has been carried out as Petroleum Institute commissioned by the support of the Ministry of International Grant-In-Aid for Scientific Research of (09555244).
a research of project of the Japan Petroleum Energy Center with the Trade and Industry as well as the the Ministry of Education Grant
6. R E F E R E N C E S
1. 2. 3. 4.
M. Nagai, T. Miyao, and T. Tsuboi, Catal. Lett., 18 (1993) 9. M. Nagai, A. Irisawa, S. Omi, J. Phys. Chem. B., 102 (1998) 7619. U.S. Ozkan, L. Zhang, P. A. Clark, J. Catal. 172 (1997) 294. K.R. McCrea, J. W. Logan, T. L. Tarbuck, J. L. Heiser, M. E. Bussell, J. Catal. 171 (1997) 255. 5. D.J. Sajkowski, S. T. Oyama, Appl. Catal. A.' Gen. 134 (1996) 339. 6. G.M. Dolce, P.E. Savage, L. T. Thompson, Energy Fuels 11 (1997) 669. 7. M. Nagai, T. Miyao, Catal. Lett., 15 (1992) 105. 8. C.W. Colling, L. T. Thompson, J. Catal. 146 (1994) 193; J.-G. Choi, J. R.
202
9. 10. 11. 12. 13. 14. 15. 16. 17. 18.
Brenner, L. T. Thompson, J. Catal., 154 (1995) 33. J.C. Schlatter, S. T. Oyama, J. E. Metcalfe, III, J. M. Lambert, Jr., Ind. Eng. Chem. Res., 27 (1988) 1648. M. Nagai, A. Miyata, M. Kiyoshi, K. Hada, K. Oshikawa, J. Catal. 182 (1999) 292. I. K. Milad, K. J. Smith, P. C. Wong, K. A. R. Mitchell, Catal. Lett. 52 (1998) 113. C. C. Yu, S. Ramanathan, S. T. Oyama, J Catal. 173 (1998) 1. S. Ramanathan, C. C. Yu, S. T. Oyama, J. Catal. 173 (1998) 10. D.-W. Kim, D.-K. Lee, S.-K. Ihm, Catal. Lett. 43 (1997) 91. H. K. Park, J. K. Lee, J. K. Yoo, E. S. Ko, D. S. Kim, K. L. Kim, App. Catal. A: Gen. 150 (1997) 21. M. Nagai, J. Jpn. Petrol. Inst. 38 (1995) 52. M. Nagai, T. Masunaga, N. Hanaoka, Energy Fuels, 2 (1988) 645. M. Nagai, A. Miyata, T. Kusagaya and S. Omi, "The Chemistry of
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
203
HDS o f D i b e n z o t h i o p h e n e a n d V a n a d y l P o r p h y r i n HDP o n b u l k Fe-Mo mixed sulphides. M.A. Luis 1,2,3, A. Rives 2, R. H u b a u t 2, B.P. E m b a i d 1, F. Gonzalez-Jimenez I a n d C. E. Scott 1,*. 1. Universidad Central de Venezuela, Centro de Catklisis Petr61eo y Petroquimica, Apartado Postal 47102, Los C h a g u a r a m o s , Caracas, Venezuela. 2. Universit6 des Sciences et Technologie de Lille, Laboratoire de Catalyse H6t6rog6ne et Homog6ne, URA CNRS n~ B&timent C3, 5 9 6 5 5 Villeneuve d'Ascq, France 3. Universidad de Carabobo, F a c u l t a d de Ciencias y Tenologia, Dpto de Quimica. * e-mail: [email protected]
Abstract. Bulk Fe-Mo mixed s u l p h i d e s were p r e p a r e d by h o m o g e n e o u s precipitation. The obtained solids were c h a r a c t e r i s e d by chemical analysis, X-ray diffraction, X-ray photoelectron spectroscopy, S7Fe M 6 s s b a u e r spectroscopy, a n d nitrogen a d s o r p t i o n (BET) for surface a r e a determinations. HDS of d i b e n z o t h o p h e n e (in a b a t c h reactor) a n d HDP of vanadyl octaethyl porphyrin, at high p r e s s u r e , were u s e d as catalytic tests. It was found t h a t Fe strongly p r o m o t e s Mo for both reactions, with a more m a r k e d synergy for the HDP t h a n for the HDS. S7Fe M 6 s s b a u e r spectroscopy suggests t h a t Fe exist in only one p h a s e , for the mixed Fe-Mo sulphides. This p h a s e could be the results of some Fe s u b s t i t u t i o n in the MoS2 s t r u c t u r e . For HDP the activity is m a x i m u m at a Fe(Fe+Mo) atomic ratio of 0.68, a n d increases linearly with the weighted average hyperfine field, which is related to c h a n g e s in the density of states of d electrons at the Fermi level, clearly suggesting t h a t synergy is related to a n electronic effect. For HDS of DBT the activity m a x i m u m is located at a Fe/(Fe+Mo) atomic ratio of 0.52. 1. INTRODUCTION There is no d o u b t t h a t h y d r o t r e a t m e n t will become increasingly i m p o r t a n t in years to come. The need for processing heavier p e t r o l e u m fractions into light distillates as a result of the decline in fuel oil d e m a n d , together with stricter e n v i r o n m e n t a l legislation regarding the m a x i m u m c o n t e n t of s u l p h u r , nitrogen a n d a r o m a t i c s in fuels, have m a d e h y d r o p r o c e s s i n g increasingly difficult. One possible way to overcome these difficulties is to develop new, more active a n d selective, or less expensive, h y d r o t r e a t i n g catalysts. Fe is
204
generally regarded as a poorer promoter t h a n Co or Ni to Mo(1), and for t h a t it has been less frequently used. However, the beneficial effect of Fe, in catalysts formulations, have been shown in previous work(2,3). Thus, w hen Fe is added to V in Fe-V mixed sulphides, a weak synergy is found for hydrodesulphurization (HDS) of thiophene(2), and toluene hydrogenation a n d vanadyl octaethyl porphyrin (VOOEP) hydrodeporphyrinization (HDP). Also, it h a s been proposed(4) t hat Fe-Mo catalysts can be as active as Co-Mo catalysts (but less t h a n Ni-Mo). Thus, it has been found t hat Fe can be an efficient Mo promoter for HDS of thiophene, provided t hat it is kept in the ferrous state during preparation and activation of the catalyst(4) . On the other hand, Ho et. al(5-7) have prepared and studied Fe u n s u p p o r t e d Mo, a n d Fe-Co-Mo and Fe-Ni-Mo catalysts formed from thermal decomposition of bis(diethylenetriamine) iron thiomolybdate. Fe-Mo sulphides were quite selective towards hydrodenitrogenation (HDN) relative to HDS. They also found th at it consisted of a single sulphided phase, which during activity testing is partially transformed into a mixture of Fe sulphide and a MoS2 like phase. They also showed(7) t h a t HDS a nd HDN activities of the Fe-Mo catalysts can be significantly increased by promotion with Ni or Co. In a preliminary work(8) we have d e m o n s t r a t e d t h a t a m a r k e d synergetic effect, for HDP of VOOEP, is present in bulk Fe-Mo mixed sulphides. M6ssbauer spectroscopy revealed the existence of two p h a s e s for the mixed sulphides, one of which was proposed to be responsible for the HDP activity. In the present work we carry out a detailed study of Fe-Mo mixed sulphides in hydrotreatment. The catalysts were characterised by chemical analysis, X-ray photoelectron and M6ssbauer spectroscopies, as well as surface area determinations (BET). Dibenzothiophene (DBT) HDS a n d VOOEP HDP, in a continuos flow system were also tested. 2. EXPERIMENTAL.
Bulk Fe-Mo mixed sulphides were prepared according to our previous reported method(2,3). Thus, an a q u e o u s solution of Fe(III) nitrate (MERCK > 99%) was slowly added with stirring to a solution of a m m o n i u m tetrathiomolybdate (STREM CHEMICAL) in 20 % v / v a m m o n i u m sulphide u n d e r nitrogen a t m os phe r e (the use of an inert at m osphere represents a change from our previous method). After filtering off the obtained solids were v a c u u m dried and sulphided in a H2S/H2 (15 % v/v) flow at 673 K for 4 h. The a m o u n t s of iron nitrated a nd a m m o n i u m tetrathyomolibdate (TTM) were worked out in order to have the desire Fe/(Fe+Mo) atomic ratio. Iron sulphide was prepared in the same m a n n e r but without any TTM in the solution, a n d m o l y b d e n u m sulphide was prepared by direct sulphidation of TTM. After the sulphidation step the Service Central D'Analyse of the CNRS (Vemaison, France) analysed the samples (for Fe, Mo and S). Surface area determinations (BET) were obtained by nitrogen adsorption, on the airexposed solids, in a Q u a n t a s o r b Instrument. X-ray photoelectron
205
s p e c t r o s c o p y (XPS) m e a s u r e m e n t s were carried out in a n AEI ES 200B s p e c t r o m e t e r equipped with a n AI anode working at 300 W. The atomic composition of the e x a m i n e d s a m p l e s were d e t e r m i n e d from the integrated Fe2p a n d Mo3d with a b a c k g r o u n d linearly s u b t r a c t e d . X-ray diffraction p a t t e r n s were obtained on a Siemens D500 diffractometer equipped with a Cu anode. S7Fe M 6 s s b a u e r s p e c t r a were recorded at room t e m p e r a t u r e in a t r i a n g u l a r s y m m e t r i c mode. The s p e c t r a were c o m p u t e r fitted with a c o m p u t e r p r o g r a m supplied by P. Bonville(9) u s i n g a hyperfine field distribution method(HPFD) capable of fitting u p to 40 s u b - s p e c t r a per distribution, with the s a m e values of the Isomer Shift (IS), the Q u a d r u p o l e Splitting (QS) a n d the full width at the half m a x i m u m of the lines. The results are expressed in a h i s t o g r a m of the proportions of the a r e a s of the different s u b s p e c t r a contributing to the complete s p e c t r u m . H y d r o d e p o r p h y r i n i z a t i o n (HDP) of vanadyl octaethyl p o r p h y r i n (VOOEP) was carried out at 573 K, in a h i g h - p r e s s u r e (80 bar) c o n t i n u o s flow (liquidsolid-gas) system. A solution of VOOEP (3 10 .4 mol I-I) in decaline containing 2%(v/v) of dimethyldisulfide was u s e d as a liquid feed. Experimental details are given elsewhere(3). Dibenzothiophene (DBT) hydrodesulfurization was carried out in a 100 cm3 b a t c h reactor. The solids were p r e s u l p h i d e d in situ with flowing H2S/H2 (10% v/v) at 573 K for 4 h, a n d at a t m o s p h e r i c p r e s s u r e . Then, 70 cm 3 of a solution of DBT in decaline (2.71x10 -2 mol I-I) were a d d e d to the reactor by m e a n s of a internal device in order to avoid a n y contact of the catalyst with air. The reaction was carried out at 573 K a n d 60 bars. Initial activities were worked out from the conversion v e r s u s time plots.
3. RESULTS.
X-ray diffraction p a t t e r n s show t h a t the p u r e iron sulphide h a s the pyrrhotite s t r u c t u r e , in a g r e e m e n t with previous reports(2,3), a n d the p u r e m o l y b d e n u m sulphide c o r r e s p o n d to a r a t h e r a m o r p h o u s MoS2. For the mixed Fe-Mo catalysts the diffraction p a t t e r n s are not very conclusive. BET surface a r e a s a n d chemical a n a l y s e s are reported in table 1. Table 1. Catalyst FeMo-0 FeMo-015 FeMo-030 FeMo-052 FeMo-068 FeMo-073 FeMo- 1
Fe/(Fe+Mo) {atomic) 0.00 0.15 0.30 0.52 0.68 0.73 1.00
Surface a r e a (m2.g -I) 33 43 32 44 49 39 6
206
BET surface a r e a s are of the s a m e o r d e r (within the e x p e r i m e n t a l error) for all the Fe-Mo mixed s u l p h i d e s , as well as for the Mo sulphide. The s u r f a c e a r e a of the Fe s u l p h i d e s is the lowest c o m p a r e to o t h e r catalysts, a n d this in a g r e e m e n t with previous findings(2,3), a n d with the fact t h a t XRD s h o w s this solid to be the b e t t e r crystallised. Activities for HDS of DBT a n d HDP of VOOEP are s h o w n in figures 1 a n d 2, r e s p e c t i v e l y . . We see t h a t for b o t h r e a c t i o n s p u r e Mo s u l p h i d e is m o r e active t h a n Fe sulphide, however, the difference is more m a r k e d for the HDS of DBT t h a n for the HDP of VOOETP. On the o t h e r h a n d a synergetic effect is o b s e r v e d in e a c h case. The m a x i m u m activity is observed for a Fe/(Fe+Mo) ratio of 0.52 for the HDS of DBT, a n d for a ratio of 0.68 for the HDP of VOOETP. We see a two-fold a n d a three-fold i n c r e a s e in activity, for HDS a n d HDP respectively, in relation to p u r e MoS2.
'7, m 12
,e,-
'm 30
~ ~
2s 20
x
9-
8
->
6
x
{ ms ~ lO
<
___e~ ~ a,
W
Q i
0,0
0,2
i
i
-r
T
0,4 0,6 0,8 Fe I(Fe+Mo) atomic
1,0
Figure 1. Specific activity for HDS of DBT
l
0,0
0,2
i
l
0,4 0,6 Fe I(Fe+Mo) atomic
i
0,8
1,0
Figure 2. Specific activity for HDP of VOOEP
Figure 3 s h o w s the fitted s p e c t r a t o g e t h e r with the c o r r e s p o n d i n g h i s t o g r a m on the right side. The H P F D ' s s h o w the s a m e s h a p e for all the catalysts, except for p u r e Fe sulphide. The HPFD is t a k e n b e t w e e n 0 a n d 300 K G a u s s fields. A r o u g h way to c h a r a c t e r i s e the d i s t r i b u t i o n is to c a l c u l a t e the weighted average of the HPFD (). It is clear t h a t the i n c r e a s e s as the iron c o n t e n t is i n c r e a s e d from .25 to .68 (see figure 4). The p a r a m e t e r s IS a n d QS are c o m m o n to all the Fe c o n t a i n i n g mixed catalyst. F r o m the s p e c t r a the values for IS (0.43+0.03 m m s-l), referred to ~-iron, a n d QS (0.03+0.01 m m s -I) c a n be w o r k e d out. This s u g g e s t s t h a t the Fe is p r o b a b l y
207
p r e s e n t as only one p h a s e in the mixed Fe-Mo solids s t u d i e d (this is m o r e evident for c a t a l y s t s from 0.25 to 0.73).
y=
0.1~
y = 0.52
y=0.68
y = 0.73
a
a a
V e l o c i t y ( m m . s "1)
~_.
H y p e r f i n e Field
,
(kG)
Figure 3. Fitted spectra (leit side) and HPFD histogram (right side) for bulk Fe-Mo sulphides.
208
XPS shows that Mo is in the form of MoS2, and that some sulphate is present in the surface of all solids. Sulphate is certainly due to some oxidation during sample handling (even though care was taken to avoid contact with moisture). It is also observed, by XPS, that surface concentration of Mo and Fe are very close to bulk concentrations (see figure 5), indicating a good homogeneity for the solids prepared.
The selectivities for tetrahydro dibenzo thiophene (THDBT), hexahydro dibenzo thiophene (HHDBT), biphenyl (BiPh), cyclohexyl benzene(PhCy) a n d
209
bicyclohexyl, for the HDS of DBT, are p r e s e n t e d in figure 6. An increase in the % of PhCy a n d BiCy a n d a decrease in the % of THDBT, for Fe promote catalysts,, in relation to MoS2(FeMo00), is observed, which is indicative of a higher h y d r o g e n a t i n g ability for the p r o m o t e d catalysts
4. DISCUSSION
The results obtained for HDP activities are c o n s i s t e n t with o u r previous reports (that is a synergetic effect is observed), but, the Fe(Fe+Mo) ratio for m a x i m u m activity is different (0,40 n o m i n a l in o u r previous p a p e r a n d 0.68 in this paper). More i m p o r t a n t are the differences in the M 6 s s b a u e r spectra, while before we identified two p h a s e s in the mixed Fe-Mo catalysts, now we see only one. However, the p h a s e identified in o u r previous work as a FeMoS active p h a s e for HDP is similar to the one p h a s e obtained in the catalysts p r e s e n t e d here. These differences could be due to differences in the p r e p a r a t i o n m e t h o d s , since before the catalysts were p r e p a r e d in air, while now they were p r e p a r e d u n d e r nitrogen a t m o s p h e r e . Obviously the u s e of an inert a t m o s p h e r e play a n i m p o r t a n t role in the type of solid obtained. Also, Ho et. ai(5-7) got only one p h a s e for b u l k Fe-Mo s u l p h i d e d catalyst, b u t according to the M 6 s s b a u e r p a r a m e t e r s their p h a s e is different to our. However, it is i m p o r t a n t to point out the good HDN activity p r e s e n t e d by this catalyst, a n d the good HDP activity we obtain with o u r preparation. On the other h a n d , the s h a p e of the M 6 s s b a u e r s p e c t r a (and their t h e r m a l behaviour, which will be reported elsewhere) are very similar to the one observed in the new Spin Density Waves (SDW) s y s t e m s (namely CuFeS2 a n d CuFeTe2), recently reported(10,11). In the aforementioned s y s t e m s Fe is a Fe 2§ in low spin state, so the magnetic properties observed are a t t r i b u t e d to the highly correlated d c o n d u c t i o n electron (itinerant antiferromagnetism). In the p r e s e n t case, the values for IS a n d QS justify the a t t r i b u t i o n of the spectra to Fe 2§ low spin, a n d the evolution of the at room t e m p e r a t u r e for the different relative c o n c e n t r a t i o n s of Fe a n d Mo are due to the c h a n g e s in the density of states of the d electrons at the Fermi level(12). It is clear from figure 5 t h a t there is a linear relation between the a n d the HDP activity. This correlation is less evident for the HDS of DBT, however, the two m a x i m u m s (for HDP a n d HDS) are very close. T h u s , one can consider t h a t we are in the p r e s e n c e of a n electronic effect (increase in the d electron density of states at the Fermi level). This could be a general explanation for the synergy found in different HDS catalysts. In fact, this new evidence allows u s to review previous proposition of electron t r a n s f e r in related systems(13-15). The electronic effect p r o p o s e d could not be, in view of the new evidence, a net electron t r a n s f e r from one metal to the other, b u t is r a t h e r a c h a n g e in the electron density of the bimetallic s y s t e m as a whole. This c h a n g e in electron density is the effect of the s u b s t i t u t i o n of one metal by a n o t h e r in the s a m e crystal s t r u c t u r e . A
210
mixing of the appropriate metal would c o n d u c e to a n o p t i m u m electron density of states of the d electrons in the c o n d u c t i o n band. 5. CONCLUSIONS.
A FeMoS p h a s e h a s been identified as the sole Fe containing p h a s e in b u l k Fe-Mo sulphided catalysts. This p h a s e could be the results of some Fe s u b s t i t u t i o n in the MoS2 structure. For HDP the activity is m a x i m u m at a Fe(Fe+Mo) atomic ratio of 0.68, a n d increases linearly with the weighted average hyperfine field , which is related to c h a n g e s in the density of d electrons at the Fermi level, clearly suggesting t h a t synergy is related to a n electronic effect. For HDS of DBT the activity m a x i m u m is located at a Fe/(Fe+Mo) atomic ratio of 0.52. 6. A C K N O W L E D G E M E N T S .
The a u t h o r s gratefully acknowledge the contribution m a d e by the F r e n c h Venezuelan PICS 324, a n d to CONICIT for its financial s u p p o r t t h r o u g h project G - 9 7 0 0 0 6 5 8 a n d BID-CONICIT QF15.
7. R E F E R E N C E S .
1. M. T e m a n , J. Catal., 104(1987)256. 2. C. E. Scott, B. P. Embaid, M. A. Luis, F. Gonzalez-Jimenez, L. Gengembre, R. H u b a u t a n d J. Grimblot, Bull. Soc. Chim. Belg., 104(1995)331. 3. C. E. Scott, B. P. Embaid, F. Gonzalez-Jimenez, R. H u b a u t a n d J. Grimblot, J. Catal., 166 (1997) 333. 4. J.L. Brito a n d A.L. Barbosa, J. Catal., 171(1997)467. 5. J.Y. Koo a n d T.C. Ho, Catl. Letters, 28(1994)99. 6. T.C. Ho, A. I. lacobson, R. R. Chianelli and C. R. F. Lund, J. Catal., 138 (1992) 351. 7. T. C. Ho, R. R. Chianelli and A. I. lacobson, Appl. Catal., 114(1994)127. 8. B. P. Embaid, M. A. Luis, C. E. Scott a n d F. Gonzalez-Jimenez, Hyperfine Interactions(C), 3(1998)96. 9. P. Bonville, S P E C / C E N Saclay, France. 10. A. Ribas, F. Gonzalez-Jimenez, L D'Onofrio, E. J a i m e s , M. Quintero a n d J. Gonzalez. Hyperfine Interactions 113( 1998)493. 11. F. Gonzalez-Jimenez, A. Ribas, E. J a i m e s , L D'Onofrio, M. Q u i n t e r o Quintero a n d J. Gonzalez. Phys B, In press. 12. P. C. H. Mitchell a n d C. E. Scott., Bull. Soc. Chim. Belg., 93(1984)619. 13. P. C. H. Mitchell a n d C. E. Scott, J. P. Bonnelle, J. G. Grimblot., J. Catal., 107(1987)482. 14. C. E. Scott, P. Betancourt, M. J. P6rez Zurita, C. Bolivar a n d J. Goldwasser., ApI~I. Catal. Submited.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
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D e s i g n o f a m o d e l a c t i v i t y t e s t / ' o r s e c o n d stage d e e p H D S catalysts. H.R. Reinhoudt a, M. van GorseP, A.D. van Langeveld a, J.A.R. van Veen b, S.T. Sie a and J.A. Moulijn a
a Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands b Shell Research and Technology Centre Amsterdam, Badhuisweg 3, 1031 CM Amsterdam, The Netherlands
Abstract The availability of a fast model feed test for the prediction of catalyst performance in a real feed would be helpful for the development of new catalysts for the deep HDS of diesel fuel. In this work, the relation between the model feed composition and its predictive value for the trend in activity of a CoMo/, NiMo/, and NiW/y-AI203 and PtPd/ASA catalyst for the deep HDS of a pre-hydrotreated gas oil was studied. It appeared that the choice of the sulfur compounds, the HzS concentration and the presence of nitrogen containing compounds strongly affects the catalyst activity ranking. It was found that the ranking and relative activity of the four catalysts for deep HDS of a pre-hydrotreated gas oil could be well predicted by applying a model feed containing 4-ethyl, 6-methyl dibenzothiophene, DMDS and carbazole.
1. INTRODUCTION The development of more active catalysts plays an important role in the realization of improved HDS (hydrodesulfurisation) processes that can meet the new diesel fuel sulfur specifications in an economically feasible way. For an efficient catalyst development, the choice of representative test conditions is crucial. The most realistic catalyst test for deep HDS would be a high pressure, trickle bed experiment with a representative gas oil as a feed. However, such experiments are less suitable for fast catalyst activity screening and it would therefore be convenient to have a fast model feed test, which predicts the catalyst performance in a real feed. Starting point of the design of a model test is the choice of a representative sulfur compound. It is known that DBT's (dibenzothiophenes) and especially 4,6-alkylated DBT's are the most important sulfur compounds in deep HDS [1,2]. In addition, in a real feed various inhibiting compounds like nitrogen- and other sulfur containing compounds and aromatics may influence the catalyst activity. Van Looij et al. [3] have shown the retarding influence of (basic) nitrogen compounds on the deep HDS reactions in a straight run gas oil. On the other hand, Kabe et al. [4] showed that addition of aromatic compounds to a light gas oil did not significantly affect the conversions of various DBT's. In this work, the importance of the choice of the type of sulfur compound and the effect of nitrogen compounds and HzS on the relative activity of a CoMo/-, NiMo/- and NiW/7-AI203 and a PtPd/ASA (Amorphous Silica Alumina) catalyst will be addressed. Moreover, the predictive value of different model feed tests for the activity trends as observed in deep HDS of a prehydrotreated straight run gas oil (P-SRGO) will be discussed.
212
2. E X P E R I M E N T A L SET-UP
2.1. Model feed experiments Experiments with model feeds were carried out in a stirred batch reactor. The catalysts were tested at 633 K and a total pressure of 6.0 MPa. The model feed consisted of a solution of either 0.15 w.% 4-ethyl, 6-methyl DBT (4-E,6-MDBT) or 0.20 w.% DBT (Fluka, >98%) in n-hexadecane (Aldrich, 99%+). The effect of H2S on the catalyst performance was tested by addition of 0.05 w.% (HzS/H2 ratio to 7.10 -3 mol tool-l) of a HzS generating compound (dimethyldisulfide, DMDS, Merck-Schuchart, >99%). To explore the effect of nitrogen containing compounds on the catalyst activity for 4-E,6-MDBT and DBT, 275 ppmw carbazole (Fluka, >98%) was added to the reaction mixture. The mixed oxide catalysts (Co(3.0)Mo(9.5)/, Ni(l.6)Mo(7.9)/ and Ni(1.2)W(15.6)/~,-AI203, loading in w.% metal) were pre-sulfided in a separate reactor in a flowing mixture of 10% H2S in H2 (60 gmol s -~) at 0.1 MPa at 673 K for lh. The Pt(1.0)Pd(1.0)/ASA catalyst was reduced in the separate reactor in flowing H2 (30 gmol s -1) at 0.1 MPa at 573 K for lh.
2.2. Real feed experiments The real feed experiments were carried out in a fixed bed continuous flow micro reactor. The catalysts were pre-treated in-situ, at conditions similar to those described for the model feed experiments. A so-called P-SRGO from Shell Pernis containing 760 ppmw S was used as a representative feed for second stage deep HDS processing [Table 1]. The catalysts were tested at 633 K and a constant total pressure of 5.0 MPa. The hydrogen to oil flow ratio was 22 mol kgoi( I. Table 1: Properties of the applied gas oil feeds
total sulfur total nitrogen total aromatics boiling point range
[ppmw] [ppmw] [mol kg-~] [K]
P-SRGO
light P-SRGO
760 60 1.10 493 - 653
550 5 (+60) 1.39 493 - 593
The gas oil samples were analysed for their sulfur content with a GC (gas chromatograph) (HP 5890 series II), equipped with a 60 m CPSIL-8 CB column (Chrompack) and a Sulfur Chemiluminescence Detector (Sievers, SCD 355). The nitrogen compounds in the P-SRGO were analysed in a GC (Fisons, 8000 series) equipped with a 60 m CPSIL-8 CB column and a Nitrogen Phosphorus Detector (NPD). To study the effect of nitrogen compounds in a real feed, an almost nitrogen free gas oil was created by splitting the P-SRGO into a light- and a heavy fraction by vacuum distillation. The nitrogen concentration of the light P-SRGO was < 5 ppm and the sulfur concentration was increased by spiking with DBT (230 ppmw), 4-MDBT (120 ppmw) and 4-E,6-MDBT (80 ppmw) to a total of 550 ppmw S (Table 1). The effect of nitrogen compounds on the catalyst performance was studied by adding carbazole (60 ppmw N) to the light P-SRGO.
3. RESULTS
3.1. Choice of a representative nitrogen compound The peaks in the GC-NPD spectrum of the P-SRGO mainly represent nitrogen containing compounds although large concentrations of hydrocarbons are also visible as small peaks, especially before 25 min. retention time (Figure 1).
213
I 10
'1
I 20
I
I 30
I
I 40
I
I 50
Retention time [rain]
Figure l" GC-NPD spectrum of the P-SRGO The nitrogen compounds are mainly localised at retention times between 30 and 40 min. The NPD spectrum of the P-SRGO resembles those of gas oil NPD analyses reported in the literature [5,6]. Based on identification of a number of the peaks by Chawla [5] and the position of carbazole in our spectrum it was established that the nitrogen compounds in the P-SRGO mainly consist of C l-alkylated carbazoles. As such alkylated carbazoles were not readily available and because the competitive effect of carbazoles is not strongly influenced by the presence of substituents [7], carbazole was chosen as a representative nitrogen compound for our P-SRGO.
3.2. Model feed experiments with 4-E,6-MDBT The first order overall reaction rate constant for the conversion of 4-E,6-MDBT without additives (blank experiment) for the different catalysts is shown in figure 2A. PtPd/ASA is by far
Figure 2: The first order reaction rate constant for conversion of 4-E,6-MDBT (A: blank, B" DMDS, C: carbazole, D: carbazole, DMDS). the most active catalyst. In figure 2B, the effect of DMDS addition is shown. The first order reaction rate constant of both PtPd/ASA and NiW/7-AI203 decreases strongly as compared to the blank experiment whilst the reaction rate constant of NiMo/7-A1203 seems unaffected. Interestingly, the reaction rate constant of CoMo/7-AI203 significantly increases at the higher HzS concentration. The presence of carbazole significantly decreases the overall reaction rate constant for conversion of 4-E,6-MDBT for all tested catalysts as compared to the blank experiment (Figure 2C). Without DMDS addition, NiMo/7-AI203 is least sensitive for carbazole
214
addition. In the presence of H2S, the strong inhibiting effect of carbazole on the activity of CoMo/T-AI203 is remarkable with a decrease in the reaction rate constant of about a factor 7 (Figure 2D). 3.3. Model feed experiments with DBT In figure 3A, the first order reaction rate constant for the conversion of DBT is shown. Again, PtPd/ASA is by far the most active catalyst.
Figure 3" The first order reaction rate constant for conversion of DBT (A: blank, B" DMDS, C: carbazole, D: carbazole, DMDS). Addition of DMDS leads to a significant increase of the reaction rate constant for DBT over NiMo- and CoMo/y-Al203 as compared to the blank experiment (Figure 3B) whereas NiW/~,-A1203 is only slightly affected. In contrast, the reaction rate constant for DBT over PtPd/ASA is almost halved as was also observed for the conversion of 4-E,6-MDBT. In figure 3C, it is shown that addition of carbazole does not significantly change the ranking of the catalysts. It appears that the conversion of DBT over NiMo/y-Al203 is the least sensitive for carbazole. A similar trend is observed in case of combined carbazole and DMDS addition (Figure 3D). 3.4. Real feed experiments with the light P-SRGO The performance of NiMo/y-Al203 and CoMo/y-Al203 for deep HDS reactions was evaluated in the light P-SRGO which contains less than 5 ppm N (blank experiment). Both catalysts showed a very high conversion of DBT and 4-MDBT. Therefore, only the conversion A
,-.., 90
[] CoMo o NIMo
B
D CoMo o NiMo
~" 90 9 .2
7. 70 g
70 o
D
I
I 5.0
I
I I I I 6.0 7.0 W H S V (g,,~ g~,t -j h -~)
I 8.0
I
I 5.0
I
I I I I 6.0 7.0 W H S V (g,,ll g~,t-~ h-~)
I 8.0
Figure 4: The overall conversion of sulfur compounds in the light P-SRGO (A blank, B spiked with carbazole).
215
of 4-E,6-MDBT will be addressed. It can be observed in Figure 4A that a higher conversion of 4-E,6-MDBT is realised with CoMo/3'-A1203. However, when carbazole is added, the conversion over CoMo/T-A1203 and NiMo/T-A1203 becomes comparable (Figure 4B). 3.5. Real feed experiments with P-SRGO The conversion of sulfur compounds in the P-SRGO was found to be second order in the total sulfur content for all catalysts. This second order rate constant is considered as the activity in second stage deep HDS and is shown in Figure 5. PtPd/ASA is about a factor 2 -3 more active than the sulfided catalysts of which NiW/7-AI203 is the most active - and CoMo/~,-AI203 the least active catalyst.
Figure 5: The second order overall reaction rate constant for conversion of sulfur compounds in the P-SRGO.
4. DISCUSSION
4.1. Activity trends for 4-E,6-MDBT and DBT, the effect of H2S. The role of H~S on sulfided hydrotreating catalysts has been studied extensively in the literature. In general, by adsorption of H2S hydrogenation sites may be converted into hydrogenolysis sites, which can change the rate of elementary reaction steps [8, 9]. At higher H2S concentrations, competitive adsorption of H2S with the reacting species becomes the dominating factor and the reaction rate is suppressed. Figures 2A and 2B illustrate that the activity of CoMo/y-A1203 for 4-E,6-MDBT increases after DMDS addition. Probably because of the low H~S partial pressure in the blank experiment, especially at low conversions, the active phase on CoMo/y-AI203 is irreversibly modified. When DMDS is added the effect of increased competitive adsorption is largely compensated. Also for DBT, the addition of DMDS is beneficial for the activity of CoMo/~-AI203 (Figure 3A). For conversion of DBT, similar effects are observed for NiMo/y-AI203 (Figure 3A and 3B). In contrast, the activity for conversion of 4-E,6-MDBT over NiW/T-A1203 is significantly lower after DMDS addition (Figure 2A and 2B). The relatively high sensitivity of NiW/T-AI203 for H2S has also been reported for hydrogenation reactions [10]. In line with these results, only a small increase in the reaction rate constant for DBT is observed for NiW/T-A1203 with DMDS addition. As may be expected, the reaction rate constant of PtPd/ASA for conversion of 4-E,6-MDBT and DBT is strongly influenced by H2S and is practically halved as compared to the blank experiment. 4.2. Activity trends for 4-E,6-MDBT and DBT, the effect of carbazole It has been well recognised that nitrogen containing compounds can influence HDS reactions by strong competitive adsorption. In general, the extent of inhibition depends on the
216
basicity of the nitrogen compounds and if present, on the position of substituents [l 1]. In this case, carbazole was chosen as a representative nitrogen compound for the P-SRGO. Despite the fact that carbazole is regarded as a non-basic nitrogen compound, it possesses a surprisingly high adsorption constant [ 12]. In our experiments we have observed that addition of carbazole indeed has a strong inhibiting effect on the reaction rate constants of DBT and 4-E,6-MDBT for all catalysts (Figure 2C and 3C) and that carbazole itself was also converted. Still, the conversion of the sulfur compounds could be well described by pseudo first order kinetics up to high levels of conversion of carbazole. When carbazole is almost completely converted after long reaction times, the reaction rate constant for 4-E,6-MDBT recovered to values close to those without carbazole. This observation leads to the conclusion that lhc primary effecl of inhibilion is due to competitive adsorption of carbazole itself and not of its reaction products. The reaction rate constants for the sulfur compounds in the presence of carbazole as shown in the figures has been determined at low carbazole conversion. As compared to the blank experiments, addition of carbazole significantly shifts the trend in activity for 4-E,6-MDBT of the different catalysts in contrast to the trend in activity for DBT. Especially, the difference in sensitivity of CoMo/7-AI203 for carbazole between DBT and 4-E,6-MDBT is remarkable and suggests that the two sulfur compounds react over different sites on CoMo/T-AI203.
4.3. The predictive value of model tests for deep HDS of P-SRGO It is clear that the addition of DMDS and/or carbazole may cause significant shifts in the order of activity of the different catalysts for HDS reactions. Moreover, it was shown that H2S and carbazole not necessarily intervene similarly with the conversion of DBT and 4-E,6-MDBT. These results emphasise that the outcome of a catalyst activity test strongly depends on the choice of feed composition. In Table 2, the predictive value of the various model tests for the catalyst performance in the P-SRGO is shown. The predictive value is expressed as the standard deviation (STD) of the normalised activity in a model test as compared to the normalised activity in the P-SRGO for the four tested catalysts. Obviously, the smaller the STD, the better the prediction. Table 2: The predictive value of model feed tests, expressed as the standard deviation (STD) of the normalised rate constant in a model test as compared to the normalised rate constant in the P-SRGO. Blank
DMDS
Carbazole
DMDS,carbazole
4-E,6-MDBT
0.17
0.15
0.21
0.11
DBT
0.14
0.28
0.08
0.37
The basis of a model test for deep HDS catalysts is the selection of a representative sulfur compound. When we compare the activity trend in the P-SRGO (Figure 5) to the trends found in the blank experiments for DBT and 4-E,6-MDBT (Figure 2A and 3A), it can be seen that both predict the general trend as observed for the P-SRGO rather well. However, in case of DBT the performance of NiW/7-AI203 is underestimated whereas in case of 4-E,6-MDBT the performance of this catalyst is considerably overrated. Based the STD of the blank experiments, no preference for one of the two sulfur containing model compounds can be expressed (Table 2). When DMDS is added to the reaction mixture, the order of activity for DBT conversion does not change. However, the relative activity of NiW/7-AI203 is worse as compared to the blank experiments and the prediction for the activity of PtPd/ASA is much too low which is
217 indicated in a higher STD. The activity trend of the mixed sulfide catalysts for 4-E,6-MDBT does change considerably with addition of DMDS leading to a grave overestimation of the activity CoMo/,/-A1203. Despite this, the predictive value of this test is even somewhat better than the blank experiment since the order of activity and the mutual differences in the absolute value of the reaction rate constant for NiMo/T-AI203 NiW/T-AI203 and PtPd/ASA nicely correspond to the P-SRGO data. The effect of carbazole addition on the activity trend for DBT conversion is limited. However, the predictive value of the model test significantly improves with addition of carbazole as compared to the blank experiment (Table 2). When we draw a comparison with the blank experiment, it is clear that the combined presence of H2S and carbazole considerably worsens the predictive value of the model test for the P-SRGO experiments. In case of the blank experiment with 4-E,6-MDBT (Figure 2A), addition of carbazole overcomes the overestimation of the activity of NiW/3t-A1203 as compared to the blank experiments. However, the relative activity of CoMo/7-AI~O3 is quite low while the activity of PtPd/ASA is rather overrated which together results is a lower predictive value (Table 2). In the presence of H2S, the relative activity of CoMo/T-AI203 for conversion of 4-E,6-MDBT is much too high. The addition of carbazole strongly reduces the activity of CoMo/3t-A1203 in the presence of H2S and strongly improves the predictive value of the model test for the results as obtained in the P-SRGO experiments. Summarising it can be stated that the blank experiments with both DBT and 4-E,6-MDBT result in a rough prediction of the activity trend in P-SRGO. In case of the DBT experiments, addition of carbazole significantly improves the prediction of the P-SRGO results. For the experiments with 4-E,6-MDBT, the addition of H2S alone deteriorates the activity trend. On the contrary, addition of carbazole clearly improves the model test both in the absence and in the presence of H2S. It can be concluded that the addition of carbazole to model feeds with either DBT or 4-E,6-MDBT, leads to a much better prediction of the catalyst performance in P-SRGO. In the presence of H2S, carbazole also considerably improves the predictive value of the model test with 4-E,6-MDBT in, in contrast to the model feed experiment with DBT. 4.4. Evaluation The presented results clearly demonstrate that the composition of a model feed can strongly influence the order of activity of a set of catalysts in deep HDS reactions. Based on these results alone, we can not define a model feed that is generally applicable for an acceptable prediction of catalyst behaviour in other oils. However, we may be able to indicate important factors that should be taken into account to obtain a representative model feed composition for deep HDS of real feeds. From a numerical point of view, the model test with DBT and carbazole gives the best prediction of the catalyst performance in P-SRGO. This is surprising since the applied P-SRGO contains almost exclusively alkylated DBT's which for a significant part have substituents on the 4 and 6 position [13]. Also, the results for deep HDS in the light P-SRGO show that the presence of nitrogen compounds strongly affects the relative activity of CoMo/y-AI203 and NiMo/7-A1203 (Figure 5A and 5B). Without carbazole, CoMo/7-A1203 is significantly more active than NiMo/ 7-A1203, which however is not predicted by the blank or DMDS model feed with DBT. This suggests that the good prediction of the DBT-cabazole model activity feed is merely based on coincidence. On the other hand, the effect of carbazole on CoMo/7-AI203 and NiMo/7-A1203 as observed in the light P-SRGO is well predicted by the 4-E,6-MDBT-DMDS model activity test. Also the prediction of this test of the behaviour of NiW/7-AI203 and PtPd/ASA is acceptable which is indicated by the low STD. Therefore, it is concluded that a combination of 4-E,6-MDBT, DMDS and carbazole results in the most reliable prediction of the performance of catalysts in deep HDS of the applied P-SRGO. In general, the nature of the sulfur compound and
218
the presence of a representative nitrogen compound are the key factors in the composition of a representative model activity test for deep HDS reactions. A systematic approach of model activity tests can provide a better understanding of the parameters that determine the catalyst performance in real feeds. Because straight run oils are very complex and the choice of representative sulfur - and nitrogen compounds will be difficult or even impossible, the described approach is especially successful for pre-hydrotreated feeds.
5. C O N C L U S I O N S It was demonstrated that the presence of H2S and carbazole has a large effect on the catalyst ranking and relative activity of various catalysts for the conversion of both DBT and 4-E,6-MDBT. The effect of H2S and carbazole on the catalyst performance is very different for DBT and 4-E,6-MDBT. Based on these observations it has been possible to formulate a representative model catalyst performance test for second stage deep HDS. The best prediction for the performance of the tested catalysts in the deep HDS of a prehydrotreated gas oil was found for a model feed containing 4-E,6-MDBT, carbazole and H2S. The selection of a representative sulfur - and nitrogen compound is crucial for the outcome of the model test.
6. A C K N O W L E D G E M E N T The research has been performed under auspices of NIOK, the Netherlands Institute for Catalysis Research, Lab Report TUD 99-4-993.
7. R E F E R E N C E S
1. A. Amorelli, Y. D. Amos, C. P. Halsig, J. J. Kosman, R. J. Jonker, M. De Wind, and J. Vrieling, Hydrocarbon Proc., June, 1992, 93. 2. S.S. Shih, S. Mizrahi, L. A. Green, and M. S. Sarli, Ind. Eng. Chem. Res., 31, 1992, 1232. 3. F. Van Looij, P. Van der Laan, W. H. J. Stork, D. J. DiCamillo, and J. Swain, Appl. Catal., 170, 1998, 1. 4. T. Kabe, K. Akamatsu, A. Ishihara, S. Otsuki, M. Godo, Q. Zhang, and W. Qian, Ind. Eng. Chem. Res., 36, 1997, 5146. 5. B. Chawla, J. Chrom. Sci., 35(March), 1997, 97. 6. M. Li, S. R. Larter, D. Stoddart, and M. Bjoroy, Anal. Chem., 64, 1992, 1337. 7. M.V. Landau, Catal.Today, 36, 1997, 393. 8. S.H. Yang and C. N. Satterfield, J. Catal., 81, 1983, 168. 9. L. Vivier, S. Kasztelan, and O. Perot, Bull. Soc. Chim. Belg., 100(11-12), 1991, 801. 10. B. H. Cooper, A. Stanislaus, and P. N. Hannerup, Hydrocarbon Proc., June, 1993, 83. 11. M. J. Girgis and B. C. Gates, Ind. Eng. Chem. Res., 30, 1991,2021. 12. M. Nagai and T. Kabe, J. Catal., 81, 1983,440. 13. H. R. Reinhoudt, R. Troost, S. Van Schalkwijk, A. D. Van Langeveld, S. T. Sie, H. Schulz, D. Chadwick, J. F. Cambra, V. H. J. De Beer, J. A. R. Van Veen, J. L. G. Fierro, and J. A. Moulijn, Stud. Surf. Sci. Catal., 106, 1997, 237.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon, G.F. Froment and P. Grange(Editors) 91999Elsevier ScienceB.V. All rights reserved.
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The influence of zeolite introduction on the HDS activity of CoMo catalysts L. Zanibelli, D. Berti, M. Ferrari, C. Flego, R. Riva EniTecnologie SpA - Via Maritano, 26 - 20097 S. Donato Mil. (MI) - Italy 1. I N T R O D U C T I O N The processing of crude oils containing increasing amounts of heteroatoms (S, N, metals) is growing in importance in order to optimise the exploitation of natural resources. At the same time more restrictive environmental regulations increase the necessity of more efficient hydrotreating technologies for decreasing the heteroatom content in the hydrocarbon fractions. The latest generation of bifunctional catalysts active in reactions such as hydrodesulphurization (HDS) are represented by metal sulfides (MoS2, WS2) and promoters (Co, Ni) supported on inert materials, with the addition of an acidic component (i.e. zeolite), in order to increase the lIDS selectivity [ 1,2]. Aim of this work is to evaluate the influence of the zeolite structure and acidity on the catalytic performances of the so modified CoMo/A1203 catalysts for the HDS reaction. Considering the dimensions of both S-containing compounds and naphtha hydrocarbon fractions, a medium pore (ZSM-5) and two large pore (Y and Beta) zeolites have been chosen. 2. E X P E R I M E N T A L
2.1. Catalyst preparation The catalysts were synthetized according to the sol-gel method [3]. The zeolites are commercial samples (ZSM-5 and Beta from Zeolyst, Y fromUnion Carbide), used in H + form. Co(NO3)E*6H20 was dissolved into CH3(CHE)3OH where also zeolite and aluminium sec-butoxide were added. (NH4)6Mo7OE4*4H20 was dissolved into an aqueous solution. The alcoholic dispersion and the aqueous solution were mixed together till the gel was formed. Then it was aged overnight, vacuum dried and calcined in air at 550~ The catalysts are listed in Table 1, with their chemical composition and the SIO2/A1203 molar ratio of the pure zeolites. Table 1 Ca_talyst chemical analysis. .............Samp..!.e...............C o.(~ CMZ-0 2.3 CMZ-1 2.8 CMZ-2 2.2 CMZ-3 3
.......~o..(~ 9.5 10.5 8.1 10.3
.........Zeo!.!..te ...........Wt.~ ..........S!O~../.A!~O~... . . . . . . ZSM-5 7.0 32.3 Beta 9.0 26.3 Y 6.2 5.1
220
2.2. Sulfidation procedure The calcined samples were sulfided in a 10% n2s/n2 mixture (60 ml/min) at 400~ for lh (heating rate from RT, 6~ [4]; after cooling at RT in sulfiding mixture the sample was cleaned and kept in N2 flow till characterization.
2.3. Characterization techniques Surface area measurements were performed by adsorption/desorption of N2 at -196~ with a Fisons Carlo Erba Sorptomatic 1900; surface area was evaluated with B.E.T. method, pore distribution with Dollymore-Heal model and pore shape according to the De Boer classification. The experimental error in the surface area measures is ca. 10%. XRD (X-Ray Powder Diffraction) data were collected using a Philips equipment with monocromatic radiation Cu Kct (~,=1.5418 A.) in step-scanning mode in the range 40<20<70 ~ (step 0.03 ~ time 10 s/step). Qualitative phase analysis was carried out using the Siemens Diffrac AT software package. TEA/(Transmission Electron Microscopy) experiments were done on a JEOL JEM 3010 microscope (acceleration voltage: 300 kV; max. resolution: 1.6 A) equipped with Link ISIS EDS (Energy Dispersive X-ray Spectroscopy) microanalysis. Adsorption measurements were performed in a pyrex volumetric apparatus equipped with pressure and vacuum detectors. The adsorbed amount of probe (l.tmol/g) was calculated by the general gas law knowing the expansion coefficient of the system. The total number of acid sites was determined by ammonia adsorption at 150~ the NO adsorption capacity (taken as a measure of the Co and Mo sulfides distribution on the surface) was measured at 21 ~ The experimental error of the measure is 3%. XPS (X-ray Photoelectron Spectroscopy) spectra were collected with a VG Escalab MKII spectrometer, using non-monochromatized A1 Ktx radiation. Binding energies were referenced to the A1 2p peak at 74 eV, that is A1 2p in A1203. The binding energies and the peak shapes of Co 2p and Mo 3d allow the oxidation state of the metals to be determined. Atomic ratios were obtained from the ratios of peak areas through appropriate sensitivity factors. The Co/AI and Mo/A1 ratios reflect the dispersion of the metals on the support [5,6,7]. The binding energy of the S 2p peak showed that only sulfides and no sulfates were present in all sulfided samples.
2.4. Catalytic activity Thiophene test allows a fast catalytic screening (3-10 h) at 300~ and atmospheric pressure, with catalyst volume=2 ml. During HDS reaction, H2 flows (2 N1/h) through a bubbling device containing thiophene at 5~ (vapour pressure of 4%). The analyses of the HDS products were carded out by on-line GC-FPD + FID. The product distribution allowed to determine the extent of liDS and hydrogenation activity defined as selectivity to butanes (Sel. C4) in the C4 hydrocarbons derived from thiophene hydrogenation. The proposed [8] scheme for such test is as follow:
221
S
-,s/
v
Model feed test uses a synthetic feed containing olefines (30wt% 1-pentene), sulfur compounds (0.25wt% thiophene) and paraffines (about 70wt%). The operating conditions have been varied in the range: 250~ T <300~ P=10bar; H2/feed = 300N1/1; 4< WHSV <10 h "1. The analysis of the products of the HDS reaction is performed on-line for the gaseous fraction (GC-FPD + FID) and off-line for the liquid products (GC-AED). The conversions for the reactions of hydrodesulfurization (HDS%) and hydrogenation (HYD%) are calculated as follows, where S is derived from organic compounds: %HDS=100*(Sin - Sour)/Sin; %HYD= 100*(n-paraffins Cs)out/(1-pentene)in; 3. RESULTS
3.1. Textural and physico-chemical analyses The main structural characteristics of the zeolites used in this work are reported in Table 2. In Table 3 the N2 physisorption data are collected for the catalysts. No significant influence of the zeolite textural characteristics on the catalysts appears, the surface area slighly increasing only for CMZ-1 and-2. The pore volume reduces from a minimum in CMZ-2 (12%) up to 41% (CMZ-3), compared to CMZ-0. Table 2 Structural characteristics of the zeolites Zeolite Structural features Channel BET area Pore Volume ..................................................................diamete r (A).........(m2/g).....................(cm3/g)_ ............ ZSM-5 medium pore; 10-MR 5.6*5.3 430 0.30 5.1"5.5 Beta large pore; 12- MR 5.5*5.5 590 0.71 7.6*6.4 Y large pore; 12- MR 7.4 890 0.41 MR=at0ms ring
222
Table 3 N222~h_~orption data Catalyst Surface Area
.,
CMZ-0 CMZ-1 CMZ-2 CMZ-3
340 380 390 330
Pore Volume
Average diameter
1.24 1.04 1.09 0.73
60 80 90 80
XRD data show that the only crystalline phases in the catalysts are the zeolites; alumina is amorphous and there is no evidence of crystalline phases containing Co or Mo. The Co and Mo distribution on the support and their interactions with it were investigated by TEM/EDS analysis. Alumina and zeolite are in close contact and the size of their particles is submicronic (between 100 and 700 nm). No Co or Mo clusters are visible by TEM observations. The metal distribution is unhomogeneous: Co and Mo are mostly dispersed on A1203 rather than on the zeolite. From EDS measurements of CMZ-1, in the alumina domain the weight percentage is around 2-3% for Co and between 2 and 16% for Mo, in the zeolite domain the weight percentage is <1% for Co and <2% for Mo. The results of the XPS analysis are summarised in Table 4. As the Mo/AI and Co/A1 ratios do not change significantly with the addition of the zeolites, the overall dispersion of Co and Mo hardly seems to be affected by the presence of the zeolites. Table 4 XPS Catalyst
CMZ-0 CMZ-1 CMZ-2 CMZ-3 2.2wt%Co-A1203 13.8wt%Co-SiO2
Support
AE: Co 2 p - S 2p
A1203 A1203+ZSM5 A1EO3+Beta
A1203+Y A1203 SiO2
617.3 617.5 617.5 617.7 616.5
Co/A1
Mo/A1
0.03 0.03 0.03 0.03 0.02 -
0.05 0.06 0.06 0.06 -
The influence of Mo on the response of Co to sulfidation is evaluated by comparing the Co 2p lineshape of the CoMo catalysts with that of a sample containing only 2.2 Co wt% on alumina. In the absence of Mo, Co becomes resistant to sulfidation even alter a 6 hour treatment. This is supposed to be due to the formation of COA1204, which is resistant to sulfidation [9,10]. The dispersion of Co was lower in this sample than in the CMZ series (Co/A1 ratio 0.02 against 0.03). This means that the chemical interaction between Co and Mo favours the dispersion of Co [11]. The binding energy difference between Co 2p and S 2p of the four catalysts (around 617.5 eV) was found to be higher than that of SiO2 supported Co (616.5 eV). According to Alstrup at al., this points to the presence of CoMoS in the ml203 supported catalysts and Co9S8 in the Co-SiO2 sample [ 12]. The adsorption capacity of NH3 and NO on the sulfided samples are summarised in Table 5. Both the total acid site density and the NO adsorption increase in the zeolite containing catalysts compared to CMZ-0.
223
Table 5 ~ a__pa__pa__pa__p_~of NO and NH3 of the sulfided samples Sample NH3 ads NO ads ..................................................... .(.~.mo!./..g...)..........,(~mo!./..g..)...... CMZ-0 485 54 CMZ-1 588 179 CMZ-2 724 276 CMZ-3 787 125
In order to evaluate the influence of the acidity of the zeolite on the system, NH3 adsorption is performed on the pure zeolites and the results are depicted in Figure 1. The increase of the total acidity of the catalyst can not be ascribed only to the pure addition of the zeolite. In fact a physical mixture of 9.0 wt% of Beta and 91.0 wt% CMZ-0 (composition of the CMZ-2 catalyst) would be expected to adsorb 563 l.tmol/g of NH3 (122 I.tmol/g of the zeolite plus 441 l.tmol/g of the CoMo/A1203 fraction), which is far lower than the measured value (724 lamol/g). Also for CMZ-1 the increase of the acid site density exceeds that expected for the physical mixtures (588 l.tmol/g instead of 533 I.tmol/g). Figure 1 Adsorption capacity ofNH3 (lamol/g) of the zeolites 1600 - NH3 1200 800 400 _
~
!'
ZSM-5
~
1
Beta
~
~
1
Y
Only small amounts of NO are adsorbed on C0-A1203 (36 lamol/g), Mo-ml203 (37 l.tmol/g) and zeolites (ca. 50 l.tmol/g), the nature of the interaction being mostly of physical type. Therefore the addition of the zeolite cannot explain the large increase in NO adsorption found for CMZ-1, -2, -3, whereas the NO uptake in CMZ-0 is in agreement with the CoMoS model [13].
3.2. Catalytic thiophene tests The catalytic results related to the thiophene test (thiophene in H2) are summarized in Table 6.
224
Table 6 Catal ic performances in thiophene test .................... Samples HDS% Sel. C4 % CMZ-0 28.4 18.5 CMZ-1 54.5 24.9 CMZ-2 54.4 28.6 CMZ-3 65.8 32.6
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HDS/SeI.C4 1.5 2.2 1.9 2.0 .
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The introduction of zeolite in the catalyst increases both thiophene HDS activity and the consecutive Ca hydrogenation (SeI.C4) and improves their ratio (HDS/Sel.C4), that means increasing more the HDS activity than the hydrogenation one. 3.3 Catalytic model feed test
Two zeolite containing catalysts (CMZ-1 and CMZ-2) have also been tested with the model feed and compared to CMZ-0. In Figure 2 thiophene HDS and pentene HYD activities are reported. The results confirm the good catalytic performances of the CMZ-1 and -2 catalysts, improving the HDS/HYD ratio. This is due to the increase in HDS versus the pentene hydrogenation (HYD): that means a competition between the two reactions onto the active sites. It is of particular interest to observe the high conversions at lower temperatures than those usually applied in industrial plants [ 14], at the same space velocity and H2/HC ratio. Figure 2 Catalytic results of model feed test (T=280~
WHSV=4.3 h~)
4. DISCUSSION The introduction of the zeolite in the CoMo/A1203 catalysts, prepared by the sol-gel method, does not cause textural changes in the zeolite or in the CoMo/A203. The zeolite is the only crystalline phase present, as confirmed by XRD. Also the dispersion of Co and Mo does not change significantly in the catalysts with the zeolites, because the metals are preferentially deposited on A1203 (from TEM/EDS observations). The textural characteristics of the zeolite do not influence significantly the catalyst, except for the pore volume. By contrast, the adsorption capacities of the catalysts are highly affected by the introduction
225
of the zeolite. Both the total acid site density and the NO adsorption increase in the zeolite containing catalysts compared to CMZ-0. The total acidity (NH3 adsorption) decreases in the order: CMZ3>CMZ2>CMZI>CMZ0. This means that the presence of the zeolite, even in low amount, influences the total number of acid sites of the catalyst. The increase in the acid site density is higher than that calculated for a physical mixture of the components (i.e. it is not a simple additive property). The hypothesis is that during gelification of the alumina matrix, the presence of the zeolite particles prevents the complete formation of the alumina network (AI-O-AI bridges). As a consequence the following calcination gives rise to the formation of "defects" (i.e. A1-OH groups or unsaturated A1 sites), which have a weak acid character, and reduces the pore volume of the catalyst because of the occlusion of a part of the alumina pores by the zeolite particles. Therefore different acid sites are present: zeolite acid sites, defects, CoMo/A1203 acid sites. The zeolite introduction in the system increases also the capacity to chemisorb NO, which follows the trend CMZ2>CMZI>CMZ3>>CMZ0. Due to the presence of the CoMoS phase in the whole CMZ series (as confirmed by XPS analysis) and since NO adsorption gives a measure of the CoMoS phase availability, the presence of the zeolite seems to improve the availability of this active phase. In agreement with the CoMoS model, the formation of atomically dispersed Co sulfide species, anchored on the edge sites of MoS2, is claimed to be the origin of a synergetic generation in the catalytic active site [12]. However it is well known in literature [15], that the system is quite complex and structure and texture of the active site phase are very sensitive to the environment. The presence of the zeolite in the modified CoMo catalysts greatly increases both NO uptake and HDS activity in a simplified catalytic thiophene test, compared to CMZ-0. On the other hand, the Sel.C4 values and acid site density show a similar trend, depicted in Figure 3. As the hydrogenation function is ascribed to the Co and Mo sites, this correlation can be explained taking into account the hydrogen spillover, favoured by the large number of defects (weak acid sites able to adsorb NH3) on the surface. Figure 3 Comparison between Sel.C4 activity (%) (thiophene test) and total acid site density from NH3 adsorption (lamol/g)
In a more complex model feed test only a general trend can be drawn. A halving of the pentene HYD and an increase in the HDS/HYD ratio are observed with the zeolite containing
226
catalysts. In these tests, the composite composition of the feed can cause a greater competition of the reactants in the pores and a lowering of the hydrogen spillover. The following mechanism is proposed: beside the CoMoS sites, able to bond both HE and the unsaturated molecules (thiophene and olefins), new active sites are constituted by the acid sites of the zeolite, able to activate [16] olefins and thiophene, which react with hydrogen supplied by the defects (weak acid sites) on the surface. CONCLUSION The introduction of the zeolite in the CoMo/A1203 catalysts, prepared by the sol-gel method, does not cause textural changes in the zeolite or in CoMo/A1203. The zeolite crystallinity is maintained even after dispersion into the gel matrix. The preservation of the CoMo/A1203 characteristics is due to the preferential deposition of the metals on A1203. By contrast, the introduction of a zeolite in the sol-gel synthesis of a CoMo/AI203 system increases the number of both the CoMoS active sites, without affecting their nature, and the acid sites (including defects, acid sites of the zeolite, CoMo/A1203 acid sites). The increase of the defects on alumina (weak acid sites) favours the hydrogen spillover, while the acid sites of the zeolite may activate the thiophene molecules, creating new active sites beside CoMoS ones.The contribution of these two mechanisms can explain the improved catalytic performances of all zeolites containing catalysts. Acknowledgements The authors are grateful to Dr. F.Bazzano, Mr. V. Arrigoni, Mr. L. Galasso, Mr. A. Monico and Mr. F. Stroppa for their helpful contribution to the experimental work. REFERENCES 1. Y. Okamoto, Cat. Today, 39 (1997) 45-59. 2. T.Becue, J.Leglise, J.M.Manoli, J.Potvin, D.Cornet, J.Catal., 179, 90, (1998). 3. L.Lebihan, C.Mauchauss~, L.Duhamel, J.Grimblot, J. Sol-Gel Sci. Techn., 2 (1994) 837842. 4. R. Iwamoto, J.Grimblot, Stud.Surf.Sci.Cat., 106, 1997, 195-210. 5. J.F. Moulder, W.F. Stickle, P.E. Sobol, K.D. Bomben, Handbook of X-ray Photoelectron Spectroscopy, Perkin Elmer Corporation, Eden Prairie, 1992. 6. D. Briggs, M.P. Seah, Practical Surface Analysis vol. 1, John Wiley & Sons, Chichester, 1990. 7. J.W. Niemantsverdriet, Spectroscopy in Catalysis, VCH Verlagsgesellschaft mbH, Weinheim, 1995. 8. MR.Blake, M.Eyre, R.B.Moyes, P.B.Wells, Bull.Soc.Chim.Belg. 90 (1981) 1293. 9. P. Gayardo, A. Mathieux, P. Grange, B. Delmon, Appl. Catal., 3 (1982) 347-376. 10. R.L. Chin, D.M. Hercules, J. Phys. Chem., 86 (1982) 3079-3089. 11. P. Gayardo, P. Grange, B. Delmon, J. Catal., 63 (1980) 201-216. 12. I. Alstrup, I. Chorkendorff, R. Candia, B.S. Clausen, H. Topsoe, J.Catal., 77 (1982) 397409. 13. H. Topsoe, B.S. Clausen, Catal. Rev. Sci. Eng. 26 (1984) 395-420. 14. H. Topsoe, B.S.Clausen, F. E.Massoth, Cat. Sci. Techn. 11 (1996) 18-22. 15. L. Portela, P. Grange, B. Delmon, Catal. Rev. Sci. Eng. 37 (1995) 699-731. 16. M. Sugioka, C. Tochiyama, Y. Matsumoto, F. Sado, Stud. Surf. Sci. Catal., 94 (1995) 544-551.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
227
H y d r o d e n i t r o g e n a t i o n Properties of S u p p o r t e d Metal Catalysts in the P r e s e n c e of H2S E. Peeters a, J.L. Zotin b, C. Geantet a, M. Breysse c and M. Vrinat a.
aInstitut de Recherches sur la Catalyse, 2 avenue A. Einstein, 69626 Villeurbanne Cedex, France
b CENPES, Petrobras S.A., Cidade Universit~ria, Qd. 7, Ilha do Fund~o,, 21949 900 Rio de Janeiro, R.J., Brazil
cLaboratoire de R~activit~ de Surface, 4 place Jussieu, Tour 54-55, 75252 Paris Cedex 05, France
Abstract The hydrodenitrogenation (HDN) of 1-4 tetrahydroquinoline (THQ) has been performed on various metal supported catalysts in the presence of a small partial pressure of H2S. The nature of the metal and the nature of the support have been investigated. Side reactions and coke formation were also studied. Pt/zirconia catalysts were found to present the best performances for HDN in the presence of H2S.
1. Introduction Conventional NiMo/alumina catalyst has a limited performance for deep denitrogenation of heavier feeds and more severe conditions are required [1-3]. The low activity for aromatic structures, saturation of heterocyclic nitrogen compounds and inhibition effects between pyridine like compounds and aniline structures are considered to be the limiting factors for these catalysts [4-6]. Noble metals are potential active phases for HDN in view of their high activity for hydrogenation of aromatic structures. However, these metals are known to be very sensitive to poisoning by sulfur containing molecules present in petroleum feeds and H2S produced with HDS. The literature concerning the conversion of nitrogen containing molecules over noble metal catalysts is scarce and sulfur free conditions were employed. A screening of the activity of supported noble metals in the hydrogenolysis of methylamine has been performed by Sinfelt [7,8]. Maximum activity was found for Ir and Rh catalysts. Pd, Ru and Pt were slightly less active and presented comparable activities. Van der Eijk et als. [9] studied under trickle
228
flow conditions the HDN of quinoline on various supported group VIII metals. Co, Ru and Rh were found as the most active catalysts but they performed also C-C hydrogenolysis. Pt and Rh were combining the highest activities and selectivities toward HDN products. Therefore, the use of noble metal catalysts for HDN could be envisaged for low sulfur feeds (below 500 ppm, for example) or in a two stage process. In the present work, we reported the performance of metals supported in the HDN of 1,2,3,4-tetrahydroquinoline (14THQ) - a generally accepted model compound for nitrogen-containing oil constituents [10], in the presence of a small partial pressure of H2S. The nature of the metal and the role of the support will be studied.
2. Experimental
Alumina (Condea), silica (Aerosil 200 m2/g), titanium oxide (Degussa P25), yttried-zirconium oxide [11] and magnesia (Merck) were used as support for the catalysts. Yttrium was used as a structure stabilizer for zirconium oxide. Platinum, palladium and ruthenium (-1 wt%) were deposited on the support by incipient wetness impregnation using, respectively, H2PtCI~, RuOC12 and Pd(NH3)n(NO3)2 as metal precursors. The performance of the different noble metals were compared by supporting them on yttried-zirconia. After drying at 393 K, all the catalysts were calcined at 723 K for l h and reduced under hydrogen at 583 K. Nickel and nickelcopper silica supported catalysts were also tested and their preparations were previously reported [12]. A NiMo/alumina (19.5 wt% MOO3, 3.8wt% NiO) industrial catalyst, used as reference catalyst, was sulfided under a 15% H2S/H2 mixture at 673 K. Textural properties of the supports were determined by BET method. The dispersion of the catalysts were characterized by hydrogen chemisorption at 673 K in a volumetric apparatus and also by transmission electronic microscopy (JEOL 2010). Temperature programmed oxidation (TPO) was performed to get information on the deactivation of the catalysts. Used catalysts is heated up to 823 K (5 K.min -1) under Ar/O2 (5%) gas mixture and exhaust gas were analyzed each 30 s by mass spectrometry. Catalytic activities for HDN of I-4-THQ were measured in a dynamic micro-reactor [13] working under a total pressure of 3 MPa, with 4 kPa of l-4-tetrahydroquinoline (THQ), at 623 K and 0,3 kPa of H2S (100 vol. ppm). Reaction products were analyzed by on line gas chromatography. The reaction network for HDN of 14THQ was previously reported [14]. The main reaction products observed were decahydoquinoline (DHQ), o-propyaniline (OPA), propylcyclohexylamine (OPA) and completely denitrogenated products: propylcyclohexane (PCH), propylcyclohexene (PCHE) and propylbenzene (PB). Some quinoline and 5-8 THQ were also observed in equilibrium with 1-4 THQ and were considered as unconverted nitrogen compounds for conversions calculations. Some catalysts also formed heavy products during the reaction, which were identified by GC-MS analysis as substituted molecules of 14THQ, bicyclohexyls and phenylamines, in accordance with the results of van der
229
Eijk et al. [9] and Satterfield et al. [15]. The reaction rates were evaluated according to a pseudo-first order rate equation.
3. R e s u l t s a n d D i s c u s s i o n
Influence of the active phase Nickel on alumina was previously found to be active in the conversion of quinoline in the absence of H2S [9]. In this work, a 20wt% Ni/SiO2 and a NiCu/SiO2 (15 wt% of Ni and 0,8 wt% of Cu) catalyst were investigated. It was reported that Cu replaces the more unsaturated sites of Ni and consequently reduces the number of sites responsible for hydrogenolyis [16]. These catalysts were initially very active and selective towards the HDN products, forming essentially light cracked products, but they strongly deactivated (Fig. 1). In fact deactivation of Ni/SiO2 catalyst was sensibly reduced by the alloying effect of Cu but finally complete deactivation was reached after 15 h of reaction. Chemical analysis and TPO showed that the metal was both poisoned by sulfur and carbonaceous deposit from coke formation. 90
|
0t0
80
-k
70
_.__ k
60
,~ 50 -r -~
!
HDN t
,~ -
\ \
NiCu/SiO 2
40 30
--,,._
20 10
0
200
400
600
800
1000
time (min)
Figure 9Catalytic activity of Ni based catalysts. It turns out that in the presence of a small partial pressure of H2S, a non noble metal based catalyst such as Ni cannot perform the HDN whereas it was an effective catalyst in the absence of hydrogen sulfide [9]. This evidences the necessity of thioresistant catalysts such as noble metals [17]. Then, a screening of the activity of Pt, Pd and Ru supported catalysts has been performed. Ru is expected to present the strongest hydrogenolysis properties, whereas for hydrogenation Pt is the best candidate. These catalysts were prepared over yttried-zirconia because this support gave better HDN selectivities (see below). Chemisorption of hydrogen over these metalhc catalysts supported on yttried zirconia (referred here as ZrY) indicated that platinum and ruthenium catalysts were well dispersed, the palladium catalyst having bigger particles (Table 1.).
230 Table 1 Dispersion of noble metal/Zr-Y catalysts Catalyst Pd/ZrY Pt/ZrY Ru/ZrY
Metal content (wt%) 0.77 0.92 1.09
Dispersion 0.14 0.41 0.36
Particle Size (nm) 8.8 2.3 3.0
The catalytic activities for Pt, Pd and Ru on yttried zirconia are compared in Fig 2. The rate constants were evaluated per surface metal atoms according the hydrogen chemisorption data. Stable conversions were obtained after 12 hours of reaction and the data presented in Fig. 2 were collected after 16 h. Under our conditions, Pt/ZrY is the most active catalyst. Pd and Ru catalysts were 3 and 4 times less active than Pt. As compared to nickel catalysts, noble metal supported yttried-zirconia presented very small formation of light cracked products.
Figure 2 - Intrinsic activity and product distributions of noble metals supported catalysts on yttried-zirconia in the HDN of 1-4 THQ, at 623 K and in presence of 100 ppm hydrogen sulfide. The differences observed in catalytic activities for these metals can be related to their own intrinsic activities for this reaction but also to their sensitivity to sulfur compounds and to the formation of coke or carbonaceous deposits during reaction. This later aspect seems important to evaluate considering the observed formation of heavy products. TPO is a useful technique for this purpose [18] and the coke formation on the catalysts was analyzed after 16h under catalytic conditions. Figure 3 displays the CO2 formed as function of the temperature for the noble metal/yttried zirconia catalysts and the quantitative results are shown in Table 2. The temperature required to oxidize the carbonaceous deposits on the catalysts increases from Pd to Pt. According to the literature, higher oxidation temperature can be
231
related to more condensed coke deposits presenting lower H/C ratios. Therefore, Pt catalyst presented, even at higher conversion level of I-4THQ, the lowest amount of carbonaceous deposit but the most stable one. Pd and Ru catalysts, in opposite, displayed higher amount of carbon deposits which are easier oxidized. So, the lower activity of these metals as compared to Pt can be attributed, at least partially, to the deactivation by coke deposits on the catalyst surface.
Ru
Pd
0 :=k
L)
0
J
300
,
i
,
,
,
400
500
600
700
800
Temperature (K)
Figure 3 - TPO of supported noble metal catalysts after HDN of 1-4 THQ at 623 K
Table 2 TPO data obtained from used noble metal catalysts supported on ZrO2-Y203 Catalyst HDN Conversion at 623 ........... Z (%) Pt/ZrY 96.2 Pd/ZrY 10.4 Ru/Zr-Y 20.8
CO2 formed (~mol.g -1) 1569 4698 5584
Carbon (%) 1.9 5.6 6.7
Combustion Peaks (Z) 660 and 691 405 and 559 584 and 633
Influence of the support. For this part of the study, platinum was selected as active phase in view of previous results. Different oxides were selected in order to vary the acidic and basic properties of the support: MgO, ZrO2-Y203, A1203, TiO2 and H-USY zeolite. The results will be compared with a sulfided commercial NiMo/alumina catalyst. Pt/HUSY catalyst was very active at the beginning of the reaction but it deactivated continuously. For this reason, this catalyst will not be considered in the following analysis. The metal content, dispersion and textural properties of the various Pt catalyst supported on different oxides are presented in Table 3. Pt/TiO2 catalyst was reduced at lower temperature in order to avoid SMSI effect. Dispersions varies from 0.3 for titania supported catalyst to 0,62 for alumina one, the most dispersed.
232
Table 3 Metal content, dispersion and surface area of supported Pt catalysts Catalysts Pt (MOO3) Dispersion Surface Area (wt%) (%) (mZ/g) Pt/ZrY 0.92 0.41 100 Pt/A1203 1.00 0.62 180 Pt/MgO 0.59 0.25 (b). 160 Pt/TiO2 0.41 0.30 (a) 50 NiMo/A1 (19.5) 175 a - catalyst reduced at 473 K, b - determined by TEM. The activities for HDN of 1-4 THQ and the corresponding reaction products yields of these catalysts are shown on Table 4 and compared in Fig. 4. Except for Pt/MgO, which displayed a low activity for this reaction, the other p l a t i n u m catalysts have similar intrinsic activities (rate per g of Pt). Considering the dispersions reported in table 3, the following order of activity was observed: Pt/TiOe > Pt/ZrY > Pt/A12Oa>Pt/MgO. However, the differences in the reaction product distributions are more important. Pt/ZrY was the most selective catalyst for the formation of denitrogenated products while DHQ was the p r e d o m i n a n t product with Pt/AleO3, this is also true with Pt/TiOe catalyst compared with Pt/ZrY at a similar conversion level [19]. SMSI effect on TiOe is expected to s t a r t slightly above the reaction t e m p e r a t u r e [20] and can be considered as limitation for the use of this kind of support. Table 4 Catalytic activities and products yields for the HDN of 1-4 THQ at 623 K Catalyst Conversion Catalytic Activity Product Yields (mol %) (mol %) kcat (a) kpt (b) HDN DHQ OPA+PCHA H.P. (c) Pt/ZrY 45.3 8.3 9.0 21.4 15.1 3.0 5.6 Pt/A1203 44.4 8.7 8.7 9.0 30.1 4.0 1.4 Pt/MgO 5.0 0.9 1.5 0.4 1.7 1.4 1.5 Pt/TiO2 26.7 3.7 9.0 3.5 17.6 2.2 3.4 NiMo/A1 38.8 6.0 14.9 15.8 3.2 4.9 a - specific rate constant (10-4.1.Sl.gcatl); b - m e t a l content based rate constant (10-2.1.s-l.gpt-1) c - H . P . - heavy products As compared to conventional sulfide catalyst, Pt/ZrY is slightly more active but also more selective in HDN products in the presence of a small p a r t i a l pressure of H2S. It is generally accepted t h a t the HDN reaction goes through the i n t e r m e d i a t e s indicated in Scheme 1. Hydrogenation and hydrogenolysis sites are required for the H D N of 14THQ. This reaction scheme obtained on sulfide catalysts can be also used with the metallic catalysts, the same products being observed.
233
Q
14"IHQ
OPA
PCItE
.... 58'IHQ
DHQ
~
1"
1 PCH
PCHA
Scheme 1. Reaction scheme for the transformation of 1-4 tetrahydroquinoline The analysis of the HDN product formation is not straightforward due to the subsequent steps involved. Considering Pt/A1 and Pt/ZrY catalysts, which have similar conversion levels but differ in HDN selectivities, we can see that Pt provides high hydrogenation properties, and this support apparently stimulates the hydrogenolysis step. One possible explanation is that hydrogenolysis is known to be a structure-sensitive reaction [21] and the morphology of Pt particles may depend on the nature of the support [22]. In this case, the mechanism of hydrogenolysis of amines on metals proposed by Sinfelt [7,23] via hydrogen deficient surface intermediates with H2S added as a poison and as a kinetic inhibitor can be suggested. We can also consider the role of the acidity of the support which is supposed to induce thioresistant properties to noble metals in [24] and a bifunctional mechanism may occur. These hypotheses are currently under investigation in our laboratory.
4. C O N C L U S I O N This exploratory work shows that, the presence of small partial pressure of H2S (in the range of 100ppm), it is possible to perform hydrodenitrogenation on metals. However, H2S restricts the use of metallic catalysts to noble metals. Among the noble metal tested, platinum supported catalysts appeared to be the most efficient able to proceed the HDN of 14 THQ with rate constants comparable to those obtained on a NiMo industrial catalyst. Apparently, Pt is the best thioresistant catalyst able to hydrogenate the aromatic cycles and to perform further hydrogenolysis without large coke deposition. Pt on Yttried zirconia support was found to present a high selectivity towards HDN products evidencing a support effect.
5. ACKNOWLEDGMENTS E. P. is grateful to Petrobras Company for providing financial support.
234
6. R E F E R E N C E S
1 2 3 4 5
T.C. Ho - Catal. Rev.- Sci. Eng., 30 (1988) 117. M. Breysse, J.L. Portefaix M. Vrinat, Catalysis Today, 10(4) (1991) 489. F. Luck- Bull. Soc. Chim. Belg., 100(11-12) (1991) 781. G. P~rot - Catal. Today, 10 (1991) 447. C.N. Satterfield e J.F. Cocchetto - Ind. Eng. Chem Process Des. Dev., 20 (1981) 165. 6 H. Toulhoat e R. Kessas, Revue de l'Institut Fran~ais du P~trole, 41 (1986) 511. 7 Sinfelt J H, Catal. Lett., 9 (1991) 159. 8 Sinfelt J H, J. Catal., 98 (1986) 513. 9 J . M . van der Eijk, H.A. Colijn, J. A. R. van Veen Proc. 9th ICC, Ed. Phillips, M.J. and Ternan, M., Calgary, 1 (1988) 50 10 J . R . Katzer, and S. Sivasubramanian, Catal. Rev. Sci. Eng., 20 (1979) 155. 11 D. Hamon M. Vrinat M. Breysse, B. Durand, M. Jebrouni, M. Roubin, P. Magnoux, T. des Couri~res T, Catal. Today, 10(4), (1991) 613. 12 J.A. Dalmon, J. Catal., 60 (1979) 325. 13 M.L. Vrinat, L. de Mourgues, Reac. Kinet. Catal. Lett., 14(4) (1980) 389. 14 C.N. Satterfield, S.H. Yang Ind. Eng. Chem. Process Des. Dev., 23, (1984) 11. 15 C.N. Satterfield M. Modell, R.A. Hites, C.J. Declerck, Ind. Eng. Chem. Process Des. Dev., 17 (2), (1978) 141 16 G.A. Martin, Catal. Rev.- Sci. Eng., 30(4) (1988) 519. 17 C . H . Bartholomew, P. K. Agrawal, J. R. Katzer Advances in Catal. 31 (1982) 135. 18 E. Furimski, F. E. Massoth Catal. Today 17 (1993) 537. 19 E. Peeters, Ph D thesis, Lyon, 1998. 20 A. Dauscher, R. Touroude, G. Maire, J. Kizling, and M. Boutonnet-Kizling, J. Catal. 143 (1993) 155. 21 R.I. Masel, Principle of adsorption and reaction on solid surfaces, John Wiley & Sons, Inc., New York ,1996. 22 P. Briot, P. Gallezot, C. Leclerq, M. Primet, Microsc. Microanal. Microstruct.,1 (1990) 149. 23 G. Meitzner, W. J. Mykytka, J. H. Sinfelt, Catal. Lett. 32 (1995) 335. 24 H.R. Reinhoudt, R. Troost, S. van Schalkwijk, A. D. van Langeveld, S. T. Sie, H. Schultz, D. Chadwick, J. Cambra, V. H. J. de Beer, J. A. R. van Veen, J. L. G. Fierro and J; A. Moulijn, Stud. Surf. Sci. Catal. 106 (1997) 237.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
235
Hydrodesulphurisation and Aromatics Hydrogenation on Straight Run Gas Oils of Maya Crude Oil. Perez, A. A., Marroquin, S. G., Betancourt, R. G., Moreno, T. A. and Aguilar, R.E. Management of Technological Process Development, Mexican Petroleum Institute, Eje Central L/tzaro Cfirdenas No. 152, AP. Postal 14-805, 07730, M6xico, D.F., +52 5 5670975
ABSTRACT Environmentally driven regulations are requiring continuous and significant changes in diesel fuel quality in many parts of the world. Further tightening of specifications are being considered for the future, with trends to lower density, sulphur, and distillation end point, as well as higher cetane number, to meet with more stringent levels of new motor power and environmental regulations. Hydroprocessing is extensively practised in petroleum industry to upgrade fuels refinery streams. In this paper, pilot plant data is presented of a moderate pressure technology developed to produce diesel fuel comply with actual and near future environmental regulations using feedstocks since medium to high contaminants content. This technology is based on CoMo and Ni-Mo ?'-alumina supported proprietary catalysts, with high activity to sulphur and aromatics compounds commonly encountered in these petroleum fractions. Discussion will be driven to show the effect of process conditions on product quality and mains difference in catalysts performance. A series of Isthmus/Maya crude oils mixtures was prepared to test catalyst efficiency, but for the sake of explicitness, only three of these are reported in this work. Hydrodesulphurisation and hydrodenitrogenation activities are quite similar in both catalysts, but aromatic saturation is much more effective in Ni-Mo version. This fact is important when domestic refineries make more intensive use of light cycle oil as a mix component in the feedstock to fulfil demand. 1. INTRODUCTION Chemical processes in refinery operations are classified into those, which reject carbon, and that adding hydrogen, as it is the case of hydrodesulphurisation. Refiners are permanently posed with the need to optimise process operations or develop news technologies, to treat of increasing amounts of heavy crude oil and thus, facing the dilemma of simultaneously satisfying growing demand and also more strict specifications fuels, like diesel. Heavy crude oil is a non-specific term that applies to crude oils with API gravity less than 20, that includes within its definition a wide range of compositions and physical properties. Typically, heavy crude oils have high levels of sulphur, nitrogen, metals (nickel and vanadium), and are also rich in condensed polyaromatic compounds, which react readily to form coke. Many processes used in refineries are catalytic, and are conceived to eliminate such impurities, even though these may reduce and cancel in time, its initial capabilities or refining efficiency [ 1].
236
Environmental regulations governing the composition of diesel fuels are being enacted in the USA and Europe [2]. Some diesel fuels specifications already in use or proposed for the near future are mentioned in Table 1. Table 1 Diesel Specifications Current specification
Auto Oil proposal 2000
European Parliament 2000
European Parliament 2005
California reference
Specific gravity max. 0.860 0.845 0.837 0.830 Sulphur max, wppm 500 350 100 50 Cetane number min 50 51 52 56 Aromatics max, v o l % . . . . Polynuclear max, wt% . . . . 90% distilled max, ~ . . . . 95% distilled max, ~ 370 360 350 340 Reprinted from: Lamboum, Gerald A., Mid-barrel optimisation, PTQ WINTER
500 48 10
Mexican specification
500 48 20
1.4 338 340 1997/98, p.47.
In the context of future specifications, the following priorities are important issues that should be addressed by refiners [3]: Sulphur limits. Needs to desulphurise hitherto acceptable streams such as kerosene; although no sulphur changes for jet fuel are proposed, there could be the need to desulphurise straight run kerosene and light gas oil from low sulphur crude oils which have been to date of acceptable quality. Distillation specification. Need to upgrade high-density streams such as 350-420~ TBP atmospheric gas oils. Unless an upgrading process that affects both boiling range and density is applied, these streams will be excluded from the diesel pool. Density specification. Requires hydrogen addition but reformer by-product hydrogen must be skilfully managed since typical aromatic saturation consumption requires six to ten times more hydrogen than desulphurisation of the same stream. In addition, current gasoline specifications are narrowing the reformer feed boiling range and hence feed rate in order to control benzene content. Aromatics saturation. High aromatics content in diesel fuel lowers the cetane index [4] and are suspected to contribute to undesired emissions in exhaust gases from diesel engines [5]. Thus deep reduction of sulphur and moderate to severe aromatics reduction of middle distillates will be required to reach the new specifications. The key requirements imposed to a process that will treat heavy blends is aromatics saturation and ring opening for density reduction, sulphur and nitrogen reduction, and not least important, an abatement in production of gas and naphtha by-products. The hydrotreating process represents the typical leading approach to fulfil above mentioned considerations; however, existing gas oil hydrotreaters designed to reduce sulphur levels are capable of reducing aromatics content only marginally [6,7]. A high-pressure version is the solution to increase aromatic hydrocarbons saturation [6,8], but few refiners have such high pressure hydrotreaters available and high pressure capacity is extremely expensive to build [9,10]. In this paper a moderate pressure IMP technology, developed and tested at pilot plant scale is presented; this process offers the above mentioned advantages in producing diesel fuel from heavy crude oils straight run gas oils with optimum quality and addressing present and near future environmental regulations. Several catalyst formulations of the type (CoMo/A1203 and NiMo/A1203) were evaluated from which one of a kind was selected for HDS pilot plant tests.
In the case when severe aromatic reduction is needed, revamping of existing units will not be of help, since much higher operating pressures are required, and installation of new processing units is practically unavoidable.
2. E X P E R I M E N T A L 2.1 Feedstock Three samples (mix crudes), with different proportions of Maya crude are prepared and ffactionated to obtain straight run gas oils (SRGO) as feedstock for the study. First one is a typical Maya/Isthmus blend used as feedstock in domestic refineries, the next sample has an increased 10% vol. of Maya crude oil, and the last one is a 100% Maya crude oil sample (table 2). Table 2 Base Crude Oils Properties Property
Specific Gravity 20/4~ Gravity, ~ Sulphur, wt% Total Nitrogen, wppm Basic Nitrogen, wppm C, insl., wt% Ramsbottom Carbon, wt%. Metals Fe/Cu, w% Ni/V, pprn
Crude 1 Base*
0.8797 28.9 2.31 2080 512 7.6 6.5 1.28/0.34 24/129
Crude 2 Base+10% vol. Maya crude
0.8859 27.8 2.52 2268 536 8.6 7.2
1.24/0.38 28/152
Crude 3 100% Maya crude
0.9231 21.3 3.8 3400 685 14.8 11.7
1.0/0.6 51/286 *Base - Maya/Isthmus 30/70
From this data, it is worth to notice that there are important changes in some key properties of Maya/Isthmus blends as compared with those of Maya 100%. Basic nitrogen increases about 28 to 34%, when compared with pure Maya sample, whereas asphaltenes (C7 insl.) content is 72 to 95w% higher and well correlated with corresponding vanadium content. Ramsbottom carbon, which is also a good measure of feedstock quality, is 62 to 80% higher. The importance of these properties is reflected in the composition of derived SRGO, that exhibit increasing amounts of compounds, that makes them more difficult to hydrotreat under conventional conditions. Thus, essentially long and short boiling range gas oil samples, with particularly quite different sulphur, nitrogen and aromatic content were obtained; their properties are reported in table 3.
238
Table 3 Straight Run Gas Oil Properties Property
Specific Gravity, 20/4~ Gravity, ~ Sulphur, wt% Total Nitrogen,, wppm Basic Nitrogen, wppm ASTM distillation, ~ IBP / 10% 50%/90% Cetane Index Total aromatics, vol% Mono-aromatics Di-aromatics Tri-aromatics Saturates Yields on crude feed, vol%
Crude 1 SRGOL
Crude 2 SRGOL
Crude 3 SRGOL
164 / 221 286 / 347 44 37.1 23.2 11.7 2.2 62.9 21.0
168 / 224 290 / 351 40 42.5 26.6 13.2 2.7 57.5 20.0
241 / 277 299 / 352 34 49.1 29.1 16.7 3.3 50.9 17.0
0.8428 35.80 1.19 226 66
0.8465 35.29 1.47 251 81
0.8687 30.85 3.16 400 182
2.2 Experimental Procedure Test runs were carried out in an isothermal 250 mL catalytic bed pilot plant unit, equipped with temperature controllers that operate on heating devices located along the reactor to provide the needed heat to start-up the hydrodesulphurisation, hydrodenitrogenation and saturation reactions. Likewise, the reaction system has a separation section that recovers the liquid product and eliminates the produced bitter gases. The non-stabilised products are sent to the light ends recovery plant for separation of fuel gas, LPG, naphtha and diesel products. The first stage in pilot plant test experimentation is insitu catalyst presulphiding. The reactor is then charged with reformated naphtha containing l wt% of carbon disulphide at LHSV = 4-5h a hydrogen flow rate of 4500 mL/min, 5.3 MPa and 500 K. Straight run gas oils obtained from crudes 1-3 are hydrotreated at LHSV from 1 to 2 hr -a, and the hydrogen flow rate was 1500 mL/min (STP), except for Crude 3 gas oil for which 3500 mL/min was used, 5.0-8.0 Mpa and 640 to 660 K. In each case, hydrotreated SRGOL samples were analysed for sulphur content by X-ray fluorescence spectrometry (ASTM-D-4294), nitrogen content (ASTM-D3279); mono-, di- and tri-aromatics as measured by HPLC (ASTMD-5186). 3. R E S U L T S AND DISCUSSION Industrial SRGOL stocks are hydrotreated in units primarily designed for conventional hydrodesulphurisation. In order to study the level of cetane improvement (aromatic saturation) that can be typically achieved in such unit, pilot plant test was conducted using three SRGOL with different Maya crude contributions. As was mentioned hydrotreaters are designed to process straight run gas oils containing from 1.0 to 1.5 wt% of sulphur, offering end products with 0.05 wt% S. Typical and heavier feedstock were studied with sulphur levels up to 3.16 wt% (crude 3 sample), and with both CoMo and Ni-Mo catalyst. Under previously mentioned reaction conditions, it was possible to obtain products with less than 0.05 wt% sulphur.
239
The first obvious benefits of gas oil hydrotreatment should be pictured as a function of sulphur and nitrogen reduction; results show that both catalysts developed a very good hydrodesulphurisation and hydrodenitrogenation activity under the tested conditions, and exhibit over 98% efficiency in sulphur as well as in nitrogen compounds elimination. Nevertheless, Co-Mo catalyst is, in the range of 340 to 360~ slightly better in sulphur removal, as Ni-Mo catalyst is in nitrogen removal, which could be evidence of a better renovation of acidic sites formed in the support of Ni-Mo than in former Co-Mo formulation. Given the very high initial sulphur and nitrogen content of the samples, particularly of Maya crude SRGO, both catalysts exhibit an outstanding performance for the above reactions. Table 4 Desulphurisation and Denitrogenation Efficiency LHSV = 1.0 340 ~
Type of SRGO ~ Sulphur, wppm Total Nitrogen, wppm Basic Nitrogen,wppm Cetane Index 350~
Type of SRGO ~ Sulphur, wppm Total Nitrogen, wppm Basic Nitrogen,wppm Cetane Index
Crude 1 42.1 79 5 2 48 Crude 1 42.3 47 3 1 48.3
360~
Type of SRGO ~ Sulphur, wppm Total Nitrogen, wppm Basic Nitrogen,wppm Cetane Index
Co-Mo/A1203 Crude 2 40.0 197 13 5 51.6 Co-Mo/A1203
Crude 2 40.1 152 11 3 52.3
Crude 3 39.0 350 19 8 52.8
Crude 1 42.7 140 3 1.2 48.3
Crude 3 39.1 308 16 5 53.5
Crude 1 42.8 110 2 0.8 48.6
Crude 3 39.1 265 14 4 54.0
Crude 1 42.9 70 1 0.6 49.0
Co-Mo/A1203
Crude 1 42.3 24 2 0.5 48.5
Crude 2 40.2 101 8 2 53.1
hr-1; 8.0 Mpa
Ni-Mo/A1203 Crude 2 40.2 240 9 3 51.9 Ni-Mo/A1203
Crude 3 39.1 375 13 5 53.2
Crude 2 40.2 190 7 2 52.6
Crude 3 39.2 310 11 3 53.9
Ni-Mo/A1203 Crude 2 40.6 150 5 1.8 53.4
Crude 3 39.4 280 9 2.5 54.7
The cetane quality of a diesel fuel is a function of its aromaticity [ 11]. Thus, to gain insight into how the cetane index (ASTM-D-976) can be improved, an understanding of how aromatic hydrocarbons are converted under different hydrotreating reaction conditions is essential. Comparing original aromatic distribution from table 3 with that obtained from Co-Mo hydrotreating at different reaction conditions (table 5), it is confirmed that tri-aromatics are first converted to simpler di-aromatic compounds, and these consecutively to mono-aromatics, by ring saturating and post- partial cracking. Table 5 Saturation Efficiency 360~
LHSV= 2.0
Pressure, MPa Type of SRGO Total aromatics, vol.% Mono-aromatics, vol.% Di-aromatics, vol.% Tri-aromatics, vol.% Saturates, vol.%
5.0 Crude 1 28.66 25.46 3.20 0.00 71.34
Co-Mo/AI203
6.0 Crude2 29.30 26.00 3.30 0.00 70.70
6.5 Crude3 33.82 30.25 3.57 0.00 66.18
5.0 Crude 1 22.00 21.06 0.94 0.00 88.00
Ni-Mo/AI203 6.0 Crude 2 23.30 22.04 1.26 0.00 76.70
6.5 Crude 3 30.90 28.45 2.45 0.00 69.10
240
A more localised mesomeric effect in mono-aromatic ring structures represents a huge obstacle to complexion of total aromatic saturation under studied conditions. As a result of this, triaromatic rings are completely converted and most of di-aromatics too, while mono-benzene compounds are present in a slightly higher concentration than in the feed when used Co-Mo. As it is expected for a consecutive reaction path mono-aromatics content first increases, goes through a maximum, and subsequently decreases as the severity of hydrogenation further increases. The initial increase of the mono-aromatics is caused by conversion of di-aromatics. At low hydrotreatment severity no saturation of the mono-aromatics occurs, and only at higher severity can the mono-aromatics be saturated. The overall benefit is represented between ten to thirty percent more saturates in the product, regardless of the reaction temperature. In Ni-Mo catalyst the trend is similar but a deeper de-aromatisation is observable; it is clear that this catalyst is by far a better option when cetane improvement is also a requirement, as it is the case of mixed SRGO and light cycle oil feeds. In all cases, there is a positive effect of pressure on aromatic hydrocarbons conversion (figure 1).
Figure 1.- FIA vs. Temperature of operation When aromatics increase then cetane index decreases. At around 360~ the minimum aromatics content is reached, resulting in a cetane index improvement of approximately from 5 to 20 points for the Ni-Mo catalyst. The activity of Ni-Mo catalyst was higher than Co-Mo catalyst for aromatics saturation under all conditions tested. Diesel Motor tests carried out at IMP labs (ASTM-D-613) and before additivation, confirm a difference of five to ten units above the cetane index, for corresponding cetane number. The hydrotreated product at 360~ from the Maya 100%, Crude 3, exhibits the best cetane number, and best lubricity, with 0.48mm wear scar diameter value. However, its oxidative corrosion test renders 10-fold formation of mud than corresponding Crude 1 and 2 samples of SRGOL.
241
4. CONCLUSIONS The challenge to work with heavy crude oils in petroleum refinery requires the development of new more active and more selective catalysts. New process outlines should also be developed that allow increasing throughput of high value products, in the light of increasing demand for environmental protection. In Crude 1 and Crude 2 SRGO, sulphur and nitrogen contents can easily be reduced to very low values at moderate hydrotreating conditions using either high activity Co-Mo or Ni-Mo catalyst. Sulphur and nitrogen contents in all products, under all tested conditions were below 500 ppmwt and 100 ppmwt for sulphur and nitrogen respectively, in both catalysts. Aromatics reduction to low levels (<30 %vol) by hydrotreating is much more challenging than sulphur removal. Saturation of polyaromatic feeds (> 10 %vol) to mono-aromatics (products with more the 20 %vol) is relatively easy. The difficulty arrives when saturation of very stable monoaromatic ring is needed, since this last stage of saturation is very slow and more energy demanding. As a result, mono-aromatics tend to accumulate in product under studied conditions, as it is evident from Crude 1 SRGO, having in feed 23.2 %vol (arom.) and in product 24.4 %vol. Higher severity conditions especially greater hydrogen partial pressures and high contact times are required to reduce aromatics in distillates to low levels. Overall results show that with Crude 3 heavy feedstock, cetane number and lubricity test improvement is higher that in lighter feedstocks. Conditions required for deep hydrogenation of distillates will also result in almost total sulphur removal. Therefore, with a high activity CoMo and NiMo or mix-based catalyst of the kind it is technically possible to achieve a high degree of aromatics saturation (< 30 %vol) and meet the 0.05 wt% sulphur specification.
REFERENCES
[1] [2] [3] [4] [51 [6] [7]
[s] [9] [ 10] [ 11]
[12]
Lee, S.L., Jonker, R., Akzo Catalysts Symposium 1991, (Hydroprocessing), pp 25-46. Crow, P. and Williams, B., Oil and Gas J., Jan. 23, 1989, p. 15. Lambourn, G.A., Mid-barrel optimisation, PTQ WINTER 1997/98, p 48. Suchanek, A.J., Oil and Gas J., May 7, 1990, p. 109. Kasztelan, S., Marchal, N., and Kressmann, S., Production of environmentally friendly middle distillates by deep sulphur and aromatics reduction, Published by John Wiley & Sons. Johnson, A. D., Oil and Gas J., May 30, 1983, p.79. McCulloch, D.C., Edgar, M.D. and Pistorious, J.T., Paper AM-87-58, NPRA, March 1987. Cooper, B. H., Stanislaus, A. and Hannerup, P.N., Preprint, Fuel Chem. Div., ACS, p.41, vol.37, 1992. Nash, R.M., Oil and Gas J., May 29, 1989, p.47. Asim, M.Y., et to the one., Paper AM-90-10, NPRA, March 1990. Ullman, T.L., Spreen, K.B., and Mason, R.L., "Effects of Cetane Number, Cetane Improver, Aromatics, and Oxygenates on 1994 Heavy Diesel Duty Engine Emissions", SAE Paper 941020 presented at the International Congress & Exposition, Detroit, Michigan, February 28-March 3, 1994. Eastwood, D., They go of Venne, H., "Strategies for revamping distillate desulfurizers to meet lower sulfur specifications", AM-90-20, NPRA, March 1990, San Antonio, TX.
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
243
Hydrocracking of vacuum gas oil on CoMo/alumina (or silicaalumina) containing zeolite Woo-Suk ChoP, Kyong-Hwan Lee b, Kyungil Choi c and Baik-Hyon Ha School of Chemical Engineering, Hanyang University, 17 Haengdang-dong, Sungdong-Ku, Seoul 133-791 aplant Division, Daewoo Engineering Company, P. O. Box 20, Bundang, Korea bEnvironment Remediation Research Center, KIST, Seoul 136-791 cSK Corporation, 140-1, Wonchondong, Yusunggu, Daejon, Korea Abstract Precipitated alumina (or silica-alumina) was mixed with zeolite Y (or mordenite), and then cobalt and molybdenum oxides were impregnated in these substrates. These laboratory-prepared catalysts and some commercial hydroprocessing catalysts for heavy oil were characterized by measuring specific surface area, pore volume and pore size distribution. Hydrocracking activities of vacuum gas oil (VGO) over these catalysts were examined in a batch reactor. The catalysts mixed with silica-alumina have larger meso- and macro-pore size and smaller surface area than those mixed with alumina. Zeolite-Y in the substrate enhances the yields of both 350~ and naphtha. The increase of micropore surface fraction of the catalysts enhances the BTX formation whether the zeolite exists or not. The silica-alumina-based catalysts showed lower coke formation than the alumina-based ones. The BTX are thought to be formed in the micropore from n-paraffin produced by VGO cracking process.
1. I N T R O D U C T I O N Hydroprocessing of heavy oil is becoming more important in refineries both for profit maximization and for stricter environmental regulation, and especially naphtha production from the heavy oil fraction is an interesting aspect [1,2]. Naphtha formation through the hydrocracking process could be controlled by not only reaction conditions but also an adequate design of the catalyst. The catalyst contains alumina, clay or silica-alumina as substrate and also dealuminated hydrogen zeolite-Y such as USY or HZSM-5 as a single or a mixed state. The commercial alumina or silica-alumina has a wide mesopore distribution (3-100
244
nm) [3]. Arahata et al. [4] reported that the mesoporous (13-18 nm) and high surface area of the catalyst is more effective for the heavy oil cracking and HDS than that of its macropore. However, Itoh et al. reported that the macro-pore volume is rather important to suppress coke formation over the AUophane [5]. It was also reported that, both in the presence and in the absence of CoMo/alumina catalyst, cracked product distribution from VR was similar [6], and it didn't depend on acid strength [7]. However, the hydrocracking over alumina or silicaalumina with zeolites shows different product distribution from previous cases of non-zeolitic catalyst, probably because the primarily formed hydrocarbon over the substrate can be further cracked inside the micropore of the zeolite. The aim of this study is to understand the roles of the pore characteristics of the substrates (alumina and silica-alumina) and zeolites (Y and mordenite) existing as the mixed phase, for the product distribution of hydrocracking. The formation of BTX was also discussed based upon the surface fraction of micropore of the catalysts. 2. E X P E R I M E N T A L 2.1. C a t a l y s t s 2.1.1. C o m m e r c i a l c a t a l y s t s The commercial catalysts such as CoMo/A(C-444), CoMoP/A(C-344) NiMo/A(shel1317) and NiW/SA-Y(C) were used to compare with the catalysts prepared in the laboratory, in which A, SA, Y and C indicate alumina, silicaalumina, zeohte-Y and commercial, respectively. First three commercial catalysts were adopted because of their large pore structure even though they are hydrotreating catalysts. 2.1.2. P r e p a r a t i o n of c a t a l y s t s 2.1.2.1. D e a l u m i n a t i o n of z e o l i t e s For the hydrothermal stability as well as for an enlargement of the pore size of the zeolite, the zeolites were dealuminated by wet-air-treatment/acid-washing, which forms secondary pore [8]. NaM (Zeolon-900, Norton) and zeohte-Y (LZY-50, Union Carbide) were ion-exchanged with 1 N NH4C1 solution at 80~ for 18 hours and washed with distilled water. This procedure was repeated three times to enhance the ion-exchange rate. The washed samples were treated with wet air at 400~ for 6 hours (95 % steam) at a flow rate of 100 cc/min and washed with 0.01 N HC1 solution to remove the non-framework aluminum species. 2.1.2.2. Zeolite + s i l i c a - a l u m i n a Silica to alumina mole ratio was controlled at 80:20 in order to provide maximum acid sites. 1 N A1C13solution was dropped into a diluted water-glass (10 times) solution to be a slurry phase of precipitated aluminum hydroxide, and then dealuminated zeolite was added for mixing two kinds of solid. This slurry
245
was aged for 12 hours at pH 3.8, and then neutralized with 0.01 N NHnOH. The washed and filtered slurry was dried at 105~ and then calcined at 550~ for 2 hours. Cobalt and molybdenum-nitrate were impregnated on the solids to be 5 wt.-% and 1 wt.-%, respectively. The impregnated catalysts were dried at 105~ and then calcined at 500~ for 3 hours. These catalysts were designated as CoMo/SA-Y and CoMo/SA-M, in which M stands for mordenite.
2.1.2.3. Zeolites + a l u m i n a Aluminum chloride (Junsei first class) was dissolved in distilled water at 80~ and then aluminum hydroxide slurry was obtained by dropping of ammonia water. Zeolite was added to the slurry solution at pH 9 and aged for 12 hours. The washed and dried samples were treated with the same procedure as the silica-alumina. These catalysts were denoted as CoMo/A-Y(50), in which zeolite content was shown in the parenthesis as %. 2.2. N i t r o g e n a d s o r p t i o n The pore and surface area of the catalyst were characterized by nitrogen adsorption at liquid nitrogen temperature using Micromeritics ASAP-2000. 2.3. F e e d Hydrotreated VGO from SK Co., the composition of which is paraffin (10-15 vol.-%), naphthene (40-50 vol.-%) and aromatics (35-50 vol.-%), was used for the hydrocracking reaction. Distillation range of the feed was 350-550~ Feed properties of VGO are as follows: API: 27.9, Sp.Gr.: 0.888, sulfur: 0.15 wt.-%, nitrogen: 60 ppm. 2.4. R e a c t i o n a p p a r a t u s and p r o c e d u r e A batch type reactor (ID: 60 mm, volume: 480 ml) of high pressure stainless steel was used for the reaction. The reactor with stainless ball (12 mm~) inside was swayed at the rate of 36 cycles/min for good mixing. The reactant was a mixture of VGO 50 g and catalyst 5 g (presulfided) and the reaction was carried out at 400~ and 160 kg/cm 2of H2 pressure. Prior to the reaction the reactor was purged 3 times with hydrogen at 80 kg/cm 2. The temperature was elevated up to 400~ with 3.3~ of increasing rate, and maintained for 5 hours, and cooled down to the ambient temperature. The product mixture was separated into gas, liquid and solid phases. The liquid phase was fractionated up to 350~ and the distillate was analyzed by gas chromatography (Shimadzu GC-14A) equipped with capillary column and FID detector. The amount of coke formed was measured by the weight loss of the catalyst after the soluble species was extracted by n-heptane. The total conversion was determined based upon the total product amount of coke, gas and 350~ The distillate consisted of n-parofl~ns, olefins, naphthenes and aromatics, which were identified with GC/MS (Shimadzu GC/MS-QP2000A).
246
3. R E S U L T S A N D D I S C U S S I O N 3.1. S p e c i f i c s u r f a c e a r e a s and p o r e s t r u c t u r e s The specific surface area, pore volume and the average pore diameter of the catalysts measured by nitrogen adsorption are shown in Table 1. Mesopore size distributions of the catalysts are also shown in Figure 1. The pore size distributions of commercial catalysts CoMo/A(C-344), CoMo/A(Shell317) and NiW/SA-Y(C) are similar in spite of the existence of molecular sieve in NiW/SAY(C). For the catalysts prepared in the laboratory, CoMo/SA-Y has a wide range of meso- and macro-pore size. However, CoMo/A-Y has small sharp mesopores which were formed as secondary pores in zeolite when dealuminated. CoMo/AY(25) has a similar mesopore distribution to the commercial ones, but has smaller pore volumes than the latter, especially small for CoMo/A-Y(33 and 50). 3.2. C a t a l y t i c r e a c t i o n s 3.2.1. A c t i v i t y and s e l e c t i v i t i e s to coke, gas and liquid p h a s e Table 2 shows hydrocracking activity of VGO and product distribution for both commercial and laboratory-prepared catalysts. Table 1 BET surface areas, pore volumes and the average pore diameters of various catalysts surface area(m2/g) pore volume (cc/g) average pore Catalysts BET Micro total Micro diameter(nm) CoMoP/A(C-344) 202 12 0.52 0.003 10.4 NiMo/A(Shell317) 170 10 0.44 0.003 10.3 NiW/SA-Y(C) 306 32 0.41 0.012 5.4 CoMo/SA 337 0.41 4.9 DY 458 379 0.32 0.17 2.2 CoMo/SA-Y(25) 153 76 0.50 0.03 13.0 CoMo/SA-Y(33) 193 100 0.52 0.04 10.0 CoMo/SA-Y(50) 200 134 0.38 0.06 6.3 CoMo/A-Y(25) 313 132 0.43 0.06 4.1 CoMo/A-Y(33) 258 103 0.36 0.06 4.0 CoMo/A-Y(50) 286 143 0.4 0.07 3.7 DM 498 0.26 1.7 CoMo/SA-M(33) 149 89 0.41 0.04 8.9 CoMo/SA-M(50) 176 117 0.42 0.05 7.8 DY and DM: dealuminated zeolite-Y and mordenite. CoMoP/A(C-344)[(Co(2.52 wt.-%)Mo(9.85 wt.-%)/alumina], NiMo/A[Mo(12 wt.-%) Ni(3 wt.-%)/alumina] and NiW/SA-Y(C)[W(5 wt.-%)Ni(5 wt.-%)/silica-alumina]
247
.Z r| ~2 rr
//CoMo/A-Y(33) ?
~D
NiW~SA~(C)i!
9
/
9
f ~CoMoP/A(C-344)
, !~
I ;i',,,'
/ I /
t t/ ~i t7 7 {
~ t~
z
/ ,,;\ it! ,,
"
/
,: ,CoMolSA-Ye.5),-~"/~
';,,,,. ',
~., ,r-.--Ak" ~,.. ';'-~'- : r . . . . . . ~
"-- ,,,. ~
--
.. ,,_._~,..,,.
%:-',--r',
10
Pore diameter /nm
(A)
i![li~ i~ CoMo/A-Y(50) CoMolSA-Y(33) . ~/CoMo/A-Y(25) CoMo/SA-Y(50) \ ./
Nilo/l(Shell317)
, ,, 100
1
10
100
Pore diameter / nm
O3)
Figure 1. Pore size distributions of commercial (A) and laboratory-made (B) catalysts. Table 2 Coke, gas, distilled fraction (<350~ and conversion obtained from hydrocracking of vacuum gas oil on various catalysts Catalyst Coke Gas Distilled fraction (<350~ Conversion wt.-% wt.-% wt.-%(Naphtha, Kerosene, Diesel) wt.-% Thermal 2.3 4.1 30. (24.8 26.8 48.4) 36.4 CoMoP/A(C-344) 0.3 3.8 31.3 (30.5 27.4 42.1) 35.4 NiMo/A(Shell317) 4.2 1.8 36.0 (30.0 27.5 42.5) 42.0 NiW/SA-Y(C) 4.0 5.9 49.1 (40.0 25.2 35.2) 59.0 CoMo/A 0.7 3.4 22.3 (17.3 37.3 45.4) 26.4 CoMo/SA 1.2 6.6 56.2 (50.8 28.6 20.6) 64.0 CoMo/DM 10.7 2.2 22.3 (42.3 35.2 36.1) 35.2 CoMo/SA-M(33) 5.0 2.9 33.1 (34.2 32.3 33.6) 41.0 DY 6.2 2.9 29.3 (52.2 27.9 19.9) 38.4 CoMo/A-Y(25) 0.4 38.1 61.1 (67.3 32.7 - ) 99.6 CoMo/A-Y(33) 2.1 10.4 72.0 (72.6 23.7 3.7) 84.5 CoMo/A-Y(50) 1.3 52.2 46.4 (82.6 19.4 4.3) 99.8 CoMo/SA-Y(25) 0.2 11.5 73.6 (66.0 26.6 7.4) 85.3 CoMo/SA-Y(33) 0.5 16.0 76.5 (70.3 23.3 6.4) 93.0 CoMo/SA:y(50) 0.3 28.0 67.0 (80.7 16.6 2.6) 95.3
248
Alumina-based commercial catalysts (C-344 and SheU317) have about 35-40 wt.-% of conversion, similar to the thermally cracked one, whereas molecular sieve based NiW/SA-Y(C) has about 60 wt.-%, which is higher than that of the former. This enhanced conversion on NiW/SA-Y(C) may result from the properties of Ni and W [7] as well as further cracking in the zeolite. For the laboratory-prepared catalysts containing dealuminated zeolite-Y, both silicaalumina-based CoMo/SA-Y and alumina-based CoMo/A-Y show higher conversion despite the different pore size distribution. In other words, the addition of zeoliteY to the alumina or the silica-alumina drastically enhances the conversion and naphtha yield. It means that the zeolite further cracked the hydrocarbons that were formed on the substrate due to either thermal or catalytic reaction. However, the zeolites (CoMo/DM and DY) without a substrate show just about 35-38 % conversion regardless of the zeolite type, which is the same level as the thermal cracking. This suggests that the high activities appeared on CoMo/SAY(or A-Y) resulted from the first tailoring of VGO on the substrates. The mordenite doesn't play the same role as zeolite-Y due to its strong coking tendency as shown in Table 2. Especially, the three-dimensional channel and larger pore size of zeolite-Y, which has a secondary mesopore of 4 nm as shown in Figure 1, enhance not only the conversion (84 or 93 wt.-%) but also gas yield. The over-cracked yield to gas on CoMo/A-Y may be due to the large and sharp secondary pore appearing at about 4 nm of zeolite-Y. High naphtha yield was observed for zeolite-Y containing catalysts, whereas yield to kerosene and diesel was favored for commercial catalysts. This also explains that cracking to light hydrocarbons occurs in the micropores of zeolite. The low yield of naphtha on HM and SA-M can be explained by the characteristics of high coking property and the one-dimensional channel of the mordenite. As shown in Table 2, the amount of coke formed on the CoMo/DM is high. It is interesting to note that the coke content on CoMo/SA-Y is slightly lower than that on the CoMo/A-Y, which may be due to the large pore size of the silica-alumina [5].
3.2.2. Naphtha yields with zeolite c o n t e n t Yields of naphtha and 350~ with respect to zeolite-Y content over the various catalysts are plotted in Figure 2. The figure shows that naphtha yield increases up to 30-40% of zeolite content and then decreases afterwards regardless of the type of substrate. It implies that the micropore of the zeolite in the substrate plays an important role for the further cracking. 3.2.3. BTX and n-paraffin yields in distilled fraction The BTX and n-paraffin yields were correlated with the ratio of micropore surface area to total surface area of the catalysts as shown in Figure 3. In order to provide various pore structures for CoMo/A-Y(33) catalyst, special modification treatment of wet-air-treatment at high temperature followed by acid extraction was carried out.
249
sot
)6o
o
350 ~
-6V__
naphtha
401~5/~1 20
0
10
I
I
,
rq
I'
i
20 30 40 Zeolite c o n t e n t / w t . - %
50
Figure 2. 350~ (circle) and naphtha yields (square) with the zeohte-Y content in the alumina (open symbols) and silica-alumina (closed symbols). 25 I
50
O
20
"o
>'15
x I-rn C
m
33
?0
10
.___
O
a_ 5 I
c-"
I
I
0.2
I
I
I
0.4
I
0.6
Micro--surface area / total surface area OQ CoMo/A--Y(33) ,/~ ~L, CoMo/A--Y 9 CoMo/A(Shel1317) m NiW/SA(C)
~
,0, CoMo/SA--Y
Figure 3. n-Paraffin and BTX yields against the ratio of micro-surface-area/total surface area of the catalysts (open symbols: n-paraffin, dosed symbols: BTX)
250 As a result, six different micropore surface area/total surface area ratio samples could be obtained out of CoMo/A-Y(33) catalyst as in Figure 3. Two commercial catalysts, non-zeolitic C-344 and slightly zeolitic NiW/SA-Y(C), were also adopted for correlation. The BTX yield proportionally increases as the micropore surface fraction increases whether the zeolite-Y exists or not, whereas the yield of n-paraffin decreases when the ratio of micropore surface area to total surface area goes up. It is noteworthy that the yield of BTX is enhanced at the expense of that of n-paraffin, therefore, it can be assumed that BTX are formed in the micropore from the n-paraffin produced in the cracking process. At high ratio region, the n-paraffin does not linearly decrease despite linear increase of BTX formation. 4. CONCLUSION 1. The catalysts mixed with silica-alumina have larger meso- and macro-pore size and smaller surface area than those mixed with alumina. 2. Zeolite-Y in the substrate enhances the yields of both 350~ and naphtha. 3. The silica-alumina-based catalysts show lower coke formation than the alumina-based ones. 4. The high fraction of micropore surface area in the catalyst enhances the BTX formation whether the zeolite exists or not. The BTX seems to be formed from n-paraffin, produced by VGO cracking, in the micropore of the substrate.
Acknowledgement: The financial support for this work by the Research Center for Catalytic Technology (Pohang University of Science and Technology) is gratefully acknowledged. 5. REFERENCES 1 M.A. Altajam and M. Ternan, Fuel, 68 (1989) 955. 2 M.G.Yang, I, Nakamura, and K. Fujimoto, Catalysis Today, 43 (1998) 273. 3 C. Song, T. Nihonmatsu and M. Nomura, Ind. Eng. Chem. Res., 30 (1991) 1726. 4 Y. Arahata, Y. Tsukada, O. Imamura, H. Takagi and M. Fujita, Sekiyu Gakkaishi, 32 (1989) 306. 5 T. Itoh and Y. Tsuchida, Applied Catalysis, 51 (1989) 213. 6 Y. Mild, S. Yamadaya and M. Oba and Y. Sugimoto, J. of Catal., 83 (1983) 371. 7 S.C. Thompson and J.F. Mathews, Applied Catalysis, 47 (1989) 45. 8 K.H. Lee and B.H. Ha, Microporous and Mesoporous Materials, 23 (1998) 211.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
251
T e s t i n g a n d c h a r a c t e r i s a t i o n o f P t / A S A a n d P t P d / A S A for d e e p H D S r e a c t i o n s . H.R. Reinhoudt a, R. Troost a, A.D. van Langeveld a, J.A.R. van Veen b, S.T. Sie a and J.A. Moulijn a
a Delft University of Technology, Julianalaan 136, 2628 BL Delft, The Netherlands b Shell Research and Technology Centre Amsterdam, Badhuisweg 3, 1031 CM Amsterdam, The Netherlands
Abstract
The development of a dedicated, second stage deep HDS reactor to meet the low sulfur levels in diesel fuel is a challenging option, since it allows to process the feed under low HzS partial pressures. Because ASA supported noble metal catalysts have a high activity in the HDS of 4,6 di-alkylated DBT's in real gas oil, they constitute a new option. The nature of the active sites on ASA supported noble metal catalysts has been studied by electron microscopy and infrared spectrometry on adsorbed CO (FTIR(CO)). It is concluded that the support acidity plays a role in the genesis of active sites for HDS. Alloying Pt and Pd on ASA results in very active catalysts for conversion of 4-E,6-MDBT and DBT. Pt/ASA has a high selectivity for hydrogenolysis whereas PtPd/ASA shows a high selectivity for hydrogenation reactions. FTIR(CO) and the catalytic activity in DBT HDS indicate the presence of two different sites on PtPd/ASA: HDS sites, similar to those present in Pt/ASA and hydrogenation sites. The former consist of electron deficient Pt which facilitates the formation of sulfur vacancies, whereas the latter are are sulfur free, metal-like sites. The strong competitive adsorption between H2S, 4-E,6-MDBT and DBT suggests that the sulfur containing compounds are all adsorbed through the sulfur atom.
1. I N T R O D U C T I O N The application of a dedicated, second stage deep HDS reactor to meet the low sulfur levels in diesel fuel is a challenging option, since it allows to process the feed under low HzS partial pressures and, hence the application of other catalysts than those conventionally applied in hydrodesulfurization reactions. Since ASA supported noble metal catalysts have a high activity in the HDS of 4,6 di-alkyalated DBT's as well as in the deep HDS of a prehydrotreated gas oil, they constitute a new option [ 1]. It is suggested in the literature (for example [2-5]) that the support acidity enhances the activity of noble metal catalysts in HDS reactions by decreasing the electron density on highly dispersed noble metal clusters. A series of Pt-based catalysts on respectively 7-A1203, amorphous silica-alumina (ASA) and a stabilised Y-zeolite (XVUSY) has been prepared. Since Pd stabilises the performance of Pt/~/-AI203 [6] in the presence of HzS, the effect of alloying Pt/ASA with Pd will be evaluated also. The catalytic activity of these catalysts for the conversion of 4-ethyl, 6-methyl dibenzothiophene (4-E,6-MDBT) and DBT was determined next to the effect of the total sulfur concentration on the activity for 4-E,6-MDBT. The nature of the active sites on ASA supported noble metal catalysts has been evaluated from electron microscopy and FTIR(CO).
252 2. E X P E R I M E N T A L
2.1. Applied catalysts The metal loading of the applied catalysts is shown in Table 1. The applied T-AI203 and ASA are standard industrial supports and the XVUSY is a standard stabilised Y-zeolite. The acidity of the supports increases in the order 7-A1203, ASA and XVUSY. Table 1" Metal loadings of the applied catalysts Catalyst
PI/AI203
Pt/ASA I
Pt loading (w.%)
1.0
0.8
Pd loading (w.%)
-
Pt/ASA II
3.0
Pt/XVUSY
Pd/ASA
PtPd/ASA 1
PtPd/ASA lI
1.0
-
1.0
0.5
1.0
1.0
1.0
2.2. Model feed experiments The catalyst pellets were crushed and sieved and 0.2 g of the particles between 125 and 250 lam was used in the reaction. Prereduction was done in-situ in pure H2 at 0.2 MPa, at 573 K for 1 h. The activity tests were carried out in a stirred batch autoclave reactor, described elsewhere [7]. The model feed consisted of 0.5 w.% DBT (Fluka >98%) or 0.15 w.% 4-E,6-MDBT in n-hexadecane (Aldrich, 99%+). If added, H2S was introduced by thermal in-situ decomposition of dimethyldisulfide (DMDS). The test with DBT was carried out at 613 K and that with 4-E,6-MDBT at 633 K, both at a total pressure of 6.0 MPa. The first order reaction rate constant for the conversion of the reactant was normalised per gram catalyst. 2.3. H R T E M characterisation High Resolution Transmission Electron Microscopy (HRTEM) was done in a Philips CM 30 ST electron microscope with a field emission gun operated at 300 kV. The catalysts were powdered and applied on holy carbon in a few droplets of ethanol, followed by drying at 300 K. Fresh and sulfided samples were transported under ambient conditions and studied directly. 2.4. Infrared measurements on adsorbed CO FTIR measurements on adsorbed CO were performed in an in-situ transmission cell [8] using a Nicolet 550 spectrometer with a resolution of 4 cm -~. The fresh sample was pressed into self supporting wafer, introduced into the cell and evacuated for 0.5 h at 423 K. Reduction and sulfidation was done in respectively pure H2 or in 10 vol.% HzS in balance H2,, at a heating rate of 600 K h -l. The maximum temperature was maintained for 1 h. Subsequently, the sample was cooled in a flow of the reaction mixture and evacuated at 300 K. The transmission cell was cooled to 130 K and CO (99.999 %, Messer Griesheim) was introduced at 50 mbar.
3. RES ULTS 3.1. Activity for 4-E,6-MDBT In Figure 1, the activity for the conversion of 4-E,6-MDBT is shown for Pt/y-Al203, Pt/ASA I and Pt/XVUSY. Clearly, the conversion of 4-E,6-MDBT increases with increasing acidity of the support. Pt/XVUSY has the highest initial activity, but, at the time scale of an experiment (8 h), it strongly deactivates. The initial reaction rate constant is shown in Figure 1. Pt/ASA I and Pt/7-A1203 have a stable activity in these experiments. The activity was also tested in the presence of 7.5.10 -3 tool HzS per mol H2, leading to a decrease in the activity of all three
253
catalysts. Pt/XVUSY still has the highest activity but again strong deactivation was observed. The sulfur tolerance, 0~, defined by us as the ratio of the rate constant in the presence of 7.5.10 -3 mol H2S per mol H2 and that in the absence of added H2S, is presented in Table 2. For Pt/XVUSY, o~ was based on the initial rate constants. For all catalysts the sulfur tolerance appears to be about 0.4.
Figure l(left): Effect of the support on the first order reaction rate constant for conversion of 4-E,6-MDBT over Pt. Figure 2 (right): The first order reaction rate constant for conversion of 4-E,6-MDBT over different ASA supported Pt(Pd) catalysts. As can be seen from Figure 2, the support itself has no activity for the conversion of 4-E,6-MDBT. Increasing the Pt loading of Pt/ASA I from 0.8 w.% to 3.0 w.% (Pt/ASA II) only doubles the reaction rate constant. Combining Pt and Pd on ASA results in a very active catalyst, whereas Pd/ASA has a much lower activity than Pt/ASA. Also the Pt/Pd atomic ratio is important since PtPd/ASA II is much more active than PtPd/ASA I, notwithstanding its lower metal loading. The sulfur tolerance a of Pt/ASA and the PtPd/ASA catalysts is 0.5 (Table 2). Table 2: The sulfur tolerance o~ Catalyst
Pt/AI203
Pt/ASA I
o~
0.4
0.5
Pt/ASA II
0.5
Pt/XVUSY
Pd/ASA
PtPd/ASA I
PtPd/ASA II
0.4
0.4
0.5
0.5
3.2. Activity and selectivity for DBT HDS Figure 3 shows the reaction rate constant and selectivity for the conversion of DBT over Pt/ASA I, PtPd/ASA I and PtPd/ASA II. The ranking of the catalysts for DBT HDS is similar to that for 4-E,6-MDBT. Significant differences exist between the selectivity over different ASA supported catalysts, since Pt/ASA I has a high hydrogenolysis selectivity as can be inferred from the yield of biphenyl (BiPh). With the PtPd/ASA catalysts, the large amounts of cyclohexylbenzene (CHB) and bicyclohexyl (BCH) formed indicate a high hydrogenation selectivity. Note that the Pt/Pd ratio does not affect the selectivity. For all experiments, the mass balance was checked. After 5 h reaction, typically less than 5 % of the converted DBT could not be covered by the three main products.
254
Figure 3" The first order reaction rate constant and the selectivity of different ASA supported noble metal catalysts in DBT HDS. 3.3. The influence of the total sulfur concentration The nature of the active phase under reaction conditions may be affected by either the amount of HzS present, or the reactant concentration. From Figure 2, it is obvious that addition of 7.5.10 -3 mol HzS per mol H2 causes a decrease in the rate constant for the HDS of 4-E,6-MDBT. To assess the nature of the effect, the HDS of 4-E,6-MDBT in the presence of H2S was followed for 3.5 h. Then, the H2S was removed by flushing the reactor with hydrogen and the reaction was continued. After flushing, a high reaction rate was recovered indicating that effect of H2S poisoning is reversible. So competitive adsorption of H2S and 4-E,6-MDBT occurs. 2.28
[ ] 4-E, 6-M DBT
[ ] DMDS
[]
DBT
1.20 11 911
(~ S:
(1911
[
1 22
44
8.9
88
88
24
Total moles of sulfur. 104 [tool]
Figure 4: The effect of the total sulfur concentration on the first order reaction rate constant for conversion of 4-E,6-MDBT over Pt/ASA I. The source of sulfur and its relative amount are indicated. The rate constant for 4-E,6-MDBT strongly increases at lower 4-E,6-MDBT concentrations as is shown in Figure 4. This means that strong self-inhibition of the reacting sulfur compounds occurs. When 50 % of the 4-E,6-MDBT is replaced by the same number of moles sulfur in the form of DMDS, the rate constant for 4-E,6-MDBT HDS is found to be quite similar. An analogous result is found when DBT is added in stead of DMDS. Clearly, the type of the sulfur compound is not relevant for its competitive effect on 4-E,6-MDBT HDS. 3.4. H R E M Characterisation On calcined Pt/ASA-I, particles are present with a size of 2-5 nm. EDX analysis shows that these particles contain Pt. In sulfided Pt/ASA I, the majority of particles has a diameter of
255
2-4 nm, while also small clusters of ca 1.2 nm could be seen. The spacing of the diffraction lines of the 2-4 nm particles indicates that they contain metallic Pt and must consist of at least 4 atomic layers. In a spent Pt/ASA I catalyst, Pt particles with similar size and nature were observed as on the freshly sulfided catalyst, even though the sample has a completely different history. On oxidic Pt/~t-Al203 catalyst no Pt particles were seen, in accord with its H2 dispersion close to 1. In contrast, on sulfided Pt/7-A1203, uniform Pt particles with a diameter of ca 2 nm and virtually no smaller particles were seen. In the fresh PtPd/ASA II catalyst mostly particles of 2-3 nm were observed. With EDX, no separate Pt or Pd particles could be observed, suggesting that a Pt-Pd alloy is present. An important difference between the sulfided and the fresh PtPd/ASA II catalyst is the number of particles, which is clearly larger on the sulfided catalyst. This indicates that also PtPd bi-metallic clusters are subject to sintering under H2S. Especially on the sulfided catalyst particles were visible in the order of 1.5 nm and smaller. 3.5. C h a r a c t e r i s a t i o n by F T I R on a d s o r b e d C O The FTIR(CO) spectrum (50 mbar) on reduced Pt/ASA I (Figure 5A) shows prominent bands at 2163 cm -1, 2086 cm -1 and 1992 cm -1 and a weak band around 2120 cm -1. Deconvolution of the spectrum revealed the presence of a weak contribution at 2147 cm -1. The FTIR(CO) spectrum (50 mbar) of Pt/ASA I after sulfiding and reduction is shown in Figure 5B. Most important difference with the spectrum of the reduced sample is the absence of the 2086 cm -1 band. The intensity of the bands at 2147 and 2120 cm -1 is increased as compared to the 2163 cm -1 band. The band at 1992 cm -~ remains present in the spectrum.
2163 ',
2163
A
2147
~ 2147 i/j
i
2107
/i/,~ 2J20
/IN.
| |
!
B
2200
2150
21 O0
2050
2000
Wave number [cm -~]
1950
2200
2150
21 O0
2050
2000
1950
Wave number [cm-l]
Figure 5 (left)" A, The FTIR(CO) (50 mbar) spectrum of reduced Pt/ASA I. B, The FTIR(CO) (50 mbar) spectrum of subsequently sulfided and reduced Pt/ASA I. Figure 6 (right): A, The FTIR(CO) (50 mbar) spectrum of reduced PtPd/ASA II. B, The FTIR(CO) (50 mbar) spectrum of subsequently sulfided and reduced PtPd/ASA II. In the deconvoluted FTIR(CO) spectrum (50 mbar) of the reduced PtPd/ASA II catalyst (Figure 6A), a strong band is observed at 2107 cm -1 with a shoulder at 2086 cm -1. Also, a band at about 1990 cm -~ is present. As in the spectrum of Pt/ASA, a band at 2163 cm -1 is present. After
256 sulfiding and subsequent reduction (Figure 6B) a band is present at 2163 cm -~. Furthermore, three small bands at 2142, 2120 and 2107 cm -~ are visible in the spectrum whereas the band at 2085 cm -l has disappeared.
4. DISCUSSION 4.1. Activity Based on the results shown in Figure 1, the acidity of the support strongly enhances the activity of Pt based catalysts in deep HDS. However there seems to be an optimum, since a strong deactivation is observed for the XVUSY based catalyst. This deactivation may be due to coke deposition, induced by the strong acidic sites. The acidity of ASA apparently combines a high activity for HDS reactions and a stable activity. Addition of DMDS (7.5.10 -3 mol HzS per mol H2) causes a significant decrease in the activity for all tested platinum catalysts. Since the sulfur tolerance is found to be similar (ca 0.4) for the three supports, it is proposed that the enhanced HDS activity of Pt on acidic supports can be attributed to a larger number of active sites, in accord with Echevskii and Ione [9]. Remarkably, similar sites seem to be also present in Pt/y-AI203, although the numbers are much lower. The influence of the support on the active phase underlines that the interaction of the active phase with the support is essential for the nature of the active sites. Hence it is logical to expect that the active sites are associated with small metal clusters. Not only the support is important, also the noble metal, its loading and the addition of a second noble metal. For Pt on ASA, an increase of a factor 3.5 in the metal loading only leads to an increase in the activity of amply a factor 2. This suggests that an important fraction of the metal is not incorporated in active sites. Although Pd/ASA itself has a lower activity than Pt/ASA, combined Pt and Pd results in very active catalysts. Also the atomic ratio of Pt and Pd does strongly affect the catalytic performance; PtPd/ASA II, which has a lower metal loading than PtPd/ASA I still shows a significantly higher activity. The high hydrogenation selectivity in DBT HDS suggests that in addition to HDS sites also a second type of active site is present on PtPd/ASA. 4.2. Nature of the active sites on Pt/ASA The role of acidic supports on the activity of noble metals in the presence of sulfur is often referred to as 'sulfur tolerance', which was by explained Sachtler and Stakheev as an electron transfer from the metal cluster to the support [10]. Since the bond between a metal and sulfur induces an electron transfer from the metal to the sulfur atom, a small positive charge on the metal lowers the stability of the metal-sulfur bond [2]. According to Anderson [ 11], a strongly induced cluster can be not larger than 1 nm. The 1.2 - 1.5 nm particles, observed on sulfided Pt/ASA I which were not present on the calcined catalyst, indicate that agglomeration of small Pt clusters has occurred during sulfiding. The beneficial role of the acidic support can possibly be attributed to the stabilisation of small metal clusters in the presence of HzS. As discussed above, no Pt particles were observed in fresh Pt/T-AI203, whereas sulfided Pt/y-AI203 contains uniformly distributed particles of ca 2 nm, which clearly points to agglomeration of Pt. Assuming that the high activity of Pt on acidic supports is caused by small Pt clusters, the next question is if the active sites on ASA supported catalysts are related to the presence of sulfur free sites. FTIR (CO) on H2 reduced Pt/ASA I shows five bands. The most prominent band at 2163 cm -l can be attributed to CO adsorption on acidic OH groups of the support. The nature of the shoulder at 2147 cm -~ is not fully clear. It can be related to a Pt a+ species [12], or condensed CO, which would give a contribution at 2143 cm -l. Lee et al. [13] identified the band at 2121 cm -l as
257
an electron deficient Pt a+ species. The band at 2086 cm -~ can be attributed to linearly adsorbed CO on Pt ~ [13]. A sulfided and subsequently reduced Pt/ASA I catalyst does not show the band at 2086 cm -l, characteristic for CO on Pt ~ Clearly treatment in pure H2 at 573 K is insufficient to restore a significant amount of sulfur free Pt ~ surface, which is in accord with H2 chemisorption measurements. Based on these findings it seems unlikely that Pt ~ sites are exposed to the gas phase under reaction conditions. The 2121 cm -~ band is still present indicating that electron deficient Pt a+ sites may play a role in the HDS reaction, in contrast to metallic sites. In summary, it is concluded that the active sites consist of sulfur vacancies on electron deficient Pt sites, which is consistent with the low hydrogenation selectivity in DBT HDS.
4.3. Active sites on PtPd/ASA HRTEM on oxidic PtPd/ASA II shows that most of the metal particles are smaller than 3 nm. In contrast, on sulfided PtPd/ASA II also smaller metal particles of about 1.5 nm were seen and in adddition the number of particles was much larger than in the oxidic samples. Based on the higher activity it is concluded that the number of active sites increases as compared to Pt/ASA. Also the nature of the active sites changes, as can be inferred form the change in selectivity. In the FTIR(CO) spectrum of the reduced sample, a band at 2107 cm -~, which is attributed to metal-like sites, is the dominant one. A shoulder at 2086 cm -l, which was also observed in the reduced Pt/ASA I, can be ascribed to linearly adsorbed CO on Pt ~ Hence, reduction must be almost complete. Analogous to Pt/ASA I, in the subsequently sulfided and reduced PtPd/ASA II, nearly no Pt ~ could be observed, although a small band remains visible at 2107 cm -~, attributed to Pd~ sites [12]. So, in contrast to sulfided Pt/ASA I, metal-like sites seem to be present on subsequently sulfided and reduced PtPd/ASA II. The observed high hydrogenation activity of PtPd/ASA is in line with this. The FTIR(CO) on sulfided PtPd/ASA II shows a more intense 2121 cm -~ band, which was assigned to Pt ~+ or Pd a+ sites [14]. As it seems, the presence of Pt a+, is important in the creation of sulfur vacancies as was also indicated for Pt/ASA I. 4.4. HDS kinetics for ASA supported noble metal catalysts Addition of H2S strongly decreases the rate constant for the conversion of 4-E,6-MDBT. Based on the reasoning above, we do not expect a progressive coverage of the active sites by sulfur to be the cause for the lower activity. The results demonstrate that competitive adsorption of HzS and 4-E,6-MDBT is more likely the cause for the lower activity since the rate constant directly increases upon removal of H2S. Hence, the adsorption constant of H2S on the active sites of Pt/ASA I must be high. A first order behaviour for the conversion of 4-E,6-MDBT is observed up to a high conversion. Since HzS is produced during the reaction, one would expect that the rate constant decreases with increasing conversion. Also, a high concentration of 4-E,6-MDBT results in a lower rate constant. Combining these observations, leads to the conclusion that the adsorption constants of 4-E,6-MDBT and H2S must be of the same order. This is in line with the above mentioned experiment in which 50% of the 4-E,6-MDBT was replaced by the same number of moles H2S or DBT. The analogy between the inhibition of 4-E,6-MDBT, DBT and HzS suggests that the mode of adsorption must be the same for the three molecules, i.e. it must occur through the sulfur atom. Obviously, this imposes restrictions on the size of the active sites. The proposition that small clusters of Pt atoms are related to the active sites on Pt/ASA I fits within these constraints.
258
5. C O N C L U S I O N S The acidity of the support plays an important role in the activity of noble metal catalysts in HDS and hydrogenation reactions. Pt/ASA and Pt/XVUSY are active catalysts in the conversion of 4-E,6-MDBT although Pt/XVUSY showed strong deactivation . ASA however, combines a high activity and a stable catalytic performance indicating that the tuning of the support acidity is important. The support acidity plays a role in the stabilisation of a larger number of active sites for HDS rather than in changing the nature of these sites. Alloying Pt and Pd on ASA results in very active catalysts for conversion of 4-E,6-MDBT and DBT. In DBT HDS, Pt/ASA has a high selectivity for hydrogenolysis whereas PtPd/ASA shows a high hydrogenation selectivity. Based on HRTEM and FTIR(CO) results it was deduced that small (< 1.5 nm), electron deficient Pt sites in small clusters most likely play an important role in the HDS activity of Pt/ASA I. The low selectivity for hydrogenated products in DBT HDS is consistent with the fact that sulfur vacancies are observed only for electron deficient Pt . FTIR(CO) and the catalytic activity in DBT HDS indicate the presence of two different active sites on PtPd/ASA: HDS sites, similar those present on Pt/ASA and hydrogenation, sulfur free metal-like sites. The strong competitive adsorption between HzS, 4-E,6-MDBT and DBT, suggests that the sulfur containing compounds are all adsorbed through the sulfur atom, which also points to the presence of small clusters as catalytically active sites.
6. A C K N O W L E D G E M E N T Dr. P. Kooyman, National Centre for HREM is gratefully acknowledged for the HREM analysis. The research has been performed under auspices of NIOK, the Netherlands Institute for Catalysis Research, Lab Report TUD 99-4-992. 7. REFERENCES
I. H.R. Reinhoudt, R. Troost, S. Van Schalkwijk, A. D. Van Langeveld, S. T. Sie, H. Schulz, D. Chadwick, J. F. Cambra, V. H. J. De Beer, J. A. R. Van Veen, J. L. G. Fierro, and J. A. Moulijn, Stud.Surf.Sci.Catal., 106, 1997,237. 2. M. Guenin, M. Breysse, R. Frety, L. Tifouti, P. Marecot, and J. Barbier, J.Catal., 105, 1987, 144. 3. G.D. Chukin, B. V. Smirnov, V. I. Malevich, M. V. Landau, V. Ya. Kruglikov, and N. V. Goncharova, Kinet.Katal., 17, 1976, 1097. 4. M.V. Landau, V. Ya. Kruglikov, N. V. Goncharova, O. D. Konoval'chikov, G. D. Chukin, B. V. Smirnov, and V. I. Malevich, Kinet.Katal., 17, 1976, 1104. 5. J. Barbier, P. Marecot, L. Tifouti, M. Guenin, and R. Frety, Appl.Catal., 19, 1985, 375. 6. T.B. Lin, C. A. Jan, and J. R. Chang, Ind.Eng.Chem.Res., 34, 1995, 4285. 7. J.-P. Janssens, PhD. Thesis, 1997, T.U. Delft, The Netherlands. 8. R. Mariscal, H. R. Reinhoudt, A. D. Van Langeveld, and J. A. Moulijn, Vibr. Spec., 16, 1998, 119. 9. G.V. Echevskii, and K. G. Ione, In: Catalysis by Zeolites., Elsevier, Amsterdam ,1980, 273. 10. W. M. H. Sachtler and A. Y. Stakheev, Catal.Today, 12, 1992, 283. 11. J. R. Anderson, In: Structure and properties of small particles, Ac. Press, London, 1975. 12. V. N. Romannikov, K. G. Ione, and L. A. Pedersen, J.Catal., 66, 1980, 121. 13. J. K. Lee and H.-K. Rhee, J. Catal., 177, 1998, 208. 14. M.M. Otten, M.J. Clayton and H.H. Lamb, J. Catal., 149, 1994, 211.
CHARA CTERlZA TION OF CATALYSTS
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
261
Probing the Electronic State of Nickel-Molybdenum Sulphide Catalysts using ortho-Xylene Hydrogenation L. Fischer, V. Harl6, S. Kasztelan IFP, Kinetics and Catalysis Division, I e t 4 Avenue de Bois Pr6au, 92852 Rueil-Malmaison Cedex, France
Ortho-xylene was hydrogenated at 300~ under 60 bar total pressure over sulphided Ni, Mo and NiMo alumina catalysts. Among the 1,2-dimethylcyclohexanes formed, the ratio of the stereoisomers cis/trans was found to be sensitive to the electronic state of the catalysts' active sites. Applied to a series of NiMo catalysts containing increasing amounts of nickel, the reaction suggests a successive decoration of MoS2 slabs with the promoter. Hydrogenation of orthoxylene is a powerful new method for probing electronic effects specifically on active sites of sulphide catalysts under typical hydrotreating conditions.
1. INTRODUCTION Catalytic activity of sulphide based hydrotreating catalysts is often related to electronic effects. The electronic structure of binary sulfides is supposed to be at the origin of differences in their activity [1]. In NiMo sulphide active phase, an electron transfer from the promoter nickel to molybdenum might be the reason for the promoted activity, as shown by X-ray photoelectron spectroscopy [2]. Electronic effects are also often proposed for the promoting effect of additives such as phosphorus or fluorine [3, 4]. However, the experimental determination of electronic effects in sulphide based catalysts remains a challenge. The classical methods employed for identification of electronic effects such as infrared spectroscopy (IR) of adsorbed CO or NO and X-ray photoelectron spectroscopy (XPS) have inherent difficulties. In particular, these methods investigate the catalyst's surface under conditions which are very far from typical hydrotreating conditions : atmospheric or vacuum treatment or pre-treatment is applied, often at high temperature in the case of IR and in presence of an electron beam in the case of XPS. These conditions may change the surface state by partial elimination of SH groups [5] and other weekly bonded surface species or even may reduce surface sites. In addition, XPS sensitivity for electronic effects is rather poor, and it is often difficult to
262
distinguish, for the same type of elements, ions of active catalytic sites from others. In general, clear evidences for electronic effects are difficult to obtain in the case of sulphide catalysts and very few have been published to our knowledge [ 1 - 3 , 6]. More evidences have been reported for metallic catalysts. For example, ortho-xylene hydrogenation has been found to provide information on electronic effects on alumina supported platinum catalysts under mild conditions (100~ lbar) [7]. The predominant formation of the stereo isomer cis - 1,2 - dimethyl cyclohexane is attributed to an active phase having a high electron density, whilst formation of a more important fraction of the trans stereoisomer is attributed to active phases with lower electron density, hence a stronger electron acceptor character. Some other authors linked the formation of either trans- or cis-l,2-dimetylcyclohexane to the strength of interaction between ortho-xylene and the catalyst surface sites [8, 9]. We will show in this preliminary report t h a t ortho-xylene hydrogenation can also be used to probe electronic effects on sulphide catalysts under typical hydrotreating conditions.
2. E X P E R I M E N T A L 2.1 P r e p a r a t i o n Ni, Mo and NiMo catalysts were prepared by pore-volume impregnation of a 7-A1203 support (240 m2/g, 0.64 cm3/g, extrudates of 1.2 mm diameter). Catalysts were dried 24 hours at 393 K and calcined for two hours in dry air at 773 K. Ni catalysts contain 3 or 8 weight percent (wt %) Ni, Mo catalysts contain 9 wt % Mo. NiMo catalysts have a constant molybdenum loading of 9 wt % and different nickel loadings of 1 to 4 wt % Ni corresponding to Ni/Mo atomic ratios between 0.15 and 0.72. Precursors were (NH4)6MoT024.4H20 (Merck) and Ni(NO3)2.6H20 (Prolabo). Impregnated metals were uniformly dispersed in the grain as found by electron microprobe analysis. Surface areas, determined by BET method, were not affected in a significant way by the impregnation with metals.
2.2 Catalytic tests Catalytic tests were performed in a fixed-bed reactor working in dynamic regime using 40 ml of catalyst (80 ml for pure Ni catalysts). Prior to test, the catalysts were sulfided for 2 hours at 623 K and 6 MPa total pressure, with a ratio HJhydrocarbons = 450 Nlfl, using a feed of 6 wt % dimethyl disulfide, 20 wt% ortho-xylene and 74 wt % cyclohexane. Catalytic activities were m e a s u r e d in conditions to favour hydrogenation and limit isomerization reactions : 573 K and HJhydrocarbons = 1000 N1/1, i.e. 4.8 MPa H2, 0.8 MPa cyclohexane, 0.2 MPa oxylene, 0.1 MPa H2S and 0.1 MPa CH4. The conversion of the ortho-xylene hydrogenation reaction (HYD) is based on the amount of all dimethyl cyclohexanes and trimethyl cyclopentanes produced. Catalytic activities are expressed in the following as pseudo first order rate constants in mol/g/h.
263
3. R E S U L T S
AND DISCUSSION
In catalysis by metals under mild conditions [7], the only products of orthoxylene hydrogenation are cis- and trans- 1,2 - dimethyl cyclohexanes. On sulphide catalysts under hydrocracking conditions, several other reactions can occur as illustrated in figure 1 [10].
I
Figure 1"
Scheme of ortho-xylene hydrogenation conditions, according to [10].
under
hydrocracking
Under the reaction conditions used in this work, the main reaction products for all catalysts are dimethyl cyclohexanes and trimethyl cyclopentanes. The concentration of toluene, C3-benzenes and their hydrogenation products are between 10 and 60 times inferior to the concentration of all other hydrogenation products of ortho-xylene and at least 1000 times inferior to the ortho-xylene concentration, so t h a t we can neglect reaction path E. According to [10], the hydrogenation velocities of all three xylenes (paths A and G) are similar. There are 20 wt % ortho-xylene and 0.12 % meta- and paraxylene as impurities in the feed. The same amount (0.10% to 0.13 %) of metaand para-xylene is found in the reaction effluents at LHSV = 2 h-1. As the concentration of ortho-xylene in the effluents does not fall below 15 % at LHSV = 2 h -1, the concentration of o-xylene is always much higher t h a n the concentration of m- and p-xylene. Consequently, we also can neglect formation of dimethyl cyclohexanes by paths F and G. Hence, it can be considered t h a t all hydrogenation products of ortho-xylene essentially come from direct hydrogenation of ortho-xylene via path A. It is useful to compare the product distribution to thermodynamic equilibrium. As thermodynamic data are not available for all trimethyl cyclopentanes and as the peaks of trans-l,3- and cis-l,4-dimethyl cyclohexanes are not separated in our chromatographic conditions, the calculation has been limited to the dimethyl cyclohexanes ~, cis- and trans-l,2, 1,1, cis-l,3 and trans1,4 ~. Results are illustrated in figure 2.
264
Figure 2 :
Distribution among dimethyl cyclohexanes in effluents after catalytic tests, compared to thermodynamic equilibrium [11]. The experimental error is estimated between _+0.1% and _+1 % absolute, the m a x i m u m error being observed for Ni catalysts because of their poor catalytic activities.
As it can be seen in figure 2, the product distribution obtained for all catalysts does not correspond to the thermodynamic equilibrium. Moreover, all the ratios ,,cis/(cis+trans)~ of 1,2-dimethyl-cyclohexanes measured in this work are in a range from 43% cis to 53% cis, far from the thermodynamic equilibrium at 23 % cis in those conditions [11, 12]. We verified whether the isomerization of desorbed products is negligible in comparison to the hydrogenation reaction. The first order isomerization activities of cis-l,2-dimethyl-cyclohexane or trans-l,2-dimethyl-cyclohexane have been measured in the same conditions as the hydrogenation of ortho-xylene. It has been found t h a t the isomerization activities are always between 10 and 100 times inferior to the hydrogenation activities. The secondary isomerization of the 1,2dimethyl-cyclohexanes can therefore be neglected. In the hydrogenation of ortho-xylene, a change in contact time has practically no effect on the selectivity for the different isomers of the hydrogenation products. This is illustrated in figure 3, by subdividing the hydrogenation products into one group containing the 1,2-dimethyl-cyclohexanes and another group containing all other hydrogenation products. Both product groups appear to be primary product groups, in agreement with results obtained under hydrocracking conditions by other authors [10].
265
40-or..r
30- -
o
~
~
1,2-dimethylcyclohexanes other hydrogenation products
lO-
o
o
I
0
10 20 30 40 overall HYD conversion (%)
Figure 3 9 Hydrogenation conversion of two product groups ,, I ~ with respect to the overall HYD conversion. NiMo catalyst (Ni/Mo = 0.5). Distinguishing the products more precisely, the ratios ,, cis/(cis+trans)~ of the 1,2 - dimethyl cyclohexanes are i n d e p e n d e n t from catalytic conversion (figure 4). The ratio r e m a i n s constant for one catalyst over a large range of conversion. More important, each catalyst appears to be characterized by a different value of the , , cis/(cis+trans) ~ molar ratio. 54'--
5 2 - -3
9 "
2 1.5 1
Mo
(9 wt%)
50-
= ,1,8+ 4:6- _0.5 o~,,i
44-
-'- Ni (3 and 8 wt%) 0.5 -~
~, NiMo (Ni/Mo=0.27) ,L
C9
42-
,, 4
~3
~2
NiMo ~" (Ni/Mo=0.5)
40 1
10 o-xylene hydrogenation [%]
100
Figure 4 9 Molar ratios , , cis/(cis+trans) ~ of the reaction products 1,2-dimethylcyclohexanes, as function of the ortho-xylene hydrogenation conversion. LHSV (h-D are m a r k e d u n d e r each point. By analogy to catalysis by metals [7], one m i g h t suppose t h a t a high electronic density of the catalyst's active site should favour the formation of cis 1,2 - dimethyl cyclohexane, whilst a low electronic density should favour the formation of the stereoisomer trans. In order to verify, if an increase of the electronic density really increases the ratio ~, cis/(cis+trans)~, a nitrogen base (1,2,3,4 - t e t r a h y d r o isoquinoleine, THIQ) was injected during a catalytic test.
266
This compound is completely denitrogenated u n d e r our reaction conditions [13]. For this reason, we can consider a m m o n i a as the donor of electrons. 1 and 3 wt % THIQ in the feed diminish strongly the ortho-xylene hydrogenation over the NiMo/A1203 catalyst at LHSV = 1 h -1 (figure 5). This result is expected, as nitrogen compounds act as strong inhibitors on hydrotreating catalysts. At the same time, the ratio ,, cis/(cis+trans),, increases. Then, the catalyst is , , washed ,, at LHSV = 3 h-1 with a feed containing no THIQ. Three hours ,, w a s h i n g ~ are not long enough to recover the original activity, but it seems t h a t activity and the cis/(cis+trans) ratio tend to the original values. Hence, the effect is reversible. Finally, THIQ is one more time injected at LHSV = 3 h -~, and the effect is the same as previously described. Indeed, it can be seen in figure 5 t h a t an electron donor favours the formation of the isomer cis, as expected. 12-
~
,-~
~~
LHSV = 1
10
8-
t
- 54
3
I
I
I
,
,,
I I
I
4--
0
=
_
-~..-
2--Y
LHSV
I
S'"
I --
-
0% T H I Q I
250
- 50 -48
~
-46
+
I
,
.1,...I
13%
1% THIQ
450
I~
,,
-52
-44
- i -
I
_
I THIQ
,,
I
I
0% THIQ.
650 time [minutes]
Activity ~
-42
II 1~ THIQ I
t
850
1050
~
40
cis / (cis+trans) [%] I
Figure 5 9 Injection of 1,2,3,4-tetrahydro-isoquinoleine ( T H I Q ) i n t o the feed during ortho-xylene hydrogenation. NiMo catalyst (Ni/Mo = 0.5). In figure 4, the comparison of catalytic activities in ortho-xylene hydrogenation leads to the known activity variations of different sulphide phases. The catalytic activity increases in the order Ni<Mo
267
Therefore, our results show t h a t the hydrogenation of ortho xylene can be used as a method for determining electronic effects on sulphide catalysts' active sites. For a series of NiMo catalysts with a constant Mo content and different ratios NifMo, a linear relation is found between hydrogenation activity and the ratio ,~ cis/(cis+trans)~ (figure 6). An increasing content of nickel increases the activity and decreases the value for ,~ cis/(cis+trans) ~>. 12 10
Ni/Mo = 0.72 ~/Ni ~\Ni Ni/Mo
Z
= 0.5 ~
Mo N>
.
9
~
.
Mo
.
6
('M~
*~ 4
oW,,~
"~Mo
/
Ni/Mo = 0.15 9 " ~ ~ o 40
I
t
42
44
cisJ(cis+trans)
I
I
I
t
46 48 50 52 54 [%] = electron density .._
Figure 6 9 Activities of NiMo catalysts (Ni/Mo = 0 to 0.72) in ortho-xylene hydrogenation, as a function of the ratio , , cis/(cis+trans)~ of 1,2dimethyl-cyclohexanes. All activities m e a s u r e d at LHSV = 2 h -1. In terms of electronic density, 53 % cis/(cis+trans) should correspond to an electron-rich pure m o l y b d e n u m active site. On the other hand, the ratio of 43 % cis should correspond to an electron-poor nickel site in decoration of a MoSe slab. In the catalysts h a v i n g an intermediate NifMo ratio of 0.15 and 0.27, MoSe slabs should be decorated only partially with nickel atoms. If we assume by simplification that, w h a t e v e r the local environment, the active sites (Mo or ~, NiMo ~,) have always the same activity and electron density, then the resulting activities and cis/(cis+trans) ratios of the i n t e r m e d i a t e catalysts are in fact m a t h e m a t i c a l m e a n values of electron-rich Mo sites and electron-poor Ni sites. The observed linear relation suggests a progressive decoration of MoSe slabs with nickel atoms. O p t i m u m decoration of the MoS2 slabs with the promoter Ni appears to be achieved at Ni/Mo = 0.5, as a higher nickel loading (NifMo = 0.72) has no effect on the catalytic conversion. There is also no effect on the ratio cis/(cis+trans). This might suggest t h a t an excess of nickel above Ni/Mo = 0.5 leads to a formation of NiS particles, in analogy to [14, 17]. Although these particles are more selective in the stereoisomer cis, their activity is very poor in comparison to the NiMo phase. Consequently, their impact on the total cis/trans ratio is not significant. The results concerning the NiMo series can therefore be interpreted in a satisfactory m a n n e r with the model of a ,, mixed NiMoS sulphide phase ,~. On the
268
other hand, they are difficult to explain with solely two distinct phases of nickel sulphide and molybdenum sulphide, as proposed in the literature [18, 19]. 4. C O N C L U S I O N Ortho-xylene hydrogenation is a powerful new method for probing electronic effects on sulphide catalysts in typical hydrotreating conditions. The electronic state observed is specific for the active sites in interaction with the reacting molecule, under the specified conditions. The richer an active site is in electronic density, the higher is the ratio cis/trans of 1,2 dimethyl cyclohexanes. In a series of NiMo catalysts containing different amounts of nickel and a constant amount of molybdenum, a linear relation between catalytic activity and the ratio cis/trans has been found, suggesting a progressive decoration of MoS2 slabs with nickel atoms when the amount of Ni increases.
REFERENCES 1. R.R. Chianelli, M. Daage, M.J. Ledoux, Adv. Catal., 40 (1994) 177. 2. S. Houssenbay, S. Kasztelan, H. Toulhoat, J. P. Bonnelle and J. Grimblot, J. Phys. Chem., 93 (1989) 7176. 3. A.N. Startsev, O. V. Klimov, A. V. Kalinkin and V. M. Mastikhin, Kinetics and Catal., 35-4 (1994) 552. 4. G. Muralidhar, F. E. Massoth and J. Shabtai, J. Catal., 85 (1984) 44. 5. C. Petit, F. Maug~, J.-C. Lavalley, Stud. Surf. Sci. Catal.,106 (1997) 157. 6. R. Hublaut, O. Poulet, S. Kasztelan, E. Payen, J. Grimblot, Prepr.- Am. Chem. Soc., Div. Pet. Chem., 39(4) (1994) 548. 7. K. Chandes, B. Didillon, Communication Europacat 1, Montpellier (1993). 8. N. Martin, G. Cordoba, A. Lopez-Gaona, M. Viniegra, React. Kinet. Catal. Lett., 44-2 (1991) 381. 9. S. Smeds, D. Murzin, T. Salmi, Appl. Cata. A, 150 (1997) 115. 10. M. Guisnet, J. Thomazeau, J.L. Lemberton, S. Mignard, J. Catal., 151 (1995) 102, and references cited therein. 11. D.R. Stull, E.F. Westrum, G.C. Sinke, The Chemical Thermodynamics of Organic Compounds, Krieger, Malabas, Florida 1987. 12. R. Maurel, G. Leclerq, P. Hell, L. Leclerq, Bull. Soc. Chim. Fr., 6 (1971) 1967. 13. P. da Silva, PhD Thesis, Paris (1998). 14. H. Tops0e, B.S. Clausen, R. Candia, C. Wivel, S.J. M0rup, J. Catal., 68 (1981) 433. 15. M.J. Yacaman, R.R. Chianelli, J.L. Gland, Paper E6.4, MRS Fall Meeting, Boston, MA 1984. 16. S.P.A. Louwers, R. Prins, Prepr. Div. Pet. Am. Chem. Soc., 35 (1990) 211. 17. H. Tops0e, B.S. Clausen, in : H. Heinemann, G. Somorjai (Eds.), Catalysis and Surface Science, Dekker, New York, 1985, p. 95. 18. B. Delmon in: D.L. Trimm et al. (Eds), Catalysis in Petroleum Refining 1889, Elsevier Amsterdam 1990, p.1. 19. B. Delmon, Bull. Soc. Chim. Belg., 95 (1995) 173.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delrnon, G.F. Froment and P. Grange (Editors) 1999Elsevier Science B.V. All rights reserved. @
IR Study of Hydrotreating Catalysts in Working Conditions :Comparison of the Acidity Present on the Sulfided Phase and on the Alumina Support. Arnaud TRAVERT and Franqoise MAUGI? Laboratoire Catalyse et Spectrochirnie, UMR CNRS 6506 Markchal Juin - F- 14050 Caen - France.
-
ISMRa, 6, Boulevard du
Abstract The design of an apparatus suitable for monitoring in situ surface sites of sulfided catalysts simultaneously by transmission FTIR spectroscopy and by activity measurements is reported. It has been used to evaluate Brmted acidity of hydrotreating catalysts in "working" conditions, i.e. at high temperature in presence of H2S and hydrocarbons in the flow. Acidity of A1203 and CoMo/A&O3has been characterized by monitoring the adsorption of 2,6dimethylpyridine (DMP) at high temperature (100-250°C). DMP reveals a higher number of Brmted acid sites on CoMo/A1203than on Al2O3.H2S adsorption leads to the creation of Brsnsted acid sites on both catalysts but in a larger extent on CoMo/Al203. Evaluation of the Brsnsted acidity of catalysts was also performed using 3,3-dimethyl-l-butene (33DME31) isomerization as test reaction, carried out in situ either in the IR reactor-cell or in a conventional glass micro-reactor. Initial activities of the activated catalysts are in good agreement with the relative amounts of Brsnsted acid sites probed by DMP. Isomerization activity of both catalysts is also promoted in presence of H2S. Simultaneous measurements of activity and DMP adsorption followed by FTIR spectroscopy during the reaction at 250°C shows that Brsnsted acid sites probed by DMP are the active sites for 33DMB1 isomerization. Deactivation of the sulfide catalyst is attributed to the formation on the sulfided phase of a species which poisons Br~nstedacid sites. Coke formation was also evidenced 1. INTRODUCTION
Activity of hydrotreating catalysts is oftently related to coordinative unsaturated sites (CUS) located on the sulfide phase. H2S which is present during the reaction can adsorb on these vacancies and competes with the other components of the reaction feed. The resulting effect depends both on the nature of the reaction (HDS, HDN or hydrogenation) and on the reaction conditions (H2S pressure, nature of the feed ...). For a long time, H2S has been considered only as an inhibitor of hydrotreating reactions. Some years ago, however, it has been reported that H2S is able to promote HDN reactions [1,2] and recently it has been demonstrated that H2S could also participate to HDS [3,4] and hydrogenation reactions [5]. These results are usually attributed to the formation of 'new' active species by H2S adsorption. Among these, su&ydril groups (-SH) with a nucleophilic or an acidic character are usually invoked. In situ detection of these species, however, is not obvious. On a RuS2 bulk catalyst, Lacroix et al. evidenced heterolytic dissociation of H2S by 'H NMR [6] and inelastic neutron scattering studies [7]. Using FTIR spectroscopy, Topsse et al. indirectly evidenced the presence of such species on Mo-based sulfide catalysts [8] and further confirmed the presence of acidic -SH groups on Mo/A1203 and CoMoIAlzO3 using FTIR spectroscopy of pyridine adsorbed at high temperature [9]. In a previous work [lo], using dimethyl-2,6-pyridine adsorbed at room temperature, we were able to confirm an increase in To whom correspondence should be addressed. E-mail: [email protected]
270
Bronsted acidity related to the presence of the sulfided phase on similar catalytic systems. Moreover, we could evidence that H2S adsorption leads to an increase in Bronsted acidity on these catalysts and that the presence o f a sulfided phase enhances this phenomenon. In order to go closer to the reaction conditions, we set up a new equipment suitable for transmission FTIR in situ monitoring o f sulfide catalysts in a wide range of experimental conditions (25 - 4 0 0 ~ 0 - 3MPa). The present paper reports the first results obtained. Acidity of sulfided A1203 and CoMo/A1203 were characterized in three steps: (i) by probe molecule adsorption at high temperature under flow, (ii) by measuring the catalytic activities in 3,3-dimethyl-l-butene isomerization in a classical glass microreactor, i.e. a test reaction specific o f weak Bronsted acidity [11,12] and (iii) by conducting the reaction in the IR reactor-cell which allowed us to characterize simultaneously the activity o f the catalyst and the adsorbed species by FTIR spectroscopy. 2. E X P E R I M E N T A L
2.1. Apparatus The principle of the cell has already been published [13]. Some modifications, however, were brought about in order to achieve a complete tightness in the range 298-673K and 0-3MPa. The reactor-cell body is made in stainless steel tubing. CaF2 windows are inserted in stainless steel holders and sealing consists in a Kalrez O-ring inserted between two PTFE rings. Air cooling limits excessive thermal gradients. KBr rods inside the reactor
Figure 1 IR reactor-cell. Leg: transverse section; Right: front section. 1 gas inlet - 2 heater - 3 cooling - 4 S.S. window holders - 5 brass screw - 6 S.S. screw- 7 FrFE ring - 8 Kalrez O-ring and PTFE back-up tings- 9 CaF2 window- 10 KBr rod - 11 Sample holder - 12 gas outlet - 13 thermocouple- 14 Sample pellet.
minimizes the dead volume (0.12 cm3). Presence of a thermocouple in contact with the sample holder allows the reactor temperature to be precisely monitored. The reactant flow penetrates by the side of the IR cell, licks the catalyst pellet (ca. 50 ~tm thick) on both faces and is evacuated by the other side. The cell behaves like a differential reactor (low conversions) and is inserted in a classical high-pressure flow apparatus directly built on the optical bench (Nicolet Magna IR 560).
271
2.2 Catalysts Two solids were studied: the alumina support (SaET = 210 m2.g1) which was provided by Rh6ne-Poulenc and CoMo/A1203 which was prepared by pore filling impregnation of A1203. Metal contents are 8.3%wt. Mo and 2.2%wt. Co. For the FTIR study, samples were pressed into self-supporting wafers (ca. 5 mg.crn-2, O = 1.6 cm). Reactivity experiments were performed either in the reactor-cell or in a conventional glass micro-reactor and in this case, catalysts were sieved (0.25 -0.31 mm).
2.3 Operating conditions Sulfidation. Activation of the catalysts were the same for FTIR and reactivity experiments. The procedure was the following : catalysts (10-15 mg) were sulfided in situ under 1 Atm. by a mixture of HE/HES (15 vol% HES) with a flow rate of 30 ml.min -1. After introduction of the sulfiding mixture at room temperature, the samples were heated at 400~ within 2 hours and maintained under these conditions for 2.5 hours. Catalysts were further flushed at the same temperature for 1 hour by He (30 ml.min "1) and cooled down under He at the desired temperature. FTIR. Dimethyl-2,6-pyridine (0.5 ~tl) was injected by pulses in the flow, either at 100~ in H2S/I-Ie flow (50 ml.min -1, 2.5 vol% H2S) or at 250~ in the reactant flow (20 ml.min "1, 2.5 vol% H2S). Spectra (256 scans, 4 crn-1 resolution) were recorded on a Nicolet Magna 560 equipped with a MCT detector. The spectra reported below correspond to a mass of 10 mg of A1203 in the sample. Reaction conditions. The reaction was carried out at 250~ the pressure of 3,3-dimethyl-1butene was 38 Torr diluted in the HES/He mixture, total flow rate was 20.0 ml.min "1 . Products analysis was performed by a gas chromatograph (Varian Star 3400 cx) equipped with a capillary column HP KCI/A1203 (50 m, 0.32 mm ID) and a FID detector. Activities are calculated according to the following equation: r = Fo X~ WAUO3 where r is the rate of disappearance of 33DMB 1 (mmol.(gA1203)-l.hl), Fo is the molar flow rate of 33DMB 1 (mmol.h-~), X is the conversion and WAuo3 the mass of alumina contained in the sample (g).
3. R E S U L T S A N D D I S C U S S I O N
3.1 DMP adsorption. Adsorption features Previous FTIR studies [ 10,14-18] reported that the various adsorption modes of DMP with the surface lead to specific spectral features in the 1660-1580 crn~ range (v8a and v8b vibrations). (i) weakly adsorbed species (H-bonded or n-coordinated) give rise to absorption bands at 1602 and 1580 crn-~ (ii) o-coordinated species (DMP-L) on CUS sites 1615-1608 and 1580 cm~ and (iii) protonated species (DMPI-I+ ) absorbs at 1650-1645 and 1625 crn~. On figure 2 are reported the spectra recorded at 100~ during a DMP pulse on A1203 (Fig. 2A) and on CoMo/A1203 (Fig. 2B). Similar features arise on both catalysts. DMP adsorption leads to the appearance of strong bands at 1615, 1602 and 1580 cm-~ and weaker bands at ca.1650 and 1625 crn-~. Within the first minutes of the pulse, bands at ca.1650 and 1625 cm~ (protonated species) strongly increase and further stabilize. Bands corresponding to
272
Figure 2. DMP adsorption at 100~ on (A) A1203and (B) CoMo/A1203. first spectrum (0-1 minute)" m following spectra (1-15 minutes) 9----first spectrum fitting A 9absorbance. coordinated species present a similar evolution. As for the weakly adsorbed species, their amount reached a maximum at ca.3 minutes and further continuously desorbed. Although the same bands are observed on both catalysts, some differences in intensity are observed : bands at ca. 1650 and 1625 cm -1 are more intense on CoMo/A1203 than on A1203 indicating that presence of the sulfide phase increases the number of Bronsted acid sites. As for the band at 1615 cm q (DMP-L), its intensity is higher for A1203 than for CoMo/Al203. This band is characteristic of strong Lewis acid sites of alumina, its lower intensity on CoMo/A1203 is explained by the coverage of the A1203 support by the sulfide phase and also bythe weaker Lewis acid sites present on the sulfide phase compared to those of the support. As for the weakly adsorbed species, their number is higher on CoMo/A1203 than on A1203 . This may be explained either by the presence of weak Bronsted acid sites on the sulfide phase (SH groups) or some weakly coordinated DMP species on the CUS sites of the sulfide phase.
Effect of H2S
When DMP adsorption is performed in presence of a flow of H/S/He, several modifications of the spectra arise. On A1103 the intensity of the ban d at ca. 1615 cm q (v8a DMP-L) strongly decreases whereas the bands at 1650 and 1625 cm~ (v8a andDMPI-I+ DMPH + resp.) increase. Appearance of a supplementary band at ca. 1635 cmq was also noted, likely attributed to weakly pDMPH + ated DMPH + species. The increase of the number of Bronsted acid sites at the expense of Lewis acid sites clearly indicates the transformation of part of Lewis acid-base pairs into Bronsted acid sites by H/S adsorption. As for the weakly adsorbed species
Figure 3. Effect of H2S on the adsorption of DMP at 100~ on "(A) A1203and (B) C o M o / A 1 2 0 3 . . . . before H2S 9 aiter H2S A 9a b s o r b a n c e .
273
their amount is slightly higher than without HES. On CoMo/A1203, the number of protonated species is also higher than previously (without HES) and the number of strongly coordinated species also decreased. At least, the number of created Bronsted acid sites follows the ranking CoMo/A1203 > A1203.
3.2 lsomerization of 3,3-dimethyl-l-butene. Activities The initial activities were measured after 1 minute of reaction and are reported in Table 1. Experiments were reproducible with an accuracy of ca.10-15% for initial activities and ca.5% for steady state activities. In our conditions, only 2,3-dimethyl-l-butene and 2,3dimethyl-2-butene were formed and no secondary reaction products were detected, either in the absence or in the presence of H2S. In the absence of H2S in the flow, CoMo/A1203 was initially more active than A1203 but quickly deactivates. Presence of HES in the flow increased activity of both A1203 and CoMo/A1203. However, in that case, initial activities are difficult to compare taking into account the experimental uncertainty and the very high deactivation rate of CoMo/A1203 in these conditions. At the steady state, A1203 was more active than CoMo/A1203 as observed in the absence of liES. FTIR spectroscopy The reaction was also carried out in the IR reactor-cell, the same products being obtained than previously. Figure 3 shows the spectra recorded during the reaction carried out on CoMo/Al203 in absence of HES. On the spectrum recorded during the first minute of reaction (spectrum b), only bands corresponding to adsorbed 33DMB1 (denoted B on the spectra) are detected. During the following minutes (spectra e), new bands at ca.1670 (sh.), and 1148 cm 1 (denoted X on the spectra) appear and stabilize after ca. 5 minutes of reaction. In parallel, broad bands at 1560 and 1460 cm -1 (with weak shoulders at 1422, 1396 and 1369
in.
Figure 4. a) Spectrum of liquid 33DMB1. b-d: Spectra recorded on CoMo/AI203 in absence of H2S during 33DMB1 isomerization at 250~ b, c, d: difference spectra recorded in the first hour of reaction. In insert, difference spectra : 15-60 minutes spectra minus spectrum recorded at 10 minutes..4: absorbance.
274
crnl ) and a weak band at 1225 c m "l (denoted C on the spectra) slowly grow during the whole duration of the reaction (cf. difference spectra d minus e, in insert). When the catalyst is flushed by He after reaction, bands due to 33DMB 1 (B) strongly decrease whereas bands due to X decrease more slowly. Bands corresponding to C species are poorly affected by the He flush. In the case of A1203, similar features arise except that species X were not detected.
Poisoning by DMP. Once a stable activity is reached, DMP is injected in the feed and its adsorption is followed by FTIR. In all I 0.004 i~ i~ i~ cases DMP adsorption induced a complete deactivation i i of the catalysts (activities fall to zero). Figure 5 shows the spectra recorded at 250~ during DMP adsorption on CoMo/A1203. Interaction with protonated species is clearly evidenced by the formation of bands at 1644 (v8a), 1620 (v8b), and 1271 cm -1 (~5(NH)). Intensity of x these bands decrease with time whereas some new 1300 700 1500 species appears on the surface, characterized by bands at Wavenumbers / cm1 1561 (broad) and 1420 crn-~ (denoted * on the spectra). DMP also displaces adsorbed species as evidenced by Figure 5. DMP adsorption on the negative intensity at 1670 crn-l (denoted X on the CoMo/A1203 after 1 hour of reaction, spectra). Similar experiments on A1203 do not reveal any 250~ 0%H2S transformation of DMP at the surface, and the amount of protonated DMP species stays constant as indicated by the intensity of v8a and v8b. This indicates that the transformation occurs on the sulfided phase. DMP adsorptions performed on activated CoMo/A1203 at temperature as low as 200~ revealed a similar transformation showing that it is not due to a reaction of DMP with the hydrocarbons present during the reaction. 3.3 Discussion Adsorption of DMP in dynamic conditions at 100~ reveals the presence of a higher number of Bronsted acid sites on CoMo/A1203 than on A1203. This is in good agreement with the results previously obtained at room temperature although in that case the sulfidation was carried out in quite different conditions [10]. The ranking obtained is also in accordance with the earlier work of Topsoe et al [9] who used pyridine as probe molecule. In the present case, however, Bronsted acidity of the AlaO3 support is by no means negligible compared to that of CoMo/A1203. This can be attributed to the nature of the probe: since DMP (pKa DMPI-Y= 6.7) is a stronger base than pyridine (pKa PyI-I+= 5.2) it allows a larger range of Bronsted acid sites to be detected. In view of the initial activities of our catalysts in 33DMB 1 isomerization, it can be concluded that DMP is a more efficient probe than pyridine to characterize weakly acidic catalysts such as those which are presently studied. These results are in line with a previous study where Bronsted acidity measured by a test reaction (2-propanol dehydration) on ZrO2-TiOa mixed oxides was better described by DMP than by pyridine [ 17]. The relative amount of Bronsted acid sites created by a flow of H2S is also in good agreement with our previous study and confirms that the sulfided phase plays a major role in H2S dissociation on such catalysts. A parallel appears between test reaction and DMP adsorption results since 33DMB1 isomerization confirms the increase of Bronsted acidity in presence of H2S. Moreover, simultaneous measurements of activities and DMP adsorption followed by FTIR during the
275 reaction at 250~ on A1203 shows that the Bronsted acid sites probed by DMP are the active sites for 33DMB1 isomerization since its adsorption leads to a complete poisoning of the reaction. The case of CoMo/AI203 is more complex since a slow transformation of the probe molecule occurs on the sulfide phase. However, only DMP species are observed at the very begining of the adsorption process and since 33DMB1 isomerization is specific of Bronsted acid sites [11 ], it is supposed that DMP transforms on the acidic - S H groups of the sulfide phase and that the resulting species remain adsorbed on these sites. Following this hypothesis, the conclusion for CoMo/A1203 is the same than for A1203 : DMP doses the active sites in 33DMB1 isomerization. Therefore, comparison between the catalytic activities obtained in 33DMB 1 isomerization and the 'number' of Bronsted acid sites detected in similar conditions should allows us to determine a relative activity of these sites in a facile Bronsted acidcatalyzed reaction. Table 1. HaS / Vol% A1203 CoMo/A1203
0 2.5 0 2.5
ro / mmol.h-l.(g A1203) "1 24.0 52.0 32.5 44.3
+ + • +
2.4 7.8 3.2 6.6
AoMart+ / c m "1. (g A1203) -1 27.5 125.0 54.1 220.3
ro / ADMPH+ / mmol.hl.cm 0.87 0.41 0.60 0.20
+ • • •
0.09 0.06 0.06 0.03
Table 1 reports the initial activities (r0) measured at 250~ in presence or in absence of H2S and the corresponding areas (ADMPH+) of the v8a and v8b vibration bands of DMPI-I + species measured at 100~ A measure of the relative activities of the Bronsted acid sites present on the catalysts can be given by the ratios: ro/ADMPH+ 9It Can be concluded from the variation of these ratios that the Bronsted acid sites on A1203 are, o n a v e r a g e , more active than those which are present on CoMo/A1203 and that on both catalysts the Bronsted acid sites created by H2S dissociation are less active than the Bronsted acid sites initially present. No specific information, however, can be obtained about the Bronsted acid sites present on the sulfide phase and on the support due to (i) the difficulty to measure precisely the coverage of the support by the sulfide phase and (ii) the possible influence of the incorporated elements on the acidity of the support (e.g. the sulfide phase, Co atoms inserted in octaedral position of the A1203 lattice...). Although CoMo/AI203 is initially more active than A1203, it quickly deactivates during the first minutes of reaction and further stabilizes. The spectra recorded within the first minutes of reaction on CoMo/A1203 show that besides 33DMB1 adsorption an unknown species (X) characterized by bands at 1670 and 1158 cm l specifically appears on this catalyst at the beginning of the reaction. It is proposed that 33DMB 1 transforms into X on the sulfided phase since X is not formed on A1203. Furthermore, X is retained a longer time than 33DMB 1 when the catalyst is flushed by He indicating that it is more strongly adsorbed than 33DMB 1. Since the amount of X increases during the deactivation of CoMo/AI203, it can be concluded that the strong deactivation of CoMo/A1203 is due to a 'strong' adsorption of X on the Bronsted acid sites of CoMo/A1203. Besides the strong initial deactivation of catalysts, a slight decrease in activity can also be noted all along the course of the reaction. This can be attributed to the 'slow' formation of species C. Bands at ca. 1560 and 1460 cm ~ can be attributed to aromatic v(C-C) vibrations and bands at 1422, 1396, 1369 and 1225 cm ~ to 6(C-H) vibrations. It is proposed that C are hydrogenated' coke' species.
276
4. CONCLUSION The main results of the present study, carried out in dynamic conditions at high temperature may be summarized as follows: - The number of Bronsted acid sites detected by DMP was found higher on a CoMo/AI203 catalyst than on A1203 and the number of Bronsted acid sites created by H2S follows the same ranking. These results are in agreement with our previous work conducted at room temperature in static conditions. - DMP doses the active sites in 33DMB1 isomerization. - The initial activities of A1203 and CoMo/A1203 in 33DMB1 isomerization in the absence of H2S are in good agreement with the number of Bronsted acid sites detected by DMP. The average intrinsic activity of these sites, however, is lower on CoMo/Al203 than on A1203. Presence of H2S in the feed promotes 33DMB 1 isomerization on both catalysts but the intrinsic activity of the created acid sites is lower than the activity of the Bronsted acid sites initially present on these catalysts. - The strong and rapid deactivation of CoMo/AI203 is attributed to the formation of an unknown X species on the sulfide phase which poisons the Bronsted acid sites of this catalyst. A slower and continuous deactivation is attributed to coke formation. In conclusion, this study emphasizes the necessity in characterizing the acidity of hydrotreating catalysts in conditions close to the reaction conditions (high temperature, in presence of H2S) and points out the complexity of sulfided catalysts since not only the surface sites of the sulfided phase but also those of the support depend of the surrounding atmosphere The other major difficulty of sulfided catalyst characterization is that these catalysts are also intrinsically active for the decomposition of nitrogenated molecules as those used as probe molecules. Their use, however, presents the advantage to really characterize the active sites of
hydrotreating catalysts.
Acknowledgements The authors wish to thank Prof. G. P6rot for catalysts supplying. The help from Dr. J. Saussey, Mr. D. Jeanne and Mr. P. Masset in the design of the IR cell is gratefully acknowledged. The authors are indebted to Dr J.L. Lemberton, Dr J. van Gestel and Dr. J.C. Lavalley for fruitful discussions. 5. R E F E R E N C E S
1 Yang S.H., Satterfield C.N., J. Catal., 81,335 (1983) 2 P6rot G., Catal. Today, 10, 447 (1991) 3 Van Gestel J., Finot L., Leglise J., Duchet J.C, Bull. Soc. Chim. Belg., 4-5, 189 (1995). 40lguin Orozco E., Vrinat M., Applied Catal. A, 170, 195 (1997) 5 Kasztelan S. and Guillaume D., Ind. Eng. Chem. Res., 33,203 (1994) 6 Lacroix M., Yuan S., Breysse M., Dor6mieux-Morin C., Fraissard J., J. Catal., 138 (1992) 409. 7 Jobic H., Clugnet C., Lacroix M., Yuan S., Miradotos C., Breysse M., J. Am. Chem. Soc. 115 (1993) 3654. 8 Topsoe N.Y., Topsoe H., Massoth F.E., J. Catal., 119, 252 (1989) 9 Topsoe N.Y., Topsoe H., J. Catal., 139, 641 (1993) 10 Petit C., Maug6 F., Lavalley J.C., Stud. Surf. Sci., 106, 157 (1997). 11 Irvine E.A., John C.S., Kemball C., Pearman, A.J., Day M.A., Sampson R.J., J. Catal., 61, 326 (1980).
277
12 Bourdillon G., Gueguen C., Guisnet M., Applied Catal., 61,123 (1990) 13 Joly J.F., Zanier-Szydlowski N., Colin S., Raatz S., Saussey J., Lavalley J.C., Catal. Today, 9, 31 (1991). 14 Jacobs P.A., Heylen C.F., J. Catal., 34, 267 (1974) 15 Mattulewicz E.RA., Kerkhof F.P.J.M., Mouljin L.A. and Reitsma H.J., J. Colloid. Interface Chem., 77, 110 (1980). 16 Corma A., Rodella C. and Fomest V., J. Catal., 88, 374 (1984) 17 Lahousse C., Aboulayt A., Maug6 F., Bachelier J., Lavalley, J.C., J. Mol. Catal., 84, 283 (1993) 18 Jolly S., Saussey J., Lavalley J.C., Zanier N., Benazzi A, Joly J.F., Ber. Bunsenges. Phys. Chem., 97, 427 (1981)
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.FromentandP. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
279
P h y s i c o c h e m i c a l c h a r a c t e r i s a t i o n o f V G O MHCK c a t a l y s t s a n d its e x t r a p o l a t i o n to c a t a l y t i c a c t i v i t y M. M. R a m i r e z de A g u d e l o , E. Mujica a n d J. A. S a l a z a r
PDVSA I N T E V E P , Apdo. 76343, Caracas 1070 a, Venezuela.
Abstract The mild hydrocracking (MHCK) is a very versatile process where a wide catalytic functionality is needed. We will try to predict the application related to the activity for the main functions: hydrodesulfurization (HDS), hydrodenitrogenation (HDN), hydrocracking (HCK), hydrogenation (HYD) and hydrogenolysis (HYDL), using different criteria derived from their physicochemical properties. The physicochemical properties measured on the 11 analyzed catalysts were the bulk and surface chemical composition, physical properties (surface area and pore volume), the surface acidity, the crystalline phases and the surface parameters (metallic dispersion and sulfiding degree). Different correlations have been established among the surface properties and the catalytic activity.
1. I N T R O D U C T I O N In principle, an MHCK operation is technically feasible in most VGO hydrotreaters. However the impact of process variables on performance has to be determined. The increase on severity, in order to achieve the desired conversion levels, highly impacts on catalyst stability by decreasing the life cycle [1].The typical life cycle of an HDT catalyst (18-24 months) would be drastically reduced under MHCK conditions, particularly due to coke formation, induced by the temperature increase required for the conversion needs. Traditional alumina supported catalysts, either CoMo or NiMo, will not exhibit any difference on distillate yield selectivity. Another alternative to increase conversion is to include an acid functionality in the catalyst. A change on support surface characteristics will certainly affect the distribution of active sites and probably their chemical nature. Catalyst selection is based not only on their conversion levels but also on the HDS activity required to achieve the specifications for low sulfur fuels. The knowledge of the effect of catalyst properties on catalyst functionality would be useful in deciding the most appropriate system for a given objective. It is the aim of the present work to discuss the physicochemical properties of commercial catalysts, which are believed to govern the catalytic functionality for the different reactions observed during an MHCK operation.
280
2. E X P E R I M E N T A L Samples of 11 commercial catalysts were characterized by chemical composition (ICP and atomic absorption), physical properties (BET, XRD ) and physicochemical properties. Catalysts were sulfided with 30 cc/min, of a 10:1 H2:H2S mixture at 350~ for 2 hours. The sulfiding degree (a) was gravimetrically measured, using the oxide catalyst weight (Wo) and its weight after sulfiding (Ws) as: W s
--
W ~
a =~.100
(1)
W o
The acidity of the sulfided catalysts was evaluated as the irreversibly NH3 adsorbed (pmol/m2). Firstly, the catalyst is saturated with a 1:1 He: NH3 mixture at 30~ 1 atm and 30 cc/min. Then, the sample is outgassed under a He flow, at different temperatures. The difference between the acidity at 100 and 400~ is taken as a measure of the inverse of the width of the curve of acid strength distribution (ASD) and the value of the acidity at 100~ as the total acidity. The XPS spectra of sulfided catalysts were recorded in a Leybold-Heraerus (LHS-11) using Mg Ka (1253.6 eV) as excitation source. The atomic surface composition was evaluated using calibrated sensitivity factors. The catalytic activity for VGO HCK was taken as the temperature for achieving 20% conversion of the 370~ + fraction and for HDS, as the temperature for 95% removal of sulfur compounds. The VGO was treated at 700 psig, LHSV = 0.5 m3/m3h and initial activity was measured after 24 h run.
3. R E S U L T S A N D D I S C U S S I O N Table 1 shows the chemical composition and physical properties of the considered catalysts. The total metal content and the Mo metallic atomic ratio, r, (Mo/Mo +Me) varies between 9.5 and 16.4 and 0.61-0.88, respectively. The Mo metallic ratio has been associated with a synergetic effect on catalytic activity. It has been shown to be optimized at different values for HDS t h a n for HYD. Some differences has been also observed for the CoMo type catalysts, compared to the NiMo type. Most of the catalysts subject of the present study are of NiMo type and contain an acidity promoter, such as F or P, or are supported on a silicoaluminate material. The XRD patterns did not show evidence for zeolite presence. The brownish - beige color of the NiMo/SIO2-A1203 supported catalysts indicates a major presence of nickel oxide rather than the typical NiA1204 green phase present in A1203 supported catalysts, which might also be a further evidence of the presence of the SIO2-A1203 mixed oxide. As can be seen, the SiO2A1203 supported catalysts showed higher surface area values t h a n the others.
281
Assuming both, to have a random distribution of the A1 atoms and a strong metal-support interaction between the metal and the A1 atoms, a high metal dispersion might be expected. Consequently, the reactivity of such highly dispersed species towards sulfiding might be different of t h a t found on A1203 supported catalysts and so the sulfiding degree might also be different. Table 1. Chemical composition, Surface Area (SA) [m2g -1] and Pore Volume (Vp) [cc g-1 ]. Catalyst Chemical Composition (wt%) SA Vp MoO3 NiO Co03 P 2 0 3 A1203 SiO2 Fr A 10.7 4.3 51.4 39.0 0.68 293 0.35 B 9.2 4.3 77.4 2.49 0.64 163 0.31 C 15.0 8.2 58.3 19.6 0.61 239 0.53 D 16.5 3.1 6.5 66.3 0.82 150 0.37 E 19.5 2.2 48.8 24.4 0.88 183 0.51 F 13.8 5.7 35.0 41.2 0.67 157 0.44 G 15.5 4.9 8.5 70.0 0.74 190 0.54 H 12.6 3.2 48.6 25.2 0.77 311 0.35 I 12.4 2.5 85.1 0.76 120 0.22 J 16.0 4.3 79.7 0.76 138 0.45 K 15.8 3.9 80.3 0.70 276 0.70 The acidity values are collected on table 2. The total acidity might be considered as the independent contribution of the sulfided phase and that of the support. However, it is reasonable to think that the total metal content could be related to a balance between acidity and hydrogenation activity in order to avoid coke formation and increase catalyst stability. If that is so, a higher total metal content would be expected for those catalysts with larger total acidity values. In Figure 1, the variation of total acidity with total metal content is presented and as it was anticipated a correlation was found. The total acidity of the catalyst (AcidityT) might be due to the independent contributions of the support (Aciditysop) and the (sulfided) metal phase (AcidityMeT), i.e.: AcidityT = Aciditysop + AcidityMeT The contribution of the support is given by a fraction of its original acidity (Aciditysop~ which is modified during impregnation, so: Aciditysop = Aciditysop o (1-klMew) and the acidity of the metal phase might be proportional to the surface metal concentration (MET), i.e. AcidityMew = k2Mew, then, Acidityw = Aciditysop o + MeT (k2- kl Aciditysop ~ According to this equation, if all the samples contain the same support, the intercept represents a measure of the support acidity and the k2 constant should be associated to the intrinsic acidity of the sulfided metal sites. Thus, it should be
282
affected by the stoichiometry of the sulfiding reaction (metal type, metal-support interaction, etc.), number and nature of species, etc. Table 2. Irreversible Acidity [~mO1NH3m -2~ 100~
300~
400~
ASD
A
7.79
0.00
0.00
7.79
B
6.76
2.47
0.00
6.76
C
8.14
1.01
0.00
8.14
D
7.92
3.43
3.03
4.89
E
12.80
3.63
0.00
12.80
F
6.57
0.84
0.00
6.57
G
4.33
1.55
0.88
3.45
H
5.98
1.18
0.00
5.98
I
2.28
0.39
0.00
2.38
J
11.90
2.57
0.00
11.90
K
6.70
0.50
0.00
6.70
F i g u r e 1: Acidity vs. Total metal content and Sulfiding degree vs. ASD
In the fluoride containing catalyst (B), strong acid sites were supposed to be present, but instead the acidity value at 400~ is nil. The F- effect on the alumina is well known to be due to the electron withdrawing character of this anion [2], which will induce the formation of strong Bronsted acid sites and destroy the amphotheric character of alumina by neutralizing the basic sites [3]. Thus, during impregnation of metals, their interaction with the strong acid sites should have taken place, which in turn will result in a better dispersed catalyst compared to a bare alumina supported one. The NiMo/SiO2-A1203 exhibited a much higher acidity than that of the NiMo/A1203 catalyst (ex. I) which has been included here for comparison purposes. The sulfided NiMo phase does not show strong acid sites when supported on A1203. Instead, acid sites are rather weak or moderate at the most, probably due to the low surface metal concentration observed on those catalysts. On the SIO2-A1203 supported catalysts, we observed that acidity of the NiMo sulfided phase type of acid sites (weak and moderate) and also a considerable increase of the total acidity. Two factors contribute to the same effect, one is the increase in dispersion (and so in the surface metallic sites, acidic in nature) and the other the support contribution. It is known that increasing the SiOe content in the mixed oxide (SIO2-A1203) increases total acidity, acidity strength and Bronsted sites [4]. Assuming the -OH- bridge Mo interaction [5] during impregnation, the Mo species will rather interact with those strong -OH- sites, neutralizing them, as mentioned before.
283
Simultaneously, and depending on sulfiding degree new acid sites will be created, namely the sulfided sites. We might then propose several hypothesis: 1. The higher the metal content, the higher the extend of neutralization of the support acid sites and consequently the higher the contribution of the sulfided phase to acidity would be. 2. The higher the dispersion, the higher the sulfiding degree and consequently the larger the total acidity would be. The higher the sulfiding degree, the lower the strong acidity and consequently the larger the ASD. In Figure 1 a confirmation of this hypothesis is shown. The first part of the second hypothesis is confirmed in Figure 2, namely the correlation between sulfiding degree (measured by XPS) and dispersion (also taken from XPS measurements as Me/A1 +Me) together with the second part of it, i.e. the variation of ASD with dispersion. In order to avoid the influence of the chemical nature of the catalysts, only SIO2-A1203 supported catalyst were considered.
F i g u r e 2. XPS sulfiding degree and F i g u r e 3. Sulfiding degree (bulk vs ASD vs. Total dispersion surface) and Metallic ratios: rs = Surface, rB -- Bulk The surface derived parameters are collected in Table 3. In the case of catalyst B, we have mentioned that the electron withdrawing character of Fmight affect the surface properties of the A1203 support, giving rise to a strong metal-support interaction, which would lead to higher dispersion. The evaluated dispersion (Table 3) for the B catalyst is, indeed, almost twice the value obtained for the NiMo/A1203 catalyst, I. We have also mentioned, that when a high dispersion is governed by a strong metal-support interaction, one might arise on a poor sulfiding caused by a weaker metal-sulfur interaction. In fact, the XPS degree of sulfiding, taken as the
284
S/MeT ratio is three times lower in the F-promoted catalyst B as compared to the catalyst I. Whil, the XPS sulfiding degree might be thought as the sulfiding stoichiometry, the gravimetric sulfiding degree is related to the sulfiding extension instead. However, a linear correlation is found between the surface parameter and the gravimetric sulfiding degree, Figure 1. A large deviation was exhibited by catalyst D, for which either the extension of sulfiding did not go up to completion or the surface S/Me ratio was too low. For alumina supported catalysts the presence of the metal spinel would contribute to higher detected values of A1 and Me surface concentrations and differences between the gravimetric and surface sulfiding degree might be expected. The values for the total dispersion (DMo + DMe) of the SIO2-A1203 supported catalyst are higher than those typically observed on alumina based catalysts (see catalyst I as reference). We already introduced the idea of this fact being due to differences in the Me-A1 and the Me-Si interactions and considered the effect of increasing amounts of SiO2 in the support. However, at high level of SiO2 the acidity begin to decrease and so the acid strength. At low SiO2 content, an increase in SiO2 will increase metal dispersion but at high SiO2 contents, the lower the total acidity and the acid strength will cause a decrease in dispersion, see Figure 4. Table 3. XPS Derived Parameters. DMo
DMe
S / M e T S/Mo
rMo
B
5,23
5,68
0,46
0,97
0,48
C
11,07
6,85
0,71
1,14
0,62
D
ND*
ND*
1,01
1,14
0,62
E
13,96
11,42
1,49
2,71
0,55
F
7,66
8,10
0,58
1,19
0,49
G
6,33
3,06
0,76
1,13
0,67
H
4,37
3,39
0,62
1,10
0,56
1,54
1,54
2,16
0,71
I 3,72 * Not detectable
15
01
0
I 19.6
24.4
39
41.2
%Si02
F i g u r e 4. Dispersion variation
Figure 3 shows the variation of the surface metallic ratio with the bulk metallic ratio, as well. We have included some previous results [7] for comparison purposes. Most of the catalysts considered in this study shift away from the straight line. For conventional HDT catalysts, the surface ratio closely resemble that of the bulk. For the SIO2-A1203 catalysts with no spinel formation, the surface concentration of the promoter metals is much higher than that observed
285
for alumina supported catalysts. The initial studies on the effect of the atomic metallic ratios on catalytic activity [8] were intended for searching the optimal value, they were based on alumina type catalysts and on bulk ratios r a t h e r t h a n on surface ratios. As we have seen, for alumina based catalyst the bulk ratio is similar to that of the surface, and the optimal synergistic ratio was determined. In view of the results shown in Figure2, any effect derived from the metallic ratio must be based on surface ratios and not on bulk ratios. Regarding HCK, it has been postulated, that if conversion is oriented towards gasoline the requirements are a strong acidity balanced to a moderate hydrogenation. The opposite situation is proposed for distillates production [9]. For the considered catalysts moderate acidity was found for all of them and a high HYD activity would be expected for those exhibiting the higher dispersion (C, D and E catalysts, for instance), but also at the adequate metallic ratio. However, we have seen how dispersion affect the sulfiding degree, which in turn will also influence the HDS and ASD. High HDS is expected for moderate XPS sulfiding degree (maximizing the number of anionic vacancies, C, D, E, F and H) but together with high metal dispersion (C, D and E), at surface ratios between 0.5 and 0.8 (C, D, E, F, G and H). A high HDN for non-basic compounds might need only poor HYD activity. Meanwhile, a high basic HDN corresponds with a high HYDL. A large number of acid sites might favor the adsorption step of the nitrogen containing compounds (B, D, E, G and J). However, the need for hydrogenation and hydrogenolysis capacities makes the selection difficult. Both, poor HYD or high HDN, would be the least stable catalyst under HCK conditions. In summary, the MHCK catalysts (even for HCK catalysts) should possess a wide functionality, which can be achieved by introducing an acidic support and widen up the ASD of the sulfided phase, for which the metal dispersion and the surface metallic ratio must be optimized towards HYD and HDS. Assuming that the estimated optimal metallic ratios determined for alumina catalysts are valid only because the existing correspondence between the surface and bulk ratios, one might extrapolate those optimal values to compare the ones found for the surface ratios of the SIO2-A1203 catalysts and, together with the preceding discussion, predict their catalytic behavior. In that case the best catalyst are as follows:
HDS: C, E
HDN: D, G HYD: C, E, H
HCK: C, E
In Figure 5 this prediction has been confirmed for HDS and HCK, so we are inclined to believe that it would be valid also for the rest of the function. Regarding HCK, we might have included the D catalyst, however its behavior on sulfiding might imply a lack of stability under MHCK conditions.
286
F i g u r e 5. Relative catalytic activity
4. C O N C L U S I O N S The knowledge of structure - reactivity correlation will certainly impact catalyst development. We have shown here the use of several physicochemical properties in trying to establish such correlations. Still, there have to be carried out further work for the search of a better knowledge of catalyst structure. A striking fact is the complexity of the mutual effect among the different properties considered here. Although, the variety of studied catalyst used different support materials and most likely were prepared by different methods, they still hold a common characteristic, i.e. they work for the same purpose under the same conditions, and so the same active sites are on their surface. References 1. J. P. van der Berg, J. P. Lucien, G. Germaine, G. L. B. Thielemans. Fuels Proc. Technol. 35, 119. 1993. 2. A. Corma, V. Fornes, E. Ortega. J. Catal. 92, 284. 1985. 3. E. A. Paukstis, P. I. Soltanov, E. N.Yunchenko, K. Jiratova. Coll. Czech. Chem. Comm. 47, 2044. 1982. 4. K. Tanabe. Solid Acid and Bases, Academic Press, NY 1970. 5. J . M . Lipsch, G.C.A. Schuit. J. Catal. 15 (2), 174. 1969. 6. M. P. Seah, W. A. Dench. Surf. Inter. Anal. 1 (1), 2. 1979. 7. M. M. Ramirez de Agudelo, Unpublished results. 8. Delmon. Proc. Intern. Conf. Chem. Uses Molyb. 4th. Golden, USA, 9-13 August. 1982. 9. W. Scott, A. G. Bridge. Adv. Chem. Ser. 103, 113. 1971.
REACTOR MODELING
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
Modelling a Hydroconversion Computational
Fluid Dynamics
289
Reactor Based
on a
Approach
M o n t s e r r a t Motas Carbonell ~ and Reginaldo Guirardello b Department of Chemical Engineering, State University of Campinas - UNICAMP. CP.6066, CEP 13083-970. Sao Paulo, Campinas, Brazil. a [email protected] b [email protected]
Abstract This paper deals with the modelling of a hydroconversion reactor, operating in a slurry bubble column mode, for the upgrade of distillation residues to lighter products. The model is based on the mass and momentum balance time-averaged equations, considering the radial dispersion of components. Simulations are carried out in two steps: first fluid dynamic variables (velocities, hold-ups and pressure drop) are obtained and then the off cracking conversion is determined showing the distribution of products and the concentration profiles in the reactor. Results show good agreement with physical behaviour. Keywords: Bubble column; slurry reactor; hydroconversion process; residue upgrading. 1. I N T R O D U C T I O N The scenario of the oil market shows an increasing need for refining petroleum with maximum yields in high-quality light products (especially naphtha and middle distillates). The level of heteroatom content is an important factor to indicate the quality of products. Recently, the stringent environmental regulations have forced the levels of contaminants allowed in the products to be reduced. In this context, the upgrading of heavy oil fractions efficiently is a major goal and the hydroconversion process is an interesting alternative to accomplish this goal (Trambouze, 1991; Eccles, 1993). In the hydroconversion reactor, the heavy oil is submitted to severe temperature (400-500 ~ and pressure (10-25 MPa) operating conditions in the presence of hydrogen and catalyst (or additive). The product obtained is composed of lighter oil fractions, with increased hydrogen content and lower level of contaminants (sulphur, nitrogen and metals). The hydroconversion technologies may be divided, according to the reactoroperating mode, in two main groups: the ebullated-bed and the slurry bubble column (Fan, 1989). The particles are fluidised in the reactor by the upward flow of oil (continuous phase) and gas (mainly hydrogen, dispersed as small bubbles). In the slurry systems, the particles are introduced and removed from the reactor mixed with the liquid (as slurry). The ebullated bed reactor operates with relatively larger particles that are charged and discharged independently of the liquid flow. The former reactors usually have no internals, while the latter reactors are provided with an internal recycle tube.
290
This paper considers the modelling of a slurry bubble column reactor for the hydroconversion of distillation residues. The model is based on a fluid dynamics approach and accounts for the thermal cracking reactions in the reactor. The simulations are carried out in two steps: first the fluid dynamic variabels are determined and then the oil cracking conversion is calculated. 2. MODELLING According to Tarmy and Coulaloglou (1992), "gas-hquid-solid reactors may be the most difficult of all reacting systems to commercialise". This is due to the complexity involved in the description of their behaviour that involves virtually all facets of reaction engineering, where a major difficulty is the understanding of the fluid dynamics phenomena. Indeed, recently a significant progress has been observed in the modelling of the fluid dynamics for bubble columns. Delnoij et al. (1997) briefly reviews these advances and presents a '"hierarchy of models" concept that consists in employing the most appropriate CFD model for the problem at hand. The problem is characterised by the size of the column, the length scale of the phenomena of interest and the prevailing flow regime. The CFD model may be the EulerianEulerian two-fluid model, the Eulerian-Lagrangian discrete bubble model and the Volume Tracking or the Marker Particle model. These models can describe the complex, time-dependent, flow structures observed in bubble columns. However, most of those work was developed under "ideal conditions", that is, in systems using air and water, at ambient temperature and pressure. Although these information do provide bases for assessing the fluid dynamics in the severe hydroconversion operating conditions, other factors, such as the effect of pressure, have to be considered. Recent reports (Luo et al., 1997) have shown that increasing pressure decreases bubble size and bubble rise velocity, leading to an increase in gas hold up. This effect of pressure on gas hold up was earlier observed by Tarmy et al. (1984) during the development and scale-up of a slurry bubble column reactor for the coal liquefaction process. According to Fan (1989), under the operating conditions of the hydroconversion and coal liquefaction process, the gas phase behaviour in these reactors bear certain similarities. Therefore, the transport properties that were reported by Tarmy et al. (1984) might provide invaluable information to the development and scale-up of hydroconversion reactors. Most of the models reported for the hydroconversion reactor are still quite simple and consider the liquid phase completely back-mixed (Mosby et al., 1986; Eccles, 1993; Abreu, 1997) or an axial dispersion in the liquid phase (Abreu, 1997). Kam et al. (1995) presented a model based on artificial neural networks as an alternative to deal with the complex behaviour of ebullated-bed. That model was trained with plant operation and laboratory research data in order to be used as an operational support tool for the process. The model developed and presented in this paper uses a fluid dynamic approach and takes into account the thermal cracking reactions of the oil in the reactor. These reactions are followed by hydrogenation of the cracked compounds. The oil and particles are treated as one pseudo-homogeneous phase (slurry) in which the sohds are uniformly distributed. This is in agreement with the recirculation pattern observed in bubble columns, with an upward flow of liquid in the center and
291
downward flow near the wall. It is considered that the gas is composed of hydrogen that flows in a homogeneous bubble regime. This assumption is in agreement with the observations of Tarmy et al. (1984) and Luo et al. (1997) regarding the behaviour of the bubbles in high pressure systems. The model is based on the mass and momentum time averaged balance equations using an Eulerian approach. The main assumption is a fully developed turbulent axial flow with azimuthal symmetry. It is also considered that the reactor is isothermal and operates under steady state conditions. The thermal cracking reactions are modelled using the lumped kinetic scheme proposed by Mosby et al. (1986), which divides the oil into lumps according to the reactivity and/or boiling point range of components. Therefore, the feed is divided in residue ('%ard" and "easy" to react) and gasofl, while the products are divided into gasoil, distillate, naphtha and gases. The model equations are obtained through the mass balance, for each reaction lump, considering both the convection and the radial dispersion phenomena. The axial dispersion phenomenon was not considered since the analysis of magnitude order showed that convection is the dominant equation term in the axial direction. The model equations written in cylindrical coordinates are presented as follows: Momentum Balance for slurry and gas: eft dVsl "~ dP 1 d @sl "r'gsl " )-~;sl " r dr dr
_.
g'Psl + 5 0 0 0 0 . e g
(1)
-esl .(Vg - V s l ) - O
1 d (eg "r'Beff dVg) dP r" dr g " dr - Sg "-~Z- g ' S g " (pg - P s l ) - 50000 "Sg-esl "(Vg - Vsl) - 0 Liquid continuity equation:
U L
97C"Rc 2.
PL
=
2- ~ -p
Rc L " ~ EL
" VL
"r.
dr
(2) (3)
0 Mass Balance equations for each lump i: 1 0 ( r eft OCi) EL 9 .VL OCi c3z r Ork ~t/s -EL c3r J k - 9.0907 1013 .exp
I - 2.4282105 1 8-~i41T
7 - ai "k'eL'PL .Ci + Z 13m,i'k'eL "PL "Cm m=l
(4) (5)
where c~i and ~m,i are constants obtained from the kinetic scheme of Mosby et al. (1986). The boundary conditions for the equations are presented in Fig. 1. The solution of the equations requires the input of the radial distribution for the hold up (s) and effective viscosity (ge~) of gas and slurry, calculated with the aid of constitutive equations. The mean gas hold up is determined with the expression proposed by Tarmy et al. (1984), while the liquid turbulent viscosity is calculated with a zero order turbulence model (Chen et al., 1985). These equations are detailed in Carbonell and Guirardello (1997) where a preliminary version of this model is presented. The main improvement in the model presented here is the consideration of the radial dispersion of components, which leads to more realistic behaviour regarding the distribution of components m the reactor, as shown in the results.
292
3. S I M U L A T I O N It is assumed that thermal cracking reactions do not interfere in the fluid dynamics due to the high back mixing of the liquid phase and the low gas absorption and formation rate. This assumption was numerically tested considering the individual properties of residue and distillate (Carbonell and Guirardello, 1997). Therefore, simulations may be divided in two consecutive steps. First, the fluid dynamic behaviour is determined giving the pressure drop in the bed (dP/dz) and the radial distribution of gas and slurry for the hold up, effective viscosity (dispersion) and velocity (Vs1, which is equal to VL). Then, the thermal cracking reactions are simulated leading to the determination of the off conversion (X) and the radial and axial distribution for the concentration of each lump (Ci). 3.1 S o l u t i o n of the e q u a t i o n s o f the m o d e l The numerical procedure used to solve Eq. 1, 2 and 3 that represent the fluid dynamic behaviour is discussed by Carbonell and Guirardello (1997). The reaction model (Eq. 4) is a system of seven partial parabolic differential equations (one for each lump) that was solved using the method of lines. The equations were discretized in the radial direction using the orthogonal collocation method with 10 collocation points (see Fig. 1). The system was then reduced to fifty six ordinary differential equations to be integrated in the axial direction. The first approach to integrate these equations considered the use of numerical methods. However, due to the high numerical stiffness of the system of differential equations, either the results diverged or the computational time was too long. Then, an analytical approach based on the determination of eigenvalues and eingenvectors was proposed and successfully implemented. The main aspects of this procedure are briefly discussed here. The model equations were divided in 7 sub-systems (one for each lump, each composed of 8 equations) that were solved independently in consecutive steps. First, the three reactant systems were solved. The analysis of the magnitude order showed that only the two last terms of the solution are significant, so that the others might be neglected. Then, the other four product systems were solved with the use of the method of undetermined coefficients for the particular solution. The equations obtained are presented as follows: 9 For "hard" and "easy" to react residue and gasoil in the feed ( i = 1; i = 2 ; i = 3 ): 8 Ci(rj, z) = Z ~ij,k. wik k=7
(6)
9 For products gasoil, distillate, n a p h t h a and gases ( i = 4; i -- 5; i - 6; i - 7 ): 8 Ci(rj,z) = Z (~ij,k" wik + bilj,k" Wlk + bi2j,k" W2k + bi3j,k" W3k + bi4j,k" W4k) k=7
(7)
and
(8)
wi k - ai k 9e ~ik:z
where: )~ikare the eingenvalues for lump i; ~i j,k are the eigenvectors for )~ik ; bil, bi2, bi3 and bi4 are the constants of the particular solution, where b43=0 and b44=0; and ai are the constants determined using the boundary conditions.
293
3.2 Reaction-Recirculation Algorithm Simulation of the reactions requires the previous knowledge of the fluid dynamic variables (velocity, hold up and effective viscosity) used in the determination of the eigenvalues and eigenvectors. These variables are calculated in the region of fully developed flow. Therefore, it was necessary to develop a procedure to correlate these information and the concentration of lumps in this region with the entrance and outlet streams of the reactor. The reactor is divided in three regions (see Fig 1.): the entrance, the outlet and the fully developed flow. Since entrance and outlet regions are much smaller, the mass balance due to reactions can be neglected and a complete back-mixed behaviour is assumed for the liquid phase in these regions. Mass balance for each lump i lead to the equations t h a t correlate Col and Csi with the concentration in the developed flow region (Carbonell and Guirardello, 1997). It was observed the occurrence of high eingenvalues for all lumps. In order to avoid overflow during simulations due to the exponential in wi, the following strategy was implemented: 9 for ~i,k < 0, wi is written as:
wi k = ai~ 9e ~'ik'z
9 for )Li,k > 0, wi is written as:
wi k = ai Lc . e ~ik'(z-Lc)
(9) (10)
Finally, the constants ai are determined with the application of the boundary conditions in z direction. This calculation is accomplished by writing the model equations (Eq. 6 and 7) for both limits of the developed flow region (at z = 0 and at z = Lc). The linear system in ai obtained is then solved together with the mass balance equations for the entrance and outlet regions, leading to the determination of ai, Ci(r~,z=0) and Ci(rj,z=Lc). The program was implemented in Fortran language (Fortran 90) and the computational time for each simulation is approximately 18 s running in a PCPentium- 133 Mhz.
294
4. RESULTS Simulations of the reactor in pilot plant scale were carried out in typical operating conditions for the hydroconversion process. The reactor is fed with a vacuum distillation residue composed of 80 % "hard" to react (mainly asphaltenes and resins) to 20% "easy" to react residue (otis). Liquid phase properties were obtained in a commercial simulator. The simulation of the fluid dynamics with the one-dimensional two-fluid model gives the pressure drop in the bed (dP/dz) and the radial distribution for the velocities, hold ups and effective viscosities. The slurry velocity profile shows the expected recirculation behaviour with the upward flow in the centre and downward flow of slurry near the walls. According to Sokolichin and Eigenberger (1994), this well known behaviour is the result of a long-time averaging of the dynamic picture with rising vortices that are generated sequentially at the entrance region of the bubble columns. The average gas hold ups obtained with the model are in agreement with the data used by Eccles (1993) in the modelling of residue hydroprocessing using ebullated bed reactors. The simulation of the thermal cracking reactions gives the residue conversion and the radial and axial distribution for the concentration of each lump i . Fig 2 and Fig 3 show the concentration profiles for "hard" to react residue (C 1) and distillate (C5), which represent the typical results obtained for reactants and products. The reactants have a decreasing concentration along the reactor while the products show the opposite trend. The concentration profiles vary more intensely in the axial than in the radial direction, as expected. Due to the effect of the entrance stream, the radial distribution of compounds is more pronounced at the bottom and flatter at the top of the reactor (Fig. 2 uses a constant ACi scale for all axial positions).
Figure 2. Reactor diagram showing the radial concentration profiles at specified axial positions (T=450~ ~L=I 7 min). Symbols indicate the calculated values.
295 0.75 The results obtained with this i }o18 model for a wide range of .......... ~........................................ [ ~ 0 16 0.70 operating conditions were t.o ......... ~, c~ ..... o.14 compared to the results given 0.65 with a model t h a t neglects the radial dispersion phenomenon 0.60 (Carbonell and Guirardello, 1997). ......... i....................................... 0.10 O It was noticed t h a t the conversion i i o 0.55 values obtained with both models .................. i o are similar, showing an average , vu 1oo6 0.50 0.0 0.2 0.4 0.6 0.8 1.0 difference of 5 % and a maximum z/Lc difference of 10 % at simulated Figure 3. Axial concentration profiles. conditions. However, it was (T=450~ I:L=I7 min) observed that the consideration of the radial dispersion phenomenon in the model is important to give a realistic physical behaviour for the concentration profiles in the reactor. Otherwise, the point where VL = 0 (r=r*) represents a physical barrier that prevents the interchange of mass between the u p w a r d and downward flow in the fully developed flow region. Consequently, there is a discontinuity in r* for the concentration profiles in the model without dispersion and the whole pattern of the simulated distribution of compounds in the reactor is affected. A parametric sensitivity analysis was performed for the determination of operating condition effects on the thermal cracking reactions. The variables with greater effects are the temperature and liquid residence time, respectively in this order. Residue conversion increases with the increase of both variables, as shown in Fig 4. The distribution of lumps at the reactor outlet stream is shown in Fig 5. .
i
50
[
=
30
. . . . . . . . . .
o
20
.................. . ~ . e . . O10/"
10
.
Gasoil
....= - . : . : i ~ m ~ - = 5
1'0
1'5
E
T = 4 2 5 ~ ..... ~ . - ....................T = 4 1 ~ i "
.........................................................
2'0 1;L (min)
2'5
3'0
35
Figure 4. Residue conversion as function of t e m p e r a t u r e and liquid residence time.
(Cs4)
-l D i s t i l l a t e ( C s 5 )
Naphta (Cs6) Gases (Cs7)
4=
E 4O
X
.
: ........
6O
o~ {D
.
Ih._l, 410 ~
In
i
i
i
=, L
425 ~ 440 ~ Temperature
455~
Figure 5. Distribution of products at the reactor outlet stream as function of temperature. (XL= 17 min)
5. C O N C L U S I O N A two-fluid model, based on a fluid dynamics approach is presented for a hydroconversion reactor operating in the slurry bubble column mode. This model accounts for the thermal cracking conditions in the reactor considering both the convection and the radial dispersion phenomena.
296
This model represents an additional improvement in obtaining an insight on how the fluid dynamics of the bed influences the reactions in a hydroconversion reactor. However, more testing and comparison with pilot plant and industrial data must be performed before it can be used for reactor design and scale up. 6. NOTATION Ci conc. of lump i, mass fraction Coi conc. of lump i at reactor entrance, mass fraction Csi conc. of lump i at the reactor outlet stream, mass fraction g gravitational acceleration, m/s 2 k rate constant, 1/s Lc reactor height, m P pressure, Pa r radial position, m r* radial position where VL=0, m Rc radius of the reactor, m T temperature, K U superficial velocity, m/s V axial velocity, m/s X conversion of residue, mass % z axial position, m
Greek letters {~i constant of the kinetic rate constant ~m,1 constant of the kinetic rate constant holdup, dimensionless ~ij,k eigenvalues for lump i ~eff effective viscosity, Pa.s p density, kg/m 3 residence time, s ~i,j,k eigenvectors of ~ij,k Subscripts g gas phase i lump i j indicates radial position k indicates eigenvalue L liquid phase sl slurry phase
7. R E F E R E N C E S
Abreu, J.C N.; Guirardello, R.; Medeiros, J. C. American Chemical Society, 213 th National Meeting, San Francisco, USA (1997) 364. Carbonell, M. M.; Guirardello, R. Chem. Engng Sci., 52 (1997) 4179. Carbonell, M.M.; Guirardello, R. Chem. Engng Sci., 53 (1998) 2479. Chen, Z.; Zheng, C.; Feng, Y.; Hofmann, H., Chem. Engng Sci., 50 (1995) 231. Delnoij, E.; Kuipers, J.A.M.; van Swaaij, W.P.M. Chem. Engng Sci., 52 (1997) 3623. Eccles, R. M. Fuel Processing Technology, 35 (1993) 21. Fan, L. S. Gas- Liquid- Solid Fluidization Engineering, Butterworths, 1989. Kam, E. K.T.; A1-Mashan, M.M.; Dashti, H. In: Catalysis in petroleum Refining and petrochemical Industries, eds. M. Absi-Halabi et al., Elsevier (1996) 283. Luo, X.; Zhang, J.; Tsuchiya, K.; Fan, L.S. Chem. Engng Sci., 52 (1997) 3693. Mosby, J.F.; Buttke,R.D.; Cox, J.A.; Nikolaides, C. Chem. Engng Sci., 41 (1986) 989. Sokolichin, A. Eigenberger, G. Chem. Engng Sci., 49 (1994) 5735. Tarmy, B.L.; Chang, M.; Coulaloglou, C.A.; Ponzi, P . R . I . Chem. E. Symposium Series No 87, Edinburg, Scotland (1984) 303. Tarmy, B.L.; Coulaloglou, C.A. Chem. Engng Sci., 47 (1992) 3231. Trambouze P. In: Chemical Reactor Technology for Environmentally Safe Reactors and Products, eds. H.I Lasa, G. Dogu, A. Ravella, Kluwer (1991) 409. ACKNOWLEDGEMENTS The authors acknowledge the financial support of C N P q - Conselho Nacional de Desenvolvimento Cientifico e Tecnol6gico (Brazil) - and PETROBRAS.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
297
AN INTEGRATED APPROACH FOR HYDROCRACKER MODELING
C S. L NARASIMHAN, M. SA U & R.P. VERMA
INDIAN OIL CORPORATION LIMITED RESEARCH AND DEVELOPMENT CENTER~ SECTOR-13, FARIDABAD, INDIA-121007
ABSTRACT" An integrated approach based on continuum theory of lumping is developed for modeling hydrocraddng kinetics & heat effects which has capabilities to predict _products yield and quality along with hydrogen oonsurnption & bed temperatureprofile. The hydrodynamics of trickle bed reactor has been modelled based on fundamental force balance approach which determines flow regimes, pressure drop and overall effectiveness factor. The kinetics along with heat and hydrodynamic effects is predicted for d_ifferent operating conditions of hydrocracker and the model predictions are compared with reported experimental data. The monitoring of industrial hydrocracker using the above approach is discussed in the paper. INTRODUCTION:
As the world refining scenario depicts stringent regulations and increasing demand for high quality and environmentally friendly petroleum products, hydrocracking, with its wide flexibility and capability to produce products to meet the present & future era demands, has emerged as the major secondary petroleum process in the refining industry. Further, hydrocracking helps the refiners to face the compounding challenge to produce desired quality products from feedstocks of increasingly deteriorating quality. The chemistry of the process is quite complex involving large number of series and parallel reactions occurring simultaneously. The reactions mainly consist of hydrogenation and cracking. A good kinetic model is essential for determination of products yield and quality at different severities. There are very few approaches on hydroeracker modeling in the open literature. They can be categorized under discrete lumping approack model compounds/structure oriented lumping approach, and continuum theory of lumping approach. Discrete lumping approach [Stangeland, 1974] [ 1] is limited with requirement of large number of model parameters and experimental data to obtain a reasonably acceptable model which would need frequent updations based on feed mix changes etc. On the other hand, the model compounds/structure oriented lumping approach [Quann & Jaffe, 1992] [2] would require large experimental data bank and computational power to build a reasonable model. An elegant approach using continuum theory of lumping has been recently attempted for
298
hydrocracking of n-paraffins [Browarzik & Kehlen, 1994] [3] & vacuum gas oils [Laxminarasimhan et al., 1996] [4]. It is always imperative to have an optimum analytical capability built in the modal so that the model is able to reasonably predict the products quality with variations in the feed quality. In the present paper, an elegant approach is presented for modelling hydrocracking kinetics with capabilities for prediction of paraffins(P), naphthenes(N) & aromatics(A) distribution in the products. This would be quite helpful as most of the products qualities are linked to the PNA distribution. The model formulation includes skewed-gaussian type primary yield distribution function chosen to describe the yield & selectivity of the hydrocracking components. The form of the primary yield distribution function has been evolved based on experimental data on yield patterns of hydrocracking of various model compounds (paraffinic, naphthenic & aromatic) reported in literature [5, 6, 7, 8, 9, 10, 11]. The approach follows the process chemistry to a reasonable extent. The present paper also describes detailed integration of kinetics, heat effects and hydrodynamics for development of hydrocracker simulator. The following aspects have been considered for integrated hydrocracker modelling:
(a) (b) (c) (d) (e) (0
Kinetics of hydrocracking of paraffins, naphthenes and aromatics based on continuum theory of lumping. Simultaneous hydrodesulfurisation kinetics with H2S inhibition effects. Hydrodenitrogenation kinetics and temporary catalyst deactivation effects due to presence of basic nitrogenous components in feed. Estimation of overall chemical hydrogen consumption across the bed by novel dynamic C/H ratio method. Estimation of heat effects by integrated approach. Hydrodynamic effects and determination of wetting efficiency, overall effectiveness factor and pressure drop.
The paper presents all the above aspects and case studies which have been conducted to benefit the operating refinery. HYDROCRACKING KINETICS: Hydrocracking kinetics comprises of reactions involving simultaneous hydrogenation and cracking. The major reactions are cracking, saturation and ring opening. The kinetics of hydrodenitrogenation (HDN) is quite important as it directly interferes with catalyst activity. Therefore, the kinetic model considers HDN kinetics which determines the extent of temporary deactivation of the catalyst. The model also incorporates novel hydrodesulphurisation (HI~S) kinetics in order to predict products sulphur.
MODEL FORMULATION The model formulation is based on continuum theory of lumping approach. [12, 4]. The complex reaction mixture is classified into continuous boiling mixtures of paraffinic, naphthenic and aromatic components. In other words, the feed characterization involves
299
TBP analysis and analysis of PNA for each narrow TBP fraction to generate separate TBP curves for Paraffins, Naphthen~and Aromatics. Based. on. the. above feed characterization and reaction trends of model compounds reported in literature, following reaction scheme has beenassumed .for Hydrocracking .of Aromatic, Naphthenic & P~raffinic components.
Hydrocracking of Aromatics A i ...........> ( A i-1 + A i-2 + . . . . . . .
+
A1) + ( Ni+
Ni-1 + . . . . . . . + N 1 ) + ( P i + Pi-1
+ ....... +
PI)
Hydrocracking of Naphtheues N i .......... >( Ni-1 + Ni-2 + ..... + N1) + ( Pi + Pi-i + .......... + P1) Hydrocracking of Paraffins Pi ........... > ( Pi-1 + P i-2 + . . . . . . . .
+P1)
where Ai, Ni, & Pi represent the yield distribution of Aromatic, Naphthenic and Paraffinic components respectively. The individual component material .balance equations for continuous mixture should satisfy the following material bataneeeonstraintskA(1)
k~,(1)
kp(1)
I P(kA,KA)29(KA).dK~+ Ip(k,,KA).D(KD.dKN § K.4
k
[. P(kN,x,,).D(IG).alG + [.m~;zc~,).D(x~).ax~ K ~r
kpO)
IP(k~,,K~,).D(Kp).dKp = 1
(t) k2
=~
(2)
k3
(3)
kp
where,
kl = fs [ fN"1 (ks) ] k2 = fN [ fp-1 (kp) ] k3 -= fA [ fl~"1(kl~) ] kA, ks, kp are monotonic function of normalised TBP, '0' defined as follows"
kA = fA ( 0 ,CZA). ricE. to. t~; k~= fs ( 0 ,czs ) . riCE .tr tn; ke - fp ( 0 ,etp ) . riCE. tr tn where, riCE is contacting efficiency and normalised TBP (0) is defined as: ( T B P - TBP(I)) / (TBP(h)- TBP(1)), kA (1), ks(l) & kp(1) denotes the lowest reactivity of Aromatic, Naphthenic and Paraffinic components respectively. CZA,CLS& CZeare model parameters to_bet determined based on experimental dat a. -rice-is-the.contacting effectiveness factordetermined from hydrodynamics, tc is the Arrhenius temperature-effect factor and t~ is nitrogen deactivation effect factor. The D(K) functions [D(KA),-D(KN) & D(Ke)] are-11aespecies t y ~ distribution fu~tions (Chou &Ho, 1987) [ 11 ] defined as follows :
D ( K ) - di . dO dO dk
(4)
The P(k,K) functions described above, determine the yields of species with reactivity k from hydrocracking of components with reactivity K, and are known as yield distribution functions. Based on the cracking pattern of various model-compounds-reported in literature [6, 7]. are shown in figures 1 and 2, it can be seen that yield distributions due to primary
300
cracking tend to follow a skewed gaussian type distribution function. Thus, the P(k,K) function is assumed to have a skewed gaussian distribution function and can be represented in the following form P(k, k) function assumes the following form
f
~k,K)= S.--.-.~ ~ exr,
~
(5)
+A+
where S, a0 ,al and 8 are model parameters.
HDN Kinetics : Nitrogenous components have to be converted dining hydrocracking reactions for better products quality. Also basic nitrogen components deactivate the catalyst and therefore they have to be removed before entering into valnexableeatalyst beds. The.HDN kine,tics is also based on continuum theory of lumping. The nitrogenous compound distribution is generated with respect to boiling point distributiata. The HDN kinetic estimations are as foltowsc dC , _ ~" ( k,,.rlc, .'~ ). p ~r' .C dt o 1 + k,p
.D ( k , )a~ ,
(6)
where D(kn) is species type distribution for nitrogen components in t-he same units as defined earlier for cracking, C, is. concentration of nitrogen in the reaction mixture and 1~ & ks are kinetic parameters. Deactivation due to presence of.nitrogen-compound-is represented by the expression : t~= f (Cn)
HDS kinetics In order to estimate the products sulphur levels,-HDS kinetics is integrated with the Hydrocracker model and is based on continuum theory of lumping approach using sulfur compound class distribution e.g., mereaptans, thiophenes,-benzothiophenes, dibenzothiophenes (DBTs) etc., in different boiling ranges. The kinetics also include H2S inhibition effects as the amount of H2S generated would inhibit both hydrotreating and hydrocracking reactions. A detailed description of the above kinetic model has been presented in one of our earlier papers [Sau et al., 1997][13].
Estimation of Hydroeen consumption t
Hydrogen consumption estimation -is based on hydrogen balance equations using novel d.y~mic C/H ratio approach for hydroeracking. Dynamic C/H ratio concept is based on determination of C/H ratio function (with respect to-TBP and PNA-distribution in the reaction mixture) along with bed of the reactor. It is observed that C/H ratio for any given petroleum mixture is a monotonic-function of boiling point- [Nelson 1988] [14]. Thus for the petroleum mixture which is considered as mixture of Paraffinic, Napthinic & Aromatic components distributed across '0', the-C/H ratio with respect, to each of them can be represented as follows: C (ka)=12 + fla(ka},C (ku)= 6+,O~(ks},C (k~)- 3+ fl~_(kl,)
301
where 13A,13N& 13Pdepend on maximum C/H ratio components present in the mixture. It may be noted that [~A, [~N, .[~tp ~ dynamically updated with hydrogenation. _Hydrogen consumption estimation is based on the following equations.
!
!E~kj~:A
]
'1
k,
]~= ~h(k)~k~k 0
F
.~
.
1
1
(7)
_
where h is average hydrogen consumption. As hydroeraeking progresses, -the parameters 13A, 13N& 13V are updated for every reactor bed dement. Estimation of Temperature Effects 9
Hydrocracking is an exothermic reaction. The heat effects vary for different components. For example, the _aromatic components generate more heat per mole of hydrogen consumption then naphthenic and paraffinic components. This effect is incorporated in the development of novel heat-effect model based on continuum theory of lumping to obtain the reactor temperatar~e profile_ The temperature dependency of each reactive species is detexmined by following Arrhenius type of expression.
k =k,.e
Rkr r,j
where E is activation energy, r is the refererw,e -valueand R is universal gas const~_ut. The reactor heat balance is presented as follows ~.o_. d-_T=v. "dz ~o dz k
~
0
o
-
-0
It
dz
0
Hydrodynamic effects:
Hydrodynamic effects are quite impoa.~t v~hileconsidering the overall reactor-model for -hydraeraeker.which is.a high pressure tfic&le bed reactor__ There. are limited number, of empirical hydrodynamic models in open literature to predict Ihe hydrodynamic effects at high pressure of hydrocrackiag.. The available empirieat equations of Larachi_etaL [151 etc., are based on moderate pressure and-have very limited range of~plicability. In view of the above, a novel hydrodynamic_ model based_on fundamental force balance approach has been developed and found to have wide range of predictability.raaging from medium pressure experimental data torelati~ely high pressure_[!61. Based on the above model, the reactor AP & liquid hold up is obtained along with flow re82me. Considering negligible, axial dispersion effects in_the eommexeial_ reactor tmits~ the overall effectiveness factor_ is .close to the. wetting efficiency, which is given by _the correlations of Al-dahhan et al. [17]. The overall effectiveness factor thus determined is
302
incorporated in the kinetic model to obtain the actual reaction rate in the commercial reactor. SOLUTION METHODOLOGY The above system of coupled integro-differential equations are solved numerically to determine the temperature profile along with the concentration profile across the reactor bed. The model equations are solved using numerical techniques to obtain the distribution of paraffins, naphthenes and aromatics in the reaction mixture as hydrocracking progresses (i.e. at different severities). Based on PNA distribution of the product mixture the yield & quality of desired product slate can be determined. Model configuration and validation The model formulations have been tested with pilot plant data from reported literature [Bennette &Boume] [ 18] and it has been found that predictions of the model match well with the experimental results (Figures 3 & 4). The integrated model has been configured to a two stage hydrocracker with recycle. The first step for configuration of the model is estimation of model parameters for both the stages individually. Test run has b e e n _ c o n d ~ to estimate the model parameters. Once the model parameters are estimated, the simulator has been used to predict historical data set of the above plant and it can be seen from figures 5, 6, 7 & 8 that the model predictions for PNA distribution matches reasonably well with the plant data. The model was used to monitor the plant for significant time continuously and figures 9, 10 & 11 show that the model predicts the overall yields, Hydrogen consumption of each stage and temperature profile (depicted through Catalyst Average Temperature of the beds (CAT)) quite consistently for both the stages. The model also predicts the sulfur and nitrogen distribution well as can be seen in Table-1. SUMMARY A novel integrated hydrocracker model has been developed based on continuum theory of lumping approach. The unique features of the modeling approach e. g. PNA distribution etc. enables the model to depict the chemistry of hydrocracking process closely and allow it to provide good process insight-s for hydrocracker operation & monitoring. The model has been configured to monitor and optimize a 2 stage commercial hydrocracker with recycle. The model predictions for yield, hydrogen consumption and ~talyst average temperature are excellent for commercial plant. This model is of immense help to the refineries for plant monitoring and prediction of yield potential for feed changeover. The model can be used for process design on the basis of kinetic data generated from pilot units. Applications of the model can be extended to catalysts evaluation and feed mix optimization. The model can also be configured to work with plant advanced control systems For overall optimization.
303
REFERENCES:
1. B.E. Stangeland, "Kinetic model prediction of hydrocracker yields", Ind. Eng. Chem. Proc. Des. Dev., 13 (1), 72 (1974). 2. R.J. Quann, & S. B. Jaffe, "Structure oriented Lumping: describing the chemistry of complex hydrocarbon mixture," Ind. Eng. Chem. Res., 31, 2483 (1992). 3. D. Browarzik, & H. Kehlen, "I-Iydrocracking process of n-alkanes by continuous kinetics", Chem. Eng. Sci., 49(6), 923 (1994). 4. C. S. Laxminarasimhan, R. P. Verma & P. A. Ramachandran, "Continuous Lumping model for simulation of hydrocracking", AIChE Journal, vo142, No. 9, 2645 (1996). 5. G.F. Froment, Mobil Workshop on chemical reactions in complex mixtures (proceedings) (1990) 6. R.F. Sullivan, C. J. Egan, & G. E. Langalois, J. of catalysis 3 185-195 (1964). 7. C.J. Egan, G. E. Langlois & R. J. White, "Selective Hydrocraeking of C9 - to C~2Alkylcyclohexanes on Acidic Catalyst. Evidence for the Paring Reaction", J. A. C. S., 84, 1204 (1962). 8. G.E. Langlois, & R. F. Sullivan, Preprints, Div of Petrol. Chem., ACS, 14 (4), D 18 - D 39 (1969). 9. H. L. Coonradt, & W. E. Garwood, Ind. & Eng. Chem. Proc. Des. Dec. , 3 (1), (1964) 10. D. K. Liguras, & D.T. Allen, Ind. & Eng. Chem Res., 28, 665-673 &674-683 (1989). 11. D. Browarizik, & H. Kehien, Chem. Eng. Sci., 49(6), 923-926 (1994) 12. M. Y. Chou, & T. C. Ho, AIChE J., 34, 1519 (1988) 13. M. Sau, C. S. L. Narasimhan, R. P. Verma, "A Kinetic Model for hydrodesulphurisation", Proceedings of First. Intl.-Symposium on Hydrotreating .& Hydrocracking of oil fraction, February 1997, Belgium. 14. W. L Nelson, 'Petroleum Refinery Engineering', 4th ed., Me. Graw- Hill Book Co. Inc., New York, 169 (1958)15. F. A. Larachi, A. Laurent, N. Midoux & G. Wild 'Experimental study of Trickle bed reactor operating at high pressure: Two phase laressure drop and liquid saturation', Chem. Eng. Sci., vo146, 5/6, pp 1233-1246, 1991. 16. M. Sau, C. S. Laxminarasimhan, R. P. Verma & P. A. Ramachandran, "Modelling of Trickel-bed and Packed Bubble Column Reactors - Hydrodynamic Aspects", presented at Chem. Engg. Congress (CHEMCON), Bombay, 1993. 17. M. H. Al-dahhan & M. P. Dudukovic, 'Catalyst wetting efficiency in trickle bed reactors at high pressure', Chem. Eng. Sci., 50, 15, pp 2377-2389, 1995. 18. R. N. Bennett, & K.H. Bourne, ACS Symposium on Advances in Distillate & Residual Oil Technology, G45-G62 (1972)
304
HYDROCRACKING OF PHENANTHRENE
HYDROCRACKING OF HEXAMETHYL CYCIZ)HEXANE 28 40
FUSEDB ALKANE
I
C
Y
C
~
CYCLOHEXANE 1
~5
MONOCY~ARBONS ,.,,.
9
~:~ ~o .i
2 S C.AimON~
4
S 9 7 OlrlqLODl~=rm ~
8
9
9
2
l~pJre-2
lqlpn~l
FEED DATA FOR PARAFFINS, NAPHTHENES & AROMATICS ( BENNETT & BOURNE, 1972)
9 9 It lO 12 CARBONNUM]~.ROIFPlK)DUCTMOLECULE
14
MODEL PREDICTIONS AT SEVERITY 2.117 (1/hr) COMPARISON WITH EXPERIMENTAL DATA Weight fraction ( B E N N E T T & B O U R N E , 1972) 1
-.-
0.8
I-
-
i
0.8
--k-Total
0.6
--~- Aromatics I Naphthems [ -*- Parafnm I i
0.4 0.2
....
0 It -200-100
r m p e r m r e O ~ ~-3
~ s
0
100
200
300
400
TBP & PNA DISTRIBUTION DATA OF FIRST STAGE PRODUCT w Total-mad
1
Armma~ks-m~l
tk8
0-6
--*-- T o t a l ---- A r o m m U c s Napi~t~mmm
0.4
---
~ L6 ~
)P'anglEm~mmll 9 Tmlal-lP,h ~
0.4
Parafllm~s
0.2
0
Q
-
0
0.2
0.4
,
0.6
0.S
Normalised TBP ~
12)
T 600
M O D E L P R E D I C T I O N S VS. P L A N T D A T A
wdot rrms~ 0.8
500
Temperarare (Deg 9
FIBre - 4
F E E D D A T A O F R E A C T O R -1 F O R PARAFFINS, N ~ & AROMATICS (COMMERCIAL PLANT DATA)
]
9
1
L2
~4
~6
HORMAI.11S~D~I~P
l~iure- 6
0-8
1
Arematic~ Plant NalPil/lltemm~ lPlam 9 ~Plaalt
306
ATTACHMENT Species Type Distribution Function D(k) The above function is Jacobian of species-to-reactivity co-ordinate transformation and accounts for the number of species having same reactivity 'k' in a given reaction mixture. In other words D(k).dk denotes the number of species with reactivity between k and k + dk. The major advantage of this approach is that there is no concentration distribution function involved while developing the kinetic expressions, as c(k,t) is just the concentration of the component with reactivity k at any given time 't'. It may be noted here that the distributive nature is shifted to the D(k) function, which when multiplied with concentration function c(k,t) gives the concentration of all the species with reactivity between k and k + dk. Mathematically the species type distribution function D(k) is represented as D (k ) = di__= di dO. dk dO dk
where 'i' is species index, '0' is normalised TBP and 'k' is reactivity. If 'i' is equally spaced then di ~-~N dO Therefore, dO D(k) = N . ~ dk
P(k,K) : Yield Distribution Function P(k,K) function is a kind of yield distribution function describing the formation of components of reactivity k(k
-7
I
FUNDAMENTALS AND REACTION MECHANISMS
I !
I
I
1
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
309
Ab-Initio E n e r g y Profiles for T h i o p h e n e H D S on the MoS2 (1010) E d g e - s u r f a c e P. Raybaud a, J. Hafner ~, G. Kresse ~ and H. Toulhoat a ~' Institut Frangais du P~trole, Division Informatique Scientifique et Math~matiques Appliqu~es, 1-4 av. de Bois-Preau, 92852 Rueil-Malmaison, France. b Universit~it Wien, Sensengasse 8, A-1090 Wien, Austria.
Abstract We establish the energy profiles for thiophene hydrodesulfurization on the edge(1010) surface for relevant model mechanisms. Two kinds of adsorption configurations (.ql and 'r]~) for the thiophene and its hydrogenated derivatives are taken into account. We have considered a supply of hydrogen either from Mo-SH surface groups or from Mo-H groups. The q5 configuration, which is likely to occur at low surface coverages, is clearly the most favorable one for the hydrogenation steps. In the 711 configuration, the role of co-adsorbed-SH groups becomes crucial for the hydrogen transfer during the hydrogenation steps. At this stage of the study, the sulfur removal step appears as the most unfavorable step for a non-promoted catalyst.
1
Introduction
A number of model mechanisms for the thiophene hydrodesulfurization (HDS) reaction have been proposed over the past [1-7]. Theoretical approaches within the Density Functional Theory (DFT) offer a new way to investigate the various mechanisms and to derive new insights on the rate determining step(s) of the HDS reaction [8]. In previous studies [8-10], we have shown that DFT within the Generalized Gradient Approximation (GGA) describes accurately the crystal and electronic structures as well as the energetics of about thirty transition metal sulfides (TMS). The structural and electronic properties of the MoS2 (1010) and (1010) edge-surfaces with an adsorbed reactant like thiophene are also well described by this technique [12, 13]. This gives us first insights for the initial step of the mechanism of the thiophene HDS reaction. The goal of the present study is to investigate the thiophene HDS mechanism for realistic MoS2 surfaces using DFT calculations. On the one hand, we will analyze the influence of the adsorption mode of thiophene and its derivatives on the active site along the subsequent steps of the catalytic cycle. Indeed this mode depends on the various experimental conditions (T, PH2S, PH~ o r Preactant, ...). On the other hand, it is of interest to take into account the role of the Mo-H or Mo-SH surface groups which are likely to be present close to the adsorbed reactant.
310
2 Methodology We perform the energy profiles of the thiophene HDS on the (1010) MoS2 surface. In the same spirit as our previous theoretical studies [12, 13], the MoS2 slab contains 72 atoms per unit cell and initially three "Coordinatively U n s a t u r a t e d Sites (CUS)". Along the reaction pathway, the co-adsorbed species on the surface sites are sulfur-bearing molecules (thiophene or derivatives), Mo-H and Mo-SH groups. A schematic model of the thiophene HDS mechanism is given in Figure 1. The hydrogenation steps are investigated in the form of successive atomic H additions (as proposed in [7]), also called monohydrogenations. The H atoms (written a s :~ t.H2 in Figure 1) are transfered either from surface Mo-H groups or from MoSH groups to the adsorbed sulfur-bearing molecules. The influence of the two hydrogen sources on the energy profile is analyzed. The ql and ~1~ adsorption modes are considered for the thiophene and its hydrogenated derivatives. The sulfur-carbon bond scission steps are investigated through a two step model. Each sulfur-carbon bond scission occurs through hydrogen addition on the C~, atom, assuming the formation of a thiolate intermediate.
Figure 1" Schematic pathways investigatigated for the thiophene HDS reaction over the MoS2 (1010) surface. The total energy calculations are performing using the Vienna Ab-initio Simulation Package (VASP) [14, 15] which determines an iterative solution of the generalized Kohn-Sham equations of the Local Density Functional Theory. The Generalized Gradient Correction are included as proposed by Perdew and Wang. The electron-ion interactions are described by Vanderbilt's ultrasoft pseudopotentials. The geometry optimization (ionic relaxation of the first upper MoS2 layer and adsorbed species) for each step of the mechanism is performed using the exact Hellmann-Feynman forces. The optimized structures along the reaction pathway are given in Figure 2.
+ Butane
Figure 2 : Optimized structures for each step of the catalytic HDS reaction path with the initially 7 5 adsorbed thiophene. The first and third monohydrogenations are turned out through H transfered from Mo-SH groups. The italic numbering corresponds to Figures 3 and 4.
+ %s
312
3 Hydrogenation steps We calculate the energy profiles of the successive monohydrogenation steps for the thiophene molecule via the hypothetical following intermediates: 2-monohydrothiophene, 2,5-dihydrothiophene, 2,3,4-trihydrothiophene or 2,3,5-trihydrothiophene and finally tetrahydrothiophene. The energy profiles in a pure gas phase and on the surface are represented on Figure 3. 3.1 In t h e gas p h a s e We establish first a reference profile for the hydrogenation path in a pure gas phase (see full line on Figure 3) for all the species involved. We take into consideration the spin corrections for the the gas monohydrothiophene and trihydrothiophene. The energy costs are +0.64 eV and +0.36 eV for the first and third monohydrogenation respectively. Although we have not determined true transition states, in a pure gas phase, these steps are likely to be rate determining for the hydrogenation process. This is in agreement with the experimental knowledge that a free thiophene molecule in a gas phase cannot be attacked by hydrides species. 3.2 On t h e s u r f a c e The effects of the sulfur coverage and of the reductive environment on the energy profile are investigated by assuming the co-adsorption of Mo-SH and Mo-H with the reflctant molecule. The presence of Mo-SH groups seems to be clearly admitted, since the early work of Maternova [16]. The formation of the Mo-SH groups could be explained by several model reactions involving H2 or H2S dissociation [17], as well as the HDS reaction of sulfur compounds (such as R-SH). It is not within the scope of this work to identify the way(s) these groups are formed. At the usual high H2 partial pressures prevalent in HDS industrial operating conditions, one expects to have all terminal S at the surface, s a t u r a t e d by H. The existence of Mo-H species is not yet experimentally established, but we assume they are present on the surface as it may be supported by our investigations. Table 1" Calculated adsorption energies (in eV) of thiophene and of its hydrogenated derivatives for the 7]1 and r]~ configurations on the bare MoS2 (1010) surface and in the presence of one Mo-H group and one Mo-SH group. Compounds thiophene 2-monohydrothiophene 2,5-dihydrothiophene 2,3,4-trihydrothiophene 2,3,5-trihydrothiophene tetrahydrothiophene
bare surface 711 qs 1.27 2.00 1.74 3.62 2.06 2.80 2.04 3.12 1.60 3.62 2.09 1.94
Mo-H, Mo-SH ql q5 0.54 0.94 n.c. n.c. 1.31 1.87 n.c. n.c. n.c. n.c. 1.38 n.c.
313
The presence of Mo-H and Mo-SH groups strongly lowers the adsorption energies of thiophene and its hydrogenated derivatives in comparison with the bare surface. Table 1 shows that, in the presence of one Mo-H and one Mo-SH group, the 'q~ mode is still more favorable. Nevertheless the corresponding adsorption energies are about I eV lower than on the bare surface. The adsorption energies for the ~]1 adsorbed compounds are about 0.7-0.8 eV lower. The hydrogen transfer from the surface to the different adsorbates is analyzed following two possible pathways. Along pathway (a) (see Figure 3), the first and third monohydrogenations (steps 3 and 5) occur with the H atom transfered from the surface Mo-H groups, whereas the second/fourth monohydrogenations through the Mo-SH groups. Along pathway (b), the first and third monohydrogenations are achieved with the H atom transfered from the Mo-SH groups. 1.0
2-MHT +0.64/--'-'~
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~
,
_
~
,
,
/
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.
.
.
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.
.
.
.
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'
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.=____._._._.Z;
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....
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.
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-,,
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-0.46-
4
.......... 115(a) .......... 115 (b)
~0~13
5
2,3,5-tilT I +H 6
7,8[.---..
Reaction path
Figure 3: Energy profiles for the four monohydrogenation steps of thiophene with preliminary H2 dissociative adsorption over the MoS2 (1010) surface giving surface Mo-H and Mo-SH groups. First and third monohydrogenations take place with H transfered (a) from the Mo-H groups, (b) from the Mo-SH groups. The italic numbering corresponds to Figure 2. (*) Energy reference: MoS2 surface with one preadsorbed sulfur + free thiophene + 8.H2. 3 . 2 . 1 7]5 c o n f i g u r a t i o n
The q~ adsorption mode appears as the more favorable for the hydrogenation process. The energy barriers observed for the gas phase totally vanish whatever the pathway (a) or (b) may be. This result points out the role of the catalytic surface
314
which favors the thiophene hydrogenation reaction by lowering the energies of the first and third monohydrogenation steps. As it is explained in [12], the high stability of the free thiophene is strongly perturbed by the donation and backdonation effects between the carbon or sulfur ~-orbitals and the molybdenum d-orbitals. Simultaneously the interactions between the monohydrothiophene, dihydrothiophene and trihydrothiophene are even stronger (see Table 1) so that the first and third monohydrogenation steps are enhanced. The energy gain for the resulting adsorbed monohydrothiophene or trihydrothiophene balances the energy cost for removing an H atom from the surface Mo-H or Mo-SH groups. The energy profile for this adsorption mode does not depend on the kind of the surface groups (Mo-H or Mo-SH) involved in the H transfer. The chemical reactivity is explained by the strong interaction between the adsorbed u n s a t u r a t e d sulfur bearing organic molecules and the CUS site.
3.2.2 J11 configuration In this case, the energy profile depends clearly on the pathway (a) or (b). The profile obtained along pathway (a) where Mo-H groups are the source of hydrogen for the first and third monohydrogenations, is quite similar to the one established for the gas phase : the energy costs for the first and third monohydrogenations are rather high: +0.48 eV and +0.62 eV respectively. They remain the unfavorable steps as it is in the gas phase. The ql adsorption configuration involves mainly the sulfur "lone pair" electrons. As a result, the stability of the aromatic ring is far less perturbed. This mode can be regarded as a medium mode between the pure gas phase and the 71~ one. The adsorption energies are smaller for all the derivatives except for the tetrahydrothiophene where the .//1 configuration becomes more favorable (see Table 1). For the pathway (b) where the first and third monohydrogenations take place through H transfer from Mo-SH groups, these steps are now slightly exothermic and the energy profile of the reaction is more favorable. The energy barriers vanish, which implies that the hydrogenation process in the .q~ mode is favored by Mo-SH groups. The explanation has to be found in the higher instability of the hydrogen atom in Mo-SH groups than the one in Mo-H groups or in the gas phase. The dissociative adsorption of the H2 molecule over a surface with one preadsorbed sulfur atom producing Mo-H and Mo-SH groups is an exothermic process of -0.38 eV (see step i in Figure 3). However for the sulfur coverage we consider here, the binding energy of one H atom in a Mo-H group is larger than the binding energy of an H atom in a Mo-SH group. As a consequence, this later hydrogen atom is more unstable and reacts preferentially with the 71~ adsorbed thiophene. For the 71~ adsorption mode, the chemical reactivity is explained by the presence of the surface neighboring Mo-SH groups. For low surface coverages, the 7/~ adsorption mode of the thiophene and its hydrogenated derivatives is favored and the hydrogenation process is enhanced by the strong interaction between the reactant and the CUS. For high surface coverages, the 'l~ 1 mode is sterically enhanced~ The hydrogenation process is still possible thanks to the high reactivity of the neighboring Mo-SH groups.
315
4
Sulfur-carbon scission step
The sulfur-carbon scission step is investigated by a s s u m i n g two successive sulfurcarbon bond scissions of the q5 adsorbed tetrahydrothiophene with H transfer and formation of a thiolate intermediate. The final product is butane. The H transfer is realized through a similar process to the hydrogenation steps considering two kinds of hydrogen sources: surface Mo-H and Mo-SH groups. -2.0
-3.0
H2S
~
THT THT+2H* -4.0
'/
+H*
10
12
11
]
0.92
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Butane
13 -6.0
+1.59[
9 _0.661+0.2 Thiol .+2H* Thiolate
-5.0
+oo4/_ Thiol
14
Reaction path
Figure 4: Energy profiles of the desulfurization step and sulfur removal step of thiophene. P r e l i m i n a r y H2 dissociative adsorption over the MoS2 surface with one adorbed sulfur atom giving one surface Mo-H groups and one Mo-SH group. The italic n u m b e r i n g corresponds to Figure 2. (*) The energy reference is the same as in Figure 3. The energy variation for the first S-C bond scission, induced by the H atom from the Mo-H group (step n u m b e r 10) is -0.66 eV (see Figure 4). The same value is obtained for a H transfer from the Mo-SH. We a s s u m e the second S-C scission does not take place immediately, but we transfer one H atom from the Mo-SH group to the thiolate i n t e r m e d i a t e to form an adsorbed thiol compound (step n u m b e r 11). The energy cost of this transfer is about +0.25 eV and is not physically impossiblen. After one more H2 dissociative adsorption (step n u m b e r 12), one more H is transfer to the thiol to produce b u t a n e and leave two Mo-SH groups (step n u m b e r 13). The sulfur-carbon bond scission step exhibits no large endothermic variation and is not energetically unfavorable. The influence of Mo-H groups or Mo-SH is negligible for this step. This is explained by the strong S-Mo bond already observed in MoS2 based catalyst [11].
316
5
Sulfur r e m o v a l step
This step (14 in Figures 2 and 4) enables the reactivation of the molybdenum site by removing the sulfur atom left by the organic molecule on the surface after the second S-C scission. One could imagine several model pathways [17] for this step depending on the kind and number of remaining surface sulfur species: Mo-S, Mo-SH or molecular adsorbed H2S. We propose that two neighboring MoSH groups are involved in the process of the sulfur removal step (see Figure 2). This minimizes the energy cost which is about +1.59 eV. This step remains energetically the most unfavorable step for the whole catalytic cycle.
6
Conclusions
Performing ab-initio calculation within Density Functional Theory and Generalized Gradient Correction, we have established the first energy profile for the reaction of thiophene HDS on realistic (1010) MoS2 edge-surfaces. The study has taken into account the effects of the co-adsorbed species and adsorption modes of the reactants. The hydrogenation steps investigated by successive monohydrogenations of the thiophene molecule, do not present any energy barriers over the catalytic surface. In opposition to the hydrogenation of the free thiophene, the adsorbed molecule is easily hydrogenated on the surface. At low sulfur coverage (or low chemical potential of sulfur, see also [18]), the thiophene is activated in the .q.5 configuration. For higher sulfur coverage (or high chemical potential of sulfur), where the q~ configuration takes place preferentially, the hydrogenation reactivity is enhanced by the neighboring MoSH groups. The sulfur-carbon bond scission can occur in a consecutive fashion without any significant energy barrier. The sulfur-removal step is the rate determining step. The theoretical study of Neurock and van Santen [8] on small Ni sulfide clusters has found that the S-C bond scission step and S-removal step can compete as rate-controlling steps. A direct comparison with our present work done on periodic systems is difficult. However we are currently pursuing our investigations with realistic model MoS2 surfaces promoted by 3d-TMS. The promoter addition modifies the sulfur-metal bond strength [19], which implies a change in the whole energy profile and a displacement of the unfavorable step. This is not contradictory to [8] and could be explained by the Sabatier principle, according to which a volcano curve appears with plotting the HDS activity against the relevant S-M bond strength across the whole TMS series [11].
Aknowledgements This work has been performed within the "GdR Dynamique Mol~culaire Quantique Appliqu~e ~ la Catalyse", a joint project of Centre National de la Recherche Scientifique (CNRS), Universitat Wien (UW), Institut Fran~ais du P~trole (IFP), TOTAL and Schuit Institute of Catalysis.
317
References [1] J.M.J.G. Lipsch and G.C.A. Schuit, J. Catal. 15, 179 (1969). [2] S. Kolboe, Canad. J. Chem. 47, 352 (1969). [3] J. Kraus and M. Zrazil, React. Kinet. Catal. Lett. 6, 475 (1977). [4] H. Kwart, G.C.A. Schuit and B.C. Gates, J. Catal. 61,128 (1980). [5] B. Delmon, J-L. Dallons, Bull. Soc. Chem. Belge. 97, 473 (1988). [6] B.C. Wiegand and C.M. Friend, Chem. Rev. 92, 491 (1992). [7] J.W. Benson, G.L. Schrader, R.J. Angelici, J. Mol. Catal. A: Chem. 96, 283 (1995). [8] M. Neurock and R.A. van Santen, J. Am. Chem. Soc. 116, 4427 (1994). [9] P. Raybaud, G. Kresse, J. Hafner and H. Toulhoat, J. Phys.: Condens. Matter 9, 11085 (1997). [10] P. Raybaud, G. Kresse, J. Hafner and H. Toulhoat, J. Phys.: Condens. Matter 9, 11107 (1997). [11] H. Toulhoat, P. Raybaud, S. Kasztelan, G. Kresse, J. Hafner and G.Kresse, Catal. Today 1625, 1 (1999). [12] P. Raybaud, J. Hafner, G. Kresse and H. Toulhoat, Surf. Sci. 407, 237 (1998). [13] P. Raybaud, J. Hafner, G. Kresse and H. Toulhoat Phys. Rev. Lett. 80, 1481 (1998). [14] G. Kresse and J. Hafner, Phys. Rev. B 47, 588 (1993); ibid. 49, 14251 (1994). [15] G. Kresse and J. Furthmfiller, Computat. Mat. Sci. 6, 15 (1996); Phys. Rev. B 54, 11961 (1996). [16] J. Maternov~, Appl. Catal. 3, 3 (1982). [17] S. Kasztelan, p29 in Hydrotreating technology for pollution control-
Catalysts, Catalysis, and Processes, edited by M.L. Occeli and R. Chianelli (1996). [18] P. Raybaud, J. Hafner, G. Kresse and S. Kasztelan, H. Toulhoat, submitted to J. Catal. [19] P. Raybaud, J. Hafner, G. Kresse and S. Kasztelan, H. Toulhoat, to be submitted.
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HydrotreatmentandHydrocrackingof Oil Fractions B. Delmon,G.F.FromentandP. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
Diffusion effects and direct C-N cleavage o-toluidine and methylcyclohexylamine NiMo/?-A12Oa and Mo(P)/?-AI203 catalysts
319
in the HDN of over sulphided
Fabio Rota and Roel Prins Laboratory for Technical Chemistry, Swiss Federal Institute of Technology (ETH), CH-8092 Ziirich
Abstract The mechanism of the C-N bond cleavage in hydrodenitrogenation over NiMo/~-A1203 and Mo(P)/~,-A1203 catalysts was studied. The major part of C-N cleavage of o-toluidine takes place via ring hydrogenation and ~-Hofmann elimination. Direct conversions from o-toluidine to toluene and from methylcyclohexylamine to methylcyclohexane were observed as well. Some of these direct conversions were, however, only apparent. They were caused by diffusion limitation, because the intermediates could not diffuse out of the pores fast enough to be detected. The influence of diffusion was greater for the more active NiMo/?-A1203 catalyst. Experiments with small catalyst particles at low conversions demonstrated that, even in the absence of diffusion limitation, direct C-N bond cleavage takes place. This is ascribed to C-N bond hydrogenolysis or to nucleophilic substitution of the NH2 group by a SH group, followed by C-S bond hydrogenolysis. 1. I N T R O D U C T I O N Hydrodenitrogenation (HDN) is an important step in hydrotreating. Several authors have studied the activity and structure of the catalysts [1] as well as the mechanisms and the kinetics of the HDN reaction [2,3]. The HDN of nitrogencontaining aromatic heterocycles goes via hydrogenation of the aromatic heterocycle, followed by ring opening and removal of ammonia by C-N bond breaking, and eventually by hydrogenation of intermediate olefins. The reaction of quinoline (Q) to intermediate propylcyclohexylamine (PCHA) and finally to propylcyclohexane (PCH) is a good example of such an HDN reaction (Figure 1) [4,5]. Previous work showed that most of the nitrogen atoms in decahydroquinoline (DHQ) and o-propylaniline (OPA) are removed via elimination, as in PCHA to propylcyclohexene (PCHE) (Figure 1) [4,6]. A significant fraction of the substrate, however, reacted by means of a different reaction. Thus, about 6% of OPA reacted to propylbenzene (PB), and 15% PCHA reacted directlv to PCH [6]. Both reactions are considered to take place as hydrogenolysis, that is reactions in which a bond is broken (a C-N bond in these cases) and the resulting radicals are s a t u r a t e d by hydrogenation. While the reaction of OPA to PB was observed directly [5], proof of the reaction of PCHA to PCH was obtained indirectly. At low conversion of DHQ, the selectivity for PCH did not go to zero, as should have happened in a s t a n d a r d PCHA to PCHE to PCH consecutive reaction. From a
320 m a t h e m a t i c a l point of view, it was, therefore, necessary to assume t h a t a direct PCHA to PCH reaction took place in addition to the simple consecutive reaction network. Such a reaction is possible at the catalyst surface. W h e n the catalyst can split the C-N bond of OPA [5], which is stronger t h a n an aliphatic single C-N bond, it should certainly be able to split the C-N bond of PCHA.
Q
THQ-1
iL
OPA PCHE
1l
THQ-5
PB
DHQ
PCHA
PCH
Figure 1. Network of the HDN of quinoline. As suggested previously [7,8], another explanation for the seemingly direct formation of PCH from PCHA may, however, be the influence of diffusion. If PCHA reacts to PCHE and PCHE reacts rapidly to PCH, then the i n t e r m e d i a t e P C H E t h a t is formed in the catalyst pores has hardly any time to diffuse out of the pores before it reacts further to PCH. As a consequence, the observed selectivity for the intermediate PCHE (as m e a s u r e d at the end of the reactor) is always underestimated, whereas t h a t for the end product P C H is always overestimated, as we observed [7-9]. If this explanation of diffusion limitation is correct, then hydrogenolysis of a single C-N bond would not have to be considered, and HDN over sulphidic catalysts would take place solely via hydrogenation and elimination reactions. In the other case, hydrogenolysis would play an i m p o r t a n t role, which, depending on the operating conditions, could be substantial. To determine the reason for the high PCH selectivity, we performed studies of the H D N of o-toluidine (TOL) and methylcyclohexylamine (MCHA), since MCHA is commercially available. The reaction of TOL to toluene (T) and the reactions of
TOL
T
k l i'"
MCHA
MCHE
~ H~
~ k4
Figure 2. Network of the HDN of o-toluidine.
/MCH
321
MCHA to methylcyclohexene (MCHE) and methylcyclohexane (MCH) (Figure 2) are equivalent to the reaction of OPA to PB and the reactions of PCHA to PCHE and PCH respectively (Figure 1). In addition to NiMo/y-A1203, Mo(P)/7-A1203 was used, because it is a better catalyst for hydrogenolysis, the selectivity for T being higher t h a n with NiMo/y-A1203 [4,8]. The selectivity to MCH should also increase over Mo(P)/7-A1203. This will enable us to determine w h e t h e r direct C-N bond cleavage indeed takes place in the reaction of MCHA to MCH.
2. E X P E R I M E N T A L The NiMo/7-A1203 catalyst used in this work contained 8 wt% Mo and 3 wt% Ni and was prepared by successive incipient wetness impregnation of 7-A1203 (CONDEA, pore volume: 0.5 cm3.g -1, specific area: 230 m2.g -1) with an aqueous solution of (NH4)GMoTOe4.4H20 (Aldrich) and then with an aqueous solution of Ni(NO3)e.6H20 (Aldrich). The Mo(P)/y-A1203 catalyst was composed of 8 wt% Mo and 3 wt% P. This catalyst was prepared by successive incipient wetness impregnation of the same support as used for NiMo/7-A1203: with an aqueous solution of (NH4)6MoTO24.4H20 (Aldrich) and then with an aqueous solution of phosphoric acid (Fluka). Both catalysts were dried in air at ambient t e m p e r a t u r e for 4 hours, then dried in an oven at 120~ for 15 hours and finally calcined at 500~ for 4 hours. The catalysts were crushed and sieved to the desired particle size before use. A sample of 0.050 g of catalyst was diluted with 8 g SiC to achieve plug-flow conditions in the continuous flow fixed bed reactor. The catalyst was sulphided in situ with a mixture of 10 % HeS in H2 at 643 K and 1.0 MPa for 4 hours. After sulphidation, the pressure was increased to 5.0 Mpa, and the liquid reactant was fed to the reactor by means of a high pressure syringe pump (ISCO 500D). All reactions were performed at 350~ unless otherwise indicated. Dimethyldisulphide was added to the liquid feed to generate H2S in the reaction stream. The composition of the gas phase reactant is given in Table 1. Table 1 Gas phase feed composition Compounds Hydrogen Octane Heptane Hydrogen Sulphide Ortho-toluidine Cyclo he xe ne
Partial Pressure (kPa) 4800 134 20 17.5 7.0 4.0
Function Reactant Solvent Internal Standard H2S Reactant Re act ant
The reaction products were analysed by on-line gas c h r o m a t o g r a p h y with a Varian 3800 GC i n s t r u m e n t equipped with a 30 m DB-5 fused silica capillary column (J & W Scientific, 0.32 mm i.d., 0.25 ~m film thickness), a flame ionisation detector (FID) and a pulsed flame photometric detector (PFPD). Space time was defined as 9 = We / nfeed, where We denotes the catalyst weight and nfeed
322 the total molar flow fed to the reactor. Space time (1:) was changed by varying the liquid and gaseous r e a c t a n t flow rates, while their relative ratios r e m a i n e d constant.
3. R E S U L T S AND D I S C U S S I O N 3.1. NiMo/7-AI~O~ catalyst The results of the HDN of TOL show t h a t the hydrogenation of the aromatic ring is the r a t e - d e t e r m i n i n g step of the reaction network (Figure 2). MCHA was hardly detectable, because it reacts so fast to MCHE and MCH; M C H E is an i n t e r m e d i a t e and MCH the final product of the HDN of TOL. About 5% of T is formed directly from TOL (Figure 3). 7000
0.7
9OMA 9MCH 9MCHE
MCH
v
0.35
3500
T J
9
0
T
i
,
4 8 Space Time (g*min/mol)
12
i
4 Time (g,mie/mol) Space
12
Figure 4. Selectivity of TOL products with NiMo catalysts.
Figure 3. Product distribution of TOL reaction with NiMo catalysts at 390~
At low space time the selectivities of MCH and MCHE did not go to 0 and 1 respectively (Figure 4). To determine w h e t h e r this was due to diffusion limitation [10], catalysts were crushed and sieved, and the reactions of TOL and MCHA 0.7
925-35 mesh
0.7
0.6
,, 60-80 mesh x 100-120 mesh
0.6
~r ~0.5
~"~--'-,~"~ '
,r
~0.4
0.4 0.3 i 0
i1"
i
0.3
8
16 Conversion (%)
_. 24
, 32
0.2-
0
8
16 Conversion (%)
24
32
Figure 5a. MCHE selectivity versus Figure 5b. MCH selectivity versus TOL conversion for different NiMo TOL conversion for different NiMo particle sizes, particle sizes. were studied under the same reaction conditions but with different catalyst particle sizes. TOL conversion did not change with the particle size of the NiMo/7-
323
A1203 catalysts, whereas product distribution did change (Figure 5). The selectivity of MCHE increased with decreasing particle size, because with small particles MCHE is more likely to leave the pores before reacting to MCH. Therefore, with smaller particles, the selectivity of MCHE is higher, t h a t of MCH lower. Nevertheless, MCH selectivity did not reach zero selectivity for the smallest particles at low conversion. To confirm t h a t diffusion played an important role, the HDN of MCHA (the precursor of MCHE) and the hydrogenation of cyclohexene were studied. MCHA reacts faster t h a n TOL, and the diffusion effect was, indeed, more severe (Figure 6). MCHE also reacts faster t h a n TOL. Instead of MCHE, cyclohexene (CHE) was added to the reaction feed. CHE has properties similar to MCHE and enabled us to follow its hydrogenation separately from that of the hydrogenation of MCHE in the HDN of MCHA. An effect of particle size and, thus, of diffusion limitation was also observed for CHE (Figure 7). 4000
7000
925-35 mesh 960-80 mesh
9100-120 mesh 9230 mesh 0
3500
0 0
_z ~ " ~ , 2 4 6 Space Time (g*min/mol)
Figure 6. MCHA partial pressure as a function of space time for different NiMo particle sizes.
2000
0
Spa 4 Time (g*mi8/mol)
12
Figure 7. CHE partial pressure as a function of space time for different NiMo particle sizes.
Since the TOL conversion did not change with particle size, the diffusion rate of TOL is higher t h a n its intrinsic hydrogenation rate (k2). The elimination rate of MCHA (k3) and the hydrogenation rates of CHE and MCHE (ks* and ks) are about equal to or higher t h a n the diffusion rate. Thus, in the latter two cases, diffusion is the rate-determining step. These results show that the selectivity of MCH does not go to zero with decreasing space time, because the relatively slow diffusion may enable MCHE to react in the pores. Because of the strong diffusional effects, it is very difficult to determine whether these effects also contribute to the seemingly direct bond cleavage when the NiMo/7-A1203 catalyst is used. A slower chemical reaction is less likely to be hindered by diffusion. Thus, the reaction of MCHA was studied with different particle sizes ofMo(P)/y-A1203, which is less active t h a n NiMo/~A1203.
324
3.2. Mo(P)/y-AI203
catalyst
7000 9 25-30 mesh
mesh
0
'.- 3500
J_ _
O3 13.
0 0
4
8
12
Space Time (g*min/mol)
Figure 8. MCHA partial pressure versus space time for different Mo(P) particle sizes.
The results of the HDN of MCHA over Mo(P)/~,-A1203 (Figure 8) show that diffusion limitation affect the reaction to a lesser extent t h a n over NiMo/y-A1203 (Figure 6) but cannot be neglected over the larger particles. Mo(P)/~-A1203 has lower activity for all reactions: 45% less conversion of TOL, 30% less conversion of CHE and 25% less conversion of MCHA. With TOL as the reactant, Mo(P)/~A1203 shows a higher selectivity for C-N cleavage t h a n does NiMo/yA1203; the selectivity for T increased from 5% to 22%. Constant selectivity reaction. This means that, under these hydrogenolysis (kl) of TOL are parallel
(Figure 9) is typical of a parallel conditions, hydrogenation (k2) and reactions. MCH selectivity increases with TOL conversion for both catalysts, albeit stronger for Mo(P)/y-A1203 (Figure 10). The selectivities at space time zero are more or less the same for both catalysts. 0.6
0.3 MoP _-_-.
>,
,~ _,
-
>,0.4 ,~
._>
~
'5 0.15
'~
o~ 0.2
NiMo
oO
--
9
9
9
NiMo
"
0
0 6 12 Conversion (%)
18
Figure 9. T selectivity as a function of conversion of TOL for NiMo and Mo(P) catalysts (230 mesh).
0
5
10 15 Conversion (%)
20
25
Figure 10. MCH selectivity as a function of TOL conversion for NiMo and Mo(P) catalysts (230 mesh).
MCHA is formed during the reaction of TOL, but its concentration is very small, because MCHA adsorbs strongly and reacts very fast. A comparison of the selectivities of MCH in the TOL reaction and of MCH in the MCHA reaction provides more information about the MCHA to MCH pathway. MCH selectivity in the TOL reaction is calculated as MCH/(MCH + MCHE). As shown in Figure 11, the trends of the two MCH selectivities are completely different. In the case of TOL HDN, the selectivity of MCH increases steeply with conversion, but in the case of MCHA the selectivity of MCH remains constant until about 80% conversion. This is because MCHA has a much higher adsorption
325 constant t h a n TOL. Thus, MCHA strongly reduces the conversion of MCHE, so t h a t it cannot react to MCH. Therefore, the MCH, which is observed 0.8 below 80% MCHA conversion, can be formed only by a direct route, which 0.6 TOL does not go through the i n t e r m e d i a t e MCHE. Simultaneous reactions of ~ 0.4 TOL with CHE and of MCHA with 0) 03 0.2 CHE were performed to prove this conclusion (Figure 12). The conversion of CHE is not reduced by the presence 25 50 75 100 Conversion (%) of TOL, whereas in the presence of 20 kPa MCHA the adsorption is so strong t h a t the CHE conversion r e m a i n s at Figure 11. MCH selectivity of Mo(P) zero for x < 4 g*min/mol. As expected, in the reaction of TOL and MCHA (230 mesh). this effect at lower MCHA pressure (7 kPa) is less severe. At high partial pressure of MCHA, paths k5 and k*5 (Figure 13) are strongly decreased; under these conditions, the reactions k3 and k4 of MCHA behave like parallel reactions. Thus, MCH can be formed only directly from MCHA and, as a consequence, its selectivity is independent of space time. From this result, we conclude t h a t the direct reaction from MCHA to MCH does, indeed, occur and is not determined solely by diffusion limitation.
/
.....
T . . . . . . .
,
$
4500 CHE
.......... ,"-
CH
3000
v
MCHE
d~
n
15oo
90.07 bar MCHA 90.07 bar Tol
0
MCHA
%
"'-..
ks
. . . . . . . . . . . . . . . . . . . .
0
4 8 Space Time (g*min/mol)
12
Figure 12. CHE partial pressure versus space time under different conditions (230 mesh).
k4
"
[~~MCH
Figure 13. MCHA H D N network and simultaneous reaction of CHE.
The questions r e m a i n as to how direct C-N bond cleavage takes place and w h e t h e r the reaction of MCHA to MCH is similar to the reaction of TOL to T. There are two m e c h a n i s m s t h a t may explain the reaction of TOL to T: hydrogenolysis and partial hydrogenation of the aromatic ring (only one double bond) followed by ~-Hofmann elimination. For the reaction of MCHA to MCH there are also two possibilities: hydrogenolysis, such as TOL to T, and nucleophilic substitution of the amine group by SH group, followed by hydrogenolysis of the C-S bond, which is known to occur with relative ease [11,12].
326
If these two pathways are similar, then their selectivities and rates (T from TOL, and MCH from MCHA) should be higher for the Mo(P)-catalyst. When the catalyst can break the C-N bond of TOL to form toluene, which is a stronger bond t h a n the aliphatic single C-N bond in MCHA, it should certainly be able to split the C-N bond of MCHA to form MCH. Indeed, the selectivity of T for Mo(P)/~Al~O3 is 22% and that for NiMo/~,-A1203 only 5%. When the reaction starts from MCHA, the selectivity of MCH is 21% with Mo(P)/~,-A1203 and 11% with NiMo/~,Al~O3. Furthermore, the rates at which MCH and T form are higher for Mo(P)/~,Al~O3. For the MCHA reaction, the rate is 1.3 times higher and for the TOL reaction 2.5 times higher for Mo(P)/~-A1203 t h a n for NiMo/~-A1203. Hydrogenolysis may be the reason for both pathways, but it is still impossible to reach a definite conclusion about the mechanism involved in those two pathways. An alternative explanation for the direct pathway from MCHA to MCH, nucleophilic substitution of NH2 by SH followed by C-S hydrogenolysis, might be checked by changing the H2S pressure and determining its effect on the pathway. 4. C O N C L U S I O N In complex networks, such as those of o-toluidine and quinoline, it is very important to ensure that all single pathways remain unaffected by diffusion limitation. In HDN catalysis, several reactions, such as NH.~ elimination of MCHA, are very fast. Thus, it is not possible to rely on the determination of the kinetic constants without making accurate checks for diffusion limitation. It has been proven that there is a direct pathway from MCHA to MCH. Thus, the pathway from PCHA to PCH must also exist. Although a comparison of the results for TOL to T and for MCHA to MCH does not eliminate the possibility of a hydrogenolysis mechanism in both reactions, other possibilities must also be considered.
5. R E F E R E N C E S
5 6 7 8 9 10 11
B. S. Clausen, H. TopsQe, R. Candia, J. Villadsen, B. Lengeler, J. AlsNielsen and F. Christensen, J. Phys. Chem., 85 (1981) 3868. G. Perot, Catal. Today, 10 (1991) 447. C. N. Satterfield and S. H. Yang, Ind. Eng. Chem. Process Des. Dev., 23 (1984) 11. S. Eijsbouts, J. N. M. van Gestel, J. A. R. van Veen, V. H. J. de Beer and R. Prins, J. Catal., 131 (1991) 412. M. Jian and R. Prins, Stud. Surf. Sci. Cat., 101 (1996) 87. M. Jian, F. Kapteijn and R. Prins, J. Catal., 168 (1997) 491. M. Jian and R. Prins, Ind. Eng. Chem. Res., 37 (1998) 834. R. Prins, M. Jian and M. Flechsenhar, Polyhedron, 16 (1997) 3235. M. Jian and R. Prins, J. Catal., 179 (1998) 18. J. J. Carberry, Chemical and Catalytic Reaction Engineering, McGrawHill, New York, 1976. G. Perot, G. Brunet, C. Canaff and H. Toulhoat, Bull. Soc. Chim. Belg., 96 (1987) 865. C. Moreau, J. Joffre, C. Saenz and P. Geneste, J. Catal., 122 (1990) 448.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
327
Theoretical Study of benzothiophene hydrodesulfurization on MoS2 S. Cristol 1, J.F. Paul 1, E. Payen 1, D. Bougeard 2 J. Hafner 3 and F. Hutschka 4 1Laboratoire de Catalyse H6t6rog6ne et Homog6ne. CNRS / U P R E S A 8010, F59655 Villeneuve d'Ascq. 2Laboratoire de Spectroscopie Infra-rouge et Raman. CNRS / UMR 8516, F-59655 Villeneuve d'Ascq. 3Institut fiir Theoretische Physik. Technische Universit~it Wien, A-1040 Wien. 4Total Raffinage Distribution, CERT, BP 27, F-76700 Harfleur.
Abstract Benzothiophene (BT) and methylbenzothiophene (MBT) desulfurization on the catalytically active MoS2 edge has been studied using density functional theory. The calculation of the stability of the (100) surface as a function of sulfur coverage indicates two potential active sites. The adsorption energies of BT, MBT and hydrogen molecules on these sites were calculated. The calculation of the binding energies of the various hydrogenated intermediates allow us to the build of energy profiles of BT and MBT desulfurization. It appears that the most endothermic step is the site regeneration (creation of the vacancy).
1. I N T R O D U C T I O N The hydrodesulfurization (HDS) is industrially performed on CoMo/A1203 or NiMo/A1203 catalysts which consist of MoS2 nanocrystallites well dispersed on an alumina support and promoted by Co or Ni atoms [1,2]. It is well admitted t h a t the active sites are located at the edges of these disulfide crystallites. Most of the theoretical work published on HDS is devoted to thiophene desulfurization. Neurock and Van Santen [3] proposed on the basis of DFT calculations on small nickel sulfide clusters that the thermodynamic limiting step is the C-S bond scission. Raybaud [4] showed from a periodic MoS2 model t h a t the most endothermic step is the creation of the vacancy on the catalyst surface.
328
These results can explain the promoting effect of nickel or cobalt as discussed in ref. 4, but they do not take into account the existence of steric effects that are known to be important in deep desulfurization. Finally, DBT has a much more pronounced aromatic character than thiophene and we have to assess how this point affects the reaction path. In this work, we report density functional calculations of the adsorption of benzothiophene (BT), methylbenzothiophene (MBT) and their hydrogenated derivatives on various MoS2 surfaces. These molecules could be good models to estimate the differences between substituted and unsubstituted molecules and so between DBT and DMDBT. 2. C O M P U T A T I O N A L M E T H O D The periodic DFT calculations were performed with the Vienna Ab-initio Simulation Package (VASP) based on plane waves [5,6] which allows a good description of the MoS2 surface, by using large supercells (9.48x20x12.294 A~). As shown in a previous study [7], a model containing two layers along the z direction, three rows in the x direction and four in the y direction (Fig.l) is suitable to give a good description of the electronic and structural properties of the perfect (100) MoS2 surface. All over this work, we used a cut-off energy of Ecut = 210 eV, a Methfessel-Paxton [8] smearing with a - 0.1 eV and F point for Brillouin zone integration. The two upper rows and the adsorbed molecules were allowed to relax during the calculation while the two lower ones were kept fixed at the bulk geometry. In order to calculate reliable adsorption energies, the nonlocal functional using generalized gradient corrections (GGA) of Perdew et al [9] was applied. With these settings, the error on the adsorption energies is less t h a n 0.1 eV.
F i g u r e 1 (perfect surface) dark balls: Mo; light balls: S.
F i g u r e 2 (most stable surface)
329 3. S U L F U R C O V E R A G E AND CATALYTIC S I T E S Figure 1 is a representation of the perfect (100) surface showing alternative rows of molybdenum (metallic edge) and sulfur atoms (sulfur edge). Industrial conditions involve the presence of H2 and H2S in the gas/liquid phase, which implies t h a t the surface could be sulfur rich or sulfur deficient. The nature of the surface in the operating conditions therefore depends on the relative chemical potential of H2 and H2S [10]. The energy of the S addition or S removal on both edges can be calculated according to the reactions (1) and the relative stabilities of the various surfaces can then be deduced. Surface 1 + H2 ~-> Surface 2 + H2S
(1)
Our calculations showed, in agreement with a previous study [4] that, in the sulfiding conditions, the most stable surface is obtained by adding three sulfur atoms on the metallic edge (Fig. 2). This surface should be catalytically inactive because the Mo atoms exposed at the surface are saturated. Adsorption of molecules can only proceed on lacunary structures obtained by removing sulfur according to reaction (1). Different coordinately unsaturated sites (CUS) can be created on both edges. The first kind of potential catalytic site is obtained by removing one (site 1) or two (site 2) sulfur atom from the metallic edge of the stable surface. The second one is obtained by removing sulfur atoms 1, 2 and 3 from the sulfur edge of the stable surface (site 3). Other CUS can be created on each edge of the surface, but we only discuss in this work the more stable ones. Table 1 summarizes the creation energy of each selected site calculated according to reaction (1) the stable surface (Fig. 2) being taken as reference. Table1 Stability of different possible catalytic sites. Surface Site 1 Energy (eV) 1.3
Site 2 3.4
Site 3 2.89
4. A D S O R P T I O N O F M O L E C U L E S The adsorption of BT and MBT was studied starting from different configurations: 111 (S) adsorption is possible on each of the aforementioned sites while ~5 (thiophene) and ~16 (benzene) are only possible on site 2. Figures 3 to 6 show the adsorption geometries and the corresponding energies are reported in table 2. It appears that the methyl group does not affect the flat adsorption of MBT whereas it induces a lowering of the energy of the ~ 1 adsorption. This is due
330
to the steric hindrance between the methyl group and the neighboring layer of the MoS2 surface. This steric interaction is more important on site 3 t h a n on the other ones.
Figure 6
Figure 5 BT adsorbed 1"11on site 3
MBT adsorbed 111 on site 3
Table 2 Adsorption energies of BT and MBT on different sites. Eads (eV) Site 1 Site 2 Site 3
111 (S) 111 (S) ~5 (thiophene) 116 (benzene) 111 (S)
BT 0.5 1.1 1.5 1.2 0.8
MBT 0.4 0.9 1.5 1.2 0.4
331
5. E N E R G Y P R O F I L E S FOR BT A N D MBT D E S U L F U R I Z A T I O N Different mechanisms of desulfurization of BT have been proposed in the literature. Different reaction pathways are thus possible as presented on figure 7. The final product of BT hydrodesulfurization is ethylbenzene (EB). This compound can be produced either by styrene (STY) hydrogenation or by 2,3dihydrobenzothiophene (DHBT) desulfurization. Most of the authors detected DHBT during the experiments so they deduced that hydrogenation of the double bond is followed by hydrogenolysis of the C-S bonds [11,12]. On the other hand, on the basis of kinetic measurements, it has also been proposed t h a t BT desulfurization could be the result of two parallel routes, one involving DHBT as intermediate, the other one involving STY [13,14]. In these studies, it was proposed t h a t STY was not detected because its hydrogenation is too fast. This was supported by experiments showing that in the hydrodesulfurization conditions, styrene yielded 100% EB. The mechanism for MBT desulfurization is likely the same, with an overall rate constant three times less t h a n BT [2]. However there are very few literature data on MBT desulfurization. We have investigated the various reaction pathways for benzothiophene desulfurization reported on figure 7. Other intermediates are also possible but they are less stable and they will be discussed elsewhere. Assuming t h a t H atoms are coming from adsorbed H2 molecule on the surface, the hydrogenation steps are investigated in the form of successive atomic H additions. On site 2, H2 dissociates as one S-H and one Mo-H while on site 3 and site 1, only S-H are present. In all cases, H2 dissociation is exothermic (0.3 to 0.7 eV). For the clarity of figure 7, BT or the intermediates are always shown in an ~1 position, but for each step a full geometry optimization was performed. The first hydrogenation produces intermediate 1 (I1). This intermediate is stabilized by flat adsorption on site 2 whereas it is not when the hydrogenation proceed on site 3. This implies the existence of an endothermic step on site 3 t h a t does not appear on site 2. The stabilization is so important on site 2 t h a t isomerization to form orthostyrenethiolate (OSTY) is athermic. Ring opening seems therefore possible only on site 3. In this case, the last step is the C-S bond scission to produce styrene. In the other reaction path, I1 is further hydrogenated to give DHBT. Desulfurization of DHBT can then proceed through two intermediates: orthoethylbenzenethiolate (OEBT) or 2-phenylethanethiolate (PET). There is no difference on a thermodynamic point of view between these two intermediates. Both C-S bond scissions are highly exothermic as is the last one to produce ethylbenzene.
332
Figure 7 Reaction path for BT desulfurization
All the results are summarized on diagram 1 and 2, which give the energy profile, including site creation from the most stable (fig.2) surface. Whatever the reaction pathway, the overall reaction is exothermic. On the molybdenum edge, creation of site 2 is assisted by BT adsorption on site 1. However, the final sulfur removal to create the CUS is more endothermic (1.61 eV) on this edge t h a n on the sulfur edge (1.33eV).
333 3.50 3.00
uxt,
x.,
u
2.50 2.00 1.50
_
1.00 ~ ~ - : N~,:site 3+S+STY
0.50 [ 0.00
,
DHBT ~ %
!
DHB'I' + H2
Stable Surface
-0.50 -1.00 site 3+S+EB -1.50 :
BT path 1 --'--BT path 2 -~-MBT1
Diagram 1 Energy profile for BT and MBT desulfurization on the sulfur edge (site 3) 3.00 BT q 1 on site 2
2.50 2.00
101e
V~BT~5
BT + H2
1.50
DHBTns~ 1.00
0.50 0.00
BT site 1
DHBT+H2
DHBT 111
~
--T
\ Stable surface
-0.50 -1.00
Site I+EB
-1.50
Diagram 2 Energy profile for BT desulfurization on molybdenum edge (site2)
334
F r o m t h e s e results, it can be deduced t h a t d e s u l f u r i z a t i o n of BT a n d MBT is possible on both edges of the MoS2 slab. The direct d e s u l f u r i z a t i o n to produce s t y r e n e is likely to occur on the sulfur edge while h y d r o d e s u l f u r i z a t i o n to produce e t h y l b e n z e n e t h r o u g h d i h y d r o b e n z o t h i o p h e n e could occur on both edges. On d i a g r a m 1 are r e p o r t e d the r e s u l t s o b t a i n e d with MBT, u s i n g t h e s a m e i n t e r m e d i a t e s , which also correspond to the most stable ones. On t h e m o l y b d e n u m edge, the e n e r g y profile of MBT d e s u l f u r i z a t i o n is a l m o s t t h e s a m e as BT desulfurization. A difference is found in the reaction proceed on the sulfur edge, w h e r e the a d s o r p t i o n e n e r g i e s of the various i n t e r m e d i a t e s are less i m p o r t a n t t h a n for BT ones. The lower d e s u l f u r i z a t i o n r a t e of s u b s t i t u t e d molecules s e e m s to be due to a lower a d s o r p t i o n constant. The sulfur r e m o v a l to produce the catalytic site is always the m o s t e n d o t h e r m i c step. 6. C O N C L U S I O N Different CUS on the surface of the MoS2 crystallites h a v e b e e n e v i d e n c e d on which BT a n d MBT d e s u l f u r i z a t i o n can proceed. It a p p e a r s t h a t both edges can p a r t i c i p a t e to the reaction a n d the limiting t h e r m o d y n a m i c step is the r e g e n e r a t i o n of the CUS. The e n e r g y cost of this step will d e p e n d on the c h e m i c a l p o t e n t i a l of H2 a n d H2S. Studies are in p r o g r e s s to t a k e this point into account by a c h e m i c a l p o t e n t i a l analysis. Acknowledge me nts This w o r k h a s b e e n p e r f o r m e d w i t h i n the g r o u p e m e n t de R e c h e r c h e E u r o p 6 e n " D y n a m i q u e mol6culaire q u a n t i q u e appliqu6e h la catlyse, l ' a d s o r p t i o n et h l'absorption", s u p p o r t e d by the I n s t i t u t F r a n ~ a i s du P6trole, the C e n t r e N a t i o n a l de la Recherche Scientifique, Total and TU Wien. 1. 2. 3. 4. 5. 6. 7. 8. 9.
H. Topsoe, R. Candia, N.Y. Topsoe and B.S. Clausen, Bull. Soc. Chim. Belg. 93 783 (1984). D. Whitehurst, T. Isoda and I. Mochida, Adv. In Catal. 42 345 (1998) and reference therein. M. Neurock and R.A. Van Santen, J. Am. Chem. Soc. 116 4427 (1994). P. Raybaud, PhD Thesis, Universit6 de Paris VI (1998). G. Kresse and J. Hafner, Phys. Rev. B 47 558 (1993); ibid. 49 14251 (1994). G. Kresse and J. Furthmfiller, Comput. Mat. Sci. 6 15 (1996). P. Raybaud, J. Hafner, G. Kresse and H. Toulhoat, Surf. Sci. 407 237 (1998). M. Methfessel and A.T. Paxton, Phys. Rev. B. 40 3616 (1989). J.P. Perdew, J.A. Chevary, S.H. Vosko, K.A. Jackson, M.R. Pedersen, D.J. Singh and C. Frolais, Phys. Rev. B 46 6671 (1992). 10. S. Cristol, J.F. Paul, E. Payen, D. Bougeard and F. Hutschka, to be published. 11. V.H.J. De Beer, J.G.J. Dahlmans and J.G.M. Smeets, J. Catal. 42 467 (1976). 12. R. Bartsch and C. Tanielian, J. Catal. 35 353 (1974). 13. E. Furimsky and H. Amberg, Can. J. Chem. 54 1507 (1975). 14. L. Shi, K.C. Tin, N.B. Wong, X.Z. Wu and C.L. Li, Fuel Sci. Tech. Int. 14 767 (1996).
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
337
Effects of Alumina-Titania Supports on the Activity of NiMo Catalysts Jolanta R. Grzechowiak a, Jacek Rynkowski b, Iwona Wereszczako-Zielifiska a aInstitute of Chemistry and Technology of Petroleum and Coal, Technical University of Wror 50-344 Wrodaw, 7/9 Gdafiska Str., Poland bInstitute of General and Ecological Chemistry, Technical University of L6d2, 90-924 L6d2, 36 Zwirki Str, Poland
ABSTRACT The objective of the study was to investigate how A1203 or A1203-TiO2 supports might influence the activity of the catalysts. The catalysts were tested in the HDS of thiophene and in the hydrotreatment of diesel oil fractions (to examine their HDS and HDN activity). Characterization of the catalysts was carried out by benzene adsorption, temperatureprogrammed reduction (TPR), X-ray diffraction (XRD) and temperature-programmed desorption of ammonia (TPDNm). Activity measurements showed that a comparable HDS activity of MoNi/AlzO3 and MoNi/A12Oa-TiO2 catalysts can be achieved by increasing the amount of NiO in MoNi/AI~O3.When use made of the A12Oa-TiO2supports, it is possible to reduce the MoO3 content in the catalysts. Of the investigated methods of TiO2 incorporation, the one involving impregnation of A1203 with TIC14 aqueous solution is best suited.
1. INTRODUCTION
In the past few years, sulphur content has become a major parameter of diesel fuel quality. Sulphur content can be reduced below 0.05 wt. % when use is made of catalysts with increased HDS activity. Different supports have been tested in order to obtain more active catalysts for the hydrotreatr~ent and hydrocracking of oil fraction [1, 2, 3]. Among these supports alumina-titania appear to be promising [4]. Many investigations have concentrated on the application of binary oxide systems (AI203-TiO2) to the supports of HDS catalysts, but the results reported are .still controversial. This should be attributed to the various method by which the supports'and the catalysts have been prepared. It is due to the influence of the preparation method that the presence of TiO2 increased or did not increase catalytic activity [5, 6, 7]. The function of TiO2 is explained, among others, by the weakening interaction between MoO3 and A1203 due to the coverage of the A1203 surface by TiO2 [8,9]. Some investigators postulate that molybdena crystallites are more easily formed on titania-alumina than on alumina carriers [10]. It was also found that the reduction of nickel in NiMo/AI203-TiO2 catalysts occurred at a lower temperature than in NiMo/AI203 catalysts [ 11 ]. The objective of the present study was to asses the utility of using aluminium oxide for the preparation of TiO2 containing catalysts in order to increase their activity.
338
2. EXPERIMENTAL
2.1. Catalysts preparation Al203-TiO2 supports containing 15 wt.% TiO2 were prepared using two methods: method A involving .impregnation of A1203 with aqueous solution of titanium tetrachloride (catalysts of series A), and method B, which involved coprecipitation. In the latter, sodium aluminate, titanium tetrachloride and nitric acid were added gradually at the same time in the presence of a portion of ammonium nitrate solution, at pH 9 and 70 ~ an additional amount of nitric acid was used for pH adjustment to 7 (catalysts of series B). The same method was used to prepare aluminum hydroxide, which was the support for the catalysts of series A and series C. The catalysts were obtained by the incipient wetness impregnation method with ammonium heptamolybdate and nickel nitrate solutions. The required amounts of nickel were incorporated by impregnation of molybdenum catalysts. The catalysts prepared via this route contained 12 wt. MoO3 (Cat..A-1, B-l), 12 wt.% MoO3 and 3 wt.% NiO (Cat. A-2, B-2, C2), as well as 16 wt. % MoO3 and 3 wt.% NiO (Cat. A-4, B-4).
2.2. Catalysts characterization The methods of TiO2 incorporation by which the supports have been prepared exert a noticeable influence on the activity, physicochemical properties, texture, reducibility, and phase composition of the catalysts. The activities of catalysts A-2 and catalysts B-2 measured in the test reaction, i.e. in the HDS of thiophene, evidenced that the method of incorporating TiO2 into the support had no effect on the conversion of thiophene (Fig. 1). Differences in activity between the two catalysts were found in the hydrotreating of the diesel oil fractions (0.85 wt. % S; 321 ppm N; distillation: IBF - 220 ~ 355 ~ 90 vol. %). According to their decreasing HDS activity (from 330 to 370 ~ the catalysts can be ordered as follows: catalyst A-2 > catalyst C-2 > catalyst B-2.
Fig. 1 HDS of thiophene (T - 400 ~
Fig. 2 HDS of diesel oil fraction (p - 4 MPa, LHSV- 3 h-1)
Thus, the efficiency of the HDN of the diesel oil fraction over catalysts B-2 (which showed the lowest HDS activity) and over catalyst A-2 (which displayed the highest HDS activity) amounted to 73 % and 69 %, respectively (Table 1).
339
Table 1 LBF and HDN activity (T-360 ~ Cat. C-2 LBF, wt. % 3.3 HDN, % 65
p- 4 MPa, LHSV-3 hl). Cat. A-2 5.8 69
Cat. B-2 3.7 73
......
As it may be inferred from the quantity of the light fraction obtained (LBF: T < 236 ~ the application of the support obtained by impregnation of A1203 with TiCI4 (catalyst A-2) brought about a noticeably increased hydrocracking activity (compared to that of catalysts C-2 or catalyst B-2). Of the inv6stigated catalysts, A-2 was characterized by the highest acidity. The applied method of support preparation (by impregnation of A1203 with TiCI4) did not account for an increase of acidity in the range of strong acid sites (Table 2). Table 2 Physicochemical properties and texture of the catalysts (12 wt. % MOO3, 3 wt. % NiO) C-2 A-2 B-2 232 268 309 Specific surface, m2/g 0.57 0.32 0.45 Pore volume, dm3/kg 0.22 0.27 0.13 1.-3nm 0.17 0.03 0.7 3-5nm 5 - 10 nm 0.07 0.04 0.06 10 - 100 nm 0.08 0.02 0.12 Mean mesopore radius, nm 3.9 2.4 3.7 0.77 0.75 0.82 Total acidyty, mmol NH3/t~ 0.23 0.20 0.17 Weak acid sites TPD~I3<250 ~ Medium acid sites, 250 ~ ~ 0.45 0.38 0.41 Strong acid sites, 450 ~ ~ 0.14 0.17 0.19 ,,,
The differences in the activity of the catalysts should be attributed to their physicochemical properties (Table 2), texture (Figs. 3 and 4), reducibility (Table 3) and phase composition (Fig. 5). Thus, the physicochemical properties and the texture of catalyst B-2 differ markedly from those of catalyst A-2 and catalyst C-2. Benzene sorption shows that catalyst B-2 is characterized by the smallest mesopore volume (1.5 to 100 nm) and the smallest mean pore radius. Pores of a radius ranging between 1.5 and 3.0 nm account for 93 % of the B-2 catalyst surface (Fig. 3). For catalysts A-2 and C-2, the proportion of surface at this pore range amounts to 46 and 61%, respectively. The investigated catalysts were found to differ in their pore volume distributions as well (Fig. 5). Only catalyst. B-2 (with a support prepared by coprecipitation) followed a monomodal pattern of pore volume distribution (its maximum corresponding to the 2 nm pore radius). Impregnation of A1203 with TIC14 solution was a factor affecting the pore.volume distribution for pore radii of up to 3 nm; the proportion of pore volume falling within this range of pore radii was by approxi~nately 10 % smaller in catalyst A-2 than in catalyst C-2 (Table 2). Irrespective of the method by which TiO2 had been incorporated into the support, catalyst A-2 and catalyst B-2 were characterized by a greater reducibility as compared to catalyst C-2 (Table: 3). For the catalysts with TiO2 content, the maximum of the lowtemperature peak was situated at temperatures below 420 ~ Re-oxidation at 700 ~ shifted the low-temperature peak towards higher temperatures (by about 25 ~ Following re-
340
oxidation at 900 ~ the least active catalyst, B-2, showed the most noticeable changes in reducibility; the shift of T~x towards higher temperatures amounting to 75 ~ The reducibility of catalyst A-2 was found to vary to a slightly greater extent than that of catalyst C-2. From the investigation it was found that NiMo catalysts did not undergo reduction at the initial temperature of the activation process (300 ~
Fig. 3. Pore surface and volume distribution. XRD (Fig. 4) of catalyst A-2 and catalyst B-2 evidenced the presence of amorphous TiO2. The method of incorporating TiO2 had an influence on the 7-A1203 phase, which was greater in catalyst A-2. The method by which TiO2 had been incorporated into the support also influenced the phase composition of the catalysts subjected to reduction up to the temperature of 900 ~ The presence of rutile and corundum was evidenced only in catalyst B-2, which was characterized by the highest Mo o content. The study revealed that the reduction of catalysts containing amorphous TiO2 (catalyst A-2 and catalyst B-2) did not contribute to the formation of crystalline MoO2 phase.
341
,
u
i
b
I
.
I
....
/e
- "
/5 ....
l
.e
]
1
+'s
se
A
ss
I
6Q
o
"~
Fig. 4 a - XRD pattern of: 1 - C-2, 2 - A-2, 3 - B-2 catalyst (450 ~ D -y-A1203, x - TiO2 (rutile). b- XRD pattern of catalysts 1 - C-2, 2 - A-2, 3 - B-2 (TPR up to 900 ~ D -7-A1203, x - TiO2 (rutile), o - Mo, $ - A1203, V- MOO2, | - NiO. As shown by the results of the HDS of the Diesel oil fraction over NiMo/A1203-TiO2 catalysts, the increase of the MoO3 content from 12 to 16 wt. % did not raise catalytic activity (Fig. 5). Activity tests for rrrolybdenum catalysts (catalyst A-1 and catalyst B-l) showed that the application of carriers obtained by impregnation of A1203 (catalysts of series A) increased the HDS activity of the catalysts (Fig. 6). 93 91 80
72
~751
=704 .O _ .i.-+
~m4 L_
77
Nm~
o
~m 4 ~550~
~3 O
55
0
8 -r
m45t 121 "r" ,101
1
34O IIA-20A-4IB-2I
370 B-41
Fig. 5. Effect of molybdenum content on catalytic activity in the HDS of oil fraction (T-340 ~ p-4 MPa, LHSV-3 hl).
340
370
Ir I A - 1 1 B - I [ Fig. 6 Activity of molybdenum catalysts in the HDS of oil fraction (T-340~ p-4MPa, LHSV-3ht).
342
3. SUMMARY
A comparison of the activity of NiMo/AI203-TiO2 catalyst (15 wt. % TiO2 in the support) with that of the NiMo/A1203 shows that when support have been prepared by impregnation of A1203 with TIC14 solution, HDS activity can be increased. The catalyst obtained via this rout is also r by a slightly increased HDN activity. The differences in activity between the investigated catalysts can be attributed, in the case of the NiMo/A1203TiO2 catalysts, to the lower volume and surface of pores with radii up to 3 nm, as well as to its higher reducibility. The physicochemical properties and the texture of the MoNi/A%O3-TiO2 catalysts substantiate its application to hydrorefining of oil fractions with higher boiling range temperature. The TiO2 containing supports obtained by the methods described in this paper evidence the presence of the amorphous TiO2 phase. The method of incorporating TiO2 affects the proportion of ~{-A1203 phase and reducibility (position of the maximum of the lowtemperature peak, content of the crystalline MoO2 and Mo phase).
4. REFERENCES [1]
[2] [3] [4] [5] [6] [7]
[s] [9] [10] [11]
J. Ramirez, S. Fuentes, G. Diaz, M. Vrinat, M. Breysse, M. Lacroix, Appl. Catal., 52 (1989) 211 M. Henker, K.P. Wendland, E.S. Shapiro, O.P. Tkachenko, Appl. Catal. 61 (1990) 253 J.G. Weissman, E.I. Ko, S.Kaytall, Appl. Catal. A.: General, 94 (1993) 45 Z.B. Wei, X.. Qin, G. Xiexian, P. Grange, B. Delmon, Appl. Catal. 75 (1991) 179 C. Pophal, F. Kameda, K. Hoshino, S. Yoshinaka, K. Segawa, Catal. Today 39 (1997) 21 J. Ramirez, A. Gutierrezale]andre J. Cat. 170 (1997) 108 Z.B. Wei, X. Qin, G. Xiexian, E.L. Sham, P. Grange, B. Delmon, Appl. Catal. 63 (1990) 305 R.B. Quincy, M. Houalla, D.M. Hercules, J.Catal. 106 (1987) 85 K. Foger, J.R. Anderson, Appl.Catal. 23 (1986) 139 W. Zhaobin, X. Qin, U. Xiexian, P. Grange, B. Delmon, Appl. Catal. 75 (1991) 179 S. Damyanov.a, A. Spo]akina, K. Jinatowa, Appl. Catal. A: General, 125 (1995) 257
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
343
Effect of Light Cycle Oil on Diesel Hydrotreatment Jorge Ancheyta-Jufirez a,b Enrique Aguilar-Rodriguez Gustavo Marroquin-Sfinchez a
a,b
Daniel Salazar-Sotelo a and
aInstituto Mexicano del Petr61eo, Eje Central Lfizaro Cfirdenas 152, M6xico 07730 D.F., MEXICO, FAX (+52-5) 368-9371
bInstituto Polit6cnico Nacional, ESIQIE, MEXICO.
Abstract. A pilot plant study was conducted to evaluate the effect of Light Cycle Oil (LCO) on product quality when it is used together with straight run gas oil (SRGO) as a hydrotreatment feedstock. Experiments were carried out over a commercial catalyst at different reaction temperatures and LHSV at constant reaction pressure and hydrogen-to-oil ratio. The experimental results show that low space-velocities and high reaction temperature are required to reach less than 500 wppm of sulphur content in the product. 1. INTRODUCTION Catalytic hydrotreatment (HDT) is one of the most important stages in the processing of crude oil distillates for fuels and petrochemical feedstock. Depending on their origin, HDT feedstocks contain moderately large concentration of heteroatoms, which are distributed over the whole boiling range and they generally increase in concentration in the higher boiling point fractions
[1]. The uses of middle distillates HDT include predominantly the desulphurization of kerosene, diesel fuel and jet fuel, and its application is growing to meet several needs, including the processing of heavier feeds and the production of high-performance lubricants [2]. New diesel specifications are being introduced around the world. The acceptable limits of sulphur content in diesel fuel have been reduced drastically from 2000 to 500 wppm in many countries due to recent environmental regulations [3]. In view of this, the demand for high-quality middle distillates has grown significantly over the past decade and continues growing in the international market. As a result, refiners world wide have started revamping or optimising the existing middle-distillate HDT units to achieve deep desulphurization [4,5]. Straight run gas oil (SRGO) obtained from atmospheric distillation remains the main source of diesel fuels. Nevertheless, the increasing presence of Fluid Catalytic Cracking units in conversion refineries produces large amounts of Light Cycle Oil (LCO) available for blending with SRGO. The common practice is to feed this LCO in a mixture with SRGO to the hydrotreater. However, LCO has a low cetane index and higher density, sulphur, nitrogen and aromatics. These properties adversely affect the quality of the resulting diesel fuel thus limiting its blending ratio [4,6].
344
The aim of this paper is to determine the improvement of SRGO-LCO blend quality by HDT under typical operating conditions using a catalyst and feedstocks recovered from industrial units. 2. E X P E R I M E N T A L
The SRGO and LCO used in this study were obtained from a refinery in Mexico. LCO contains high sulphur (3.5 wt%) and aromatics content (72 wt%) and low cetane index (26.2) and API gravity (12.8), while SRGO is a representative middle distillate (1.31 wt% sulphur, 33.3 wt% aromatics, cetane index of 54.4 and 32.3~ Three feedstocks were prepared with SRGO and LCO streams: Feed 1:100 vol% SRGO, Feed 2:80 vol% SRGO and 20 vol% LCO, and Feed 3:50 vol% SRGO and 50 vol% LCO. The HDT studies were conducted under steady-state operation in a fixed-bed hydrotreatment pilot plant over a Co-Mo/7-AI203 catalyst (2.4 wt% Co, 10.5 wt% Mo, 1.35 wt% P, surface area of 176 mZ/g and pore volume of 0.51 cm3/g). The hydrodesulphurization of the three feedstocks was carried out at the following operating conditions: temperature, 350-380~ and LHSV: 1-2 h -1. The pressure and hydrogen-to-oil ratio for all runs were 54 kg/cm 2 and 2000 ft3/bbl respectively, using pure hydrogen in a once-through mode. 3. R E S U L T S A N D D I S C U S S I O N
The product quality showed a decrease in sulphur, nitrogen and aromatic carbon contents, and an increase in API gravity and cetane index when the temperature is increased and the LHSV is decreased as can be seen in Figures 1 and 2. Feed 3 required higher reaction temperature and lower LHSV to reach less than 500 wppm sulphur because it has higher sulphur content compared with Feeds 1 and 2. The sulphur content of the hydrotreated product is approximately proportional to the LHSV and hence the contact time. Therefore, the heavier the HDT feedstock the lower the LHSV required. 1200
1800 Feed 3 1400
1000.
Feed 2
3
c.~
800 .
~ 1000
600 .
600
400 340
I
I
I
I
350
360
370
380
Feed 1
200 390
Temperature, ~ Figure 1. Temperature effect on product sulphur content
0.5
I
!
!
1
1.5
2
2.5
LHSV Figure 2. Space velocity product sulphur content
effect on
345
The hydrotreated products for the three feedstocks used in this study presented a yellow-green fluorescence at 54 kg/cm 2 and reaction temperatures higher than 370~ This implies that high temperatures and low pressure favour colour degradation of diesel fuel during HDT reactions. The colour stability is more related to thermal severity than other variables in the hydrotreater. In a low-pressure unit, high reactor temperature can cause severe deterioration of the colour of the product [4]. Therefore, highest activity catalyst should be used to maintain the reaction temperature as low as possible. The composition of the feedstock to be treated in a HDT plant will have a significant impact on the unit performance. The situation is more difficult for feeds containing LCO [6]. The blending of only 20 vol% of LCO in SRGO increases the sulphur content from 1.31 to 1.78 wt%, aromatics content from 33.3 to 42.1 wt% and reduces API gravity from 32.2 to 28.9 and cetane index from 54.4 to 46.5. The cetane index although negatively affected by the presence of 20 vol% of LCO remains high (> 50). However, for Feeds 2 and 3, aromatics content is always greater and cetane index lower than those obtained with SRGO. A group of selected experimental runs, which reached less than 500 wppm of sulphur content at 54 kg/cm 2, is presented in Table 1. It can be seen that low space velocities are required for the three feedstocks. For Feed 1, if LHSV is increased from 1.0 to 1.5, a 20~ increase in reaction temperature will be needed to reach <500 wppm sulphur content in the product. In the case of Feeds 2 and 3, an increase of reaction temperature of 10 and 20~ respectively, compared with SRGO (350~ at the same LHSV (1.0 h-l), is required to reach less than 500 wppm sulphur content in the product.
Table 1 Specific operating condition Feedstock Operating conditions Temperature, ~ LHSV Product properties Cetane Index Sulfur, wppm Nitrogen, wppm Aromatics, wt%
to reach < 500 wppm of sulphur in the product 1 1 2
3
350 1.0
370 1.5
360 1.0
370 1.0
59.1 334 30 29.3
59.0 444 48 27.2
52.4 494 34 35.2
44.0 414 75 48.9
The use of conventional HDS catalyst for HDT of SRGO-LCO blends in traditional middledistillate hydrotreaters would require severe operating conditions such as high temperature, low space velocity and high pressure as was stated previously. Such severe processing conditions generally lead to rapid catalyst deactivation, shorter cycle lengths and reduced throughput. Very highly active HDS catalysts have to be used to achieve deep hydrodesulphurization under normal operating conditions in traditional HDT units without major revamping. HDS intensity can be increased by working at lower LHSV, thus new HDS reactors will have to be added to maintain production capacity.
346
High pressure HDT is one approach to improve diesel fuel quality, however, few refineries have such high pressure plants available and high pressure capacity is extremely expensive to build. Reactor pressure can not be increased in existing plants due to physical constraints of maximum total pressure, thus the purification of make up hydrogen and recycled hydrogen should be used in order to increase hydrogen partial pressure. The use of a new hydrogenation step added to a conventional HDS unit is also a possibility. 4. CONCLUSIONS The effect ofhydrotreament of SRGO-LCO blends on diesel fuel quality has been studied over a commercial Co-Mo/7-A1203 catalyst under typical operating conditions. The data showed that in general product quality improved as the reaction temperature increased or space velocity decreased. Colour degradation of the hydrotreated products was found at high reaction temperature (>370~ and low pressure (54 kg/cm2). Similar sulphur and nitrogen product contents were observed using the three feedstocks by changing the LHSV and reaction temperatures. However, aromatics were always greater and cetane lower for SRGO-LCO blends compared with SRGO. The importance for studying the impact of LCO on product quality when blended with SRGO in HDT feedstocks arises from the need for efficient design and simulation of commercial units and a better understanding of catalyst behaviour and the effect of the process conditions. This also allows a better insight into the behaviour of an existing reactor. New and future specifications of diesel fuels can be achieved through hydrotreatment, however, particular process conditions are required when treating diesel fuels blends because cracked products, mainly LCO, have a very strong impact upon the required operating severity and heat requirements along the reactor. 5. R E F E R E N C E S
1. 2. 3. 4. 5. 6.
J.R. Anderson and M. Boudart, HDT catalysis. Catalysis: Sci. & Tech. Germany. 11 (1996). M.J. Girgis and B.C. Gates, Ind. Eng. Chem. Res. 30 (1991) 2021-2058. J.A. Anabtawi and S.A. Ali, Ind. Eng. Chem. Res. 30 (1991) 2586-2592. J.A. Anabtawi, S.A. Ali and M.A Ali, Energy Sources 18 (1993) 203-214. M.K. Andari, F. Abu-Seedo, A. Stanislaus and H.M. Qabazard, Fuel 75 (1996) 1664-1670. S. Kasztelan, N. Marchall and S. Kressmann, Proc. of the 14th World Pet. Cong. (1994) 19-26.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
347
Effects of H y d r o g e n Sulphide on the Hydrodesulphurization of an Industrial HDS Feedstock in a Fixed-Bed Pilot Plant
Jorge Ancheyta-Jufirez a,b, Enrique Aguilar-Rodriguez a,b, Daniel Salazar-Sotelo Gerardo Betancourt-Rivera a and Germfin Quiroz-Sosa a
a,
a Instituto Mexicano de1 Petr61eo, Eje Central Lfizaro Cfirdenas 152, M6xico 07730 D.F., MEXICO, FAX (+52-5) 368-9371
b
Instituto Polit6cnico Nacional, ESIQIE, MEXICO
Abstract In this work we report the experimental results obtained in an isothermal fixed-bed hydrotreatment reactor using an industrial feedstock consisting of 85 vol% of straight run gas oil and 15 vol% of light cycle oil. All tests were carried out with a commercial CoMo/AI203 catalyst at total pressure of 54 kg/cm 2, LHSV of 1-2 hr l, temperature of 350370~ and constant hydrogen-to-oil ratio of 1800 ft3/bbl. The effect of HzS concentration in gas on product quality was studied in the range 0-10 mol%. The experimental results show that the inhibiting effect of HzS on sulphur and nitrogen content is lower at high temperature. Sulphur and nitrogen removals are strongly influenced by hydrogen sulphide. 1. INTRODUCTION Hydrotreatment process (HDT) is one of the most important technologies in a modem refinery. HDT process removes contaminant materials from petroleum distillates, such as sulphur, nitrogen, olefins and aromatics, by selectively reacting with Hz in a catalyst bed at elevated temperature [ 1]. Hydrodesulphurization (HDS) is the most common of the HDT reactions. The degree to which sulphur can be removed depends on the type of feedstock, catalyst and operating conditions. The HDS reaction results in the production of hydrogen sulphide (H2S). High levels of H2S concentration act as a temporary poison and tend to inhibit the HDS reaction to different extent depending on the type of feedstock and the reaction conditions [2]. Also, excessive H2 purge may be necessary in commercial units to hold HzS in the recycle gas at a reasonable level in order to maintain the hydrogen partial pressure of the reactor [3]. Various studies reported in the literature with model reactions show that HDS reaction is moderately affected by H2S [2,4,5] and that, in complex reaction schemes, involving many hydrogenation steps, the product distribution is strongly modified upon addition of H2S [2,6].
348
However, few studies about the effect of hydrogen sulphide on the hydrotreatment of complex feedstocks, i.e. real gas oils, have been reported. The objective of this study is to investigate the effect of hydrogen sulphide on the hydrodesulphurization of middle distillates recovered from industrial units over a commercial Co-Mo/AI203 catalyst. 2. E X P E R I M E N T A L
The feedstock used in this study was a blend containing 85 vol% of straight run gas oil (SRGO) and 15 vol% of light cycle oil (LCO), recovered from an HDS industrial unit (33.4~ Cetane index of 48.3, 1.5 wt% S, 255 wppm N, 31.2 vol% aromatics). The catalyst used was a commercial available Co-Mo/~,-A1203 sample (12.2 wt% Mo, 3 wt% Co, surface area of 220 mZ/g and pore volume of 0.46 cm3/g). The hydrotreatment was performed in a fixed-bed pilot plant. The reactor was operated in isothermal mode by independent temperature control of a three-zone electric furnace. The hydrotreatment was carried out at constant reaction pressure and hydrogen-to-oil ratio without hydrogen recycle (54 kg/cm 2 and 1800 ft3/bbl respectively). Reaction temperature, LHSV and HzS concentration in gas effects were studied in the range of 350370~ 1-2 h -1 and 0-10 mol%, respectively. 3. RESULTS AND DISCUSSION An increase in API gravity and cetane index and a decrease in sulphur, nitrogen and aromatic carbon contents can be observed when the reaction temperature is increased. The decrease in LHSV resulted in improved product quality (Table 1).
Table 1 Effect of operating conditions on product quality Run no. 1 Operating conditions Temperature, ~ 350 LHSV, h -1 1.5 Product properties API gravity 35.7 Cetane Index 50.7 Sulfur, wppm 435 Nitrogen, wppm 146 Aromatics, vol% 29.4
for 0 vol% of H2S concentration in gas. 2 3 4 5 360 1.0
360 1.5
360 2.0
370 1.5
36.3 51.2 230 33 27.4
36.0 51.1 310 91 28.7
35.6 50.5 376 122 29.0
36.1 51.2 200 58 28.1
The combined effect of temperature and LHSV for different H2S concentration in gas on sulphur and nitrogen contents of the product is presented in Figure 1. It can be seen that the hydrodesulphurization reaction is strongly influenced by hydrogen sulphide.
349
Commercial experiences have shown that this effect is found since 2 mol% of H2S concentration in gas [3]. This occurs because the H2S is absorbed on catalyst surface and it competes on the same site with sulphur compounds in the feedstock [7]. The inhibiting effect of H2S on sulphur and nitrogen contents is lower at high temperature. Sulphur and nitrogen removals are strongly influenced by hydrogen sulphide. At low temperature, the effect of hydrogen sulphide on sulphur content is more important when H2S concentration is modified from 0 to 4 tool% compared with 4 to 10 mol%, as can be observed in Figure 1. Other product properties, such as cetane index, aromatics content and API gravity, are moderately modified by HzS in gas, cetane index changed from 50.1 to 51.2, aromatics from 27.4 to 29.7 and API gravity from 35.2 to 36.3. 250
1100 .............
900
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(
200
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700
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t
300
o
- ......
-==
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100 50
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I
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500
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)
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4. C O N C L U S I O N S The influence of hydrogen sulphide on the hydrotreatment of an industrial feedstock (85 vol% of SRGO and 15 vol% of LCO) was carried out with a commercial Co-Mo/A1203 catalyst in a fixed-bed pilot plant under industrial operating conditions. An increase of reaction temperature and a decrease in space velocity resulted in improved product quality (increase in API gravity and cetane index and a decrease in sulphur, nitrogen and aromatic carbon contents). The effect of H2S concentration in gas (0-10 mol%) on product quality was lower at high temperature. Product sulnhur content was strongly influenced bv hvdro~en sulnhide and nitrogen con Figure 1. Effect of HzS concentration in gas on product sulfur and nitrogen contents q (O) 350~ (e) 360~ (r-q) 370~ (--) LHSV = 1, ( .... ) LHSV= 2. re moaerately moalIlea oy u2b in gas.
350 5. R E F E R E N C E S
1. 2. 3. 4.
R.A. Meyers, Handbook of petroleum refining processes. Mc Graw-Hill (1997). J. Van Gestel, J. Leglise and J.C. Duchet, Appl. Catal. 92 (1992) 143-154. NPRA Q&A, Part B : Hydrotreating, Question 36 (1994). S. Gultekin, S.A. Ali and C.N. Satterfield, Ind. Eng. Chem. Proc. Des. Dev. 23 (1984) 179. 5. M.L. Vrinat, Appl. Catal. 6 (1983) 137. 6. T.C. Ho, Catal. Rev.-Sci. Eng. 30(1988) 117. 7. J. Leglise, J. van Gestel and J.C. Duchet, Symp. on Adv. in hydrotreating Cat. 2 0 8 th ACS Nat. Meet. Washington D.C., Aug. (1994).
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
Catalytic properties o f WS2 catalysts prepared by in s i t u d e c o m p o s i t i o n t e t r a a l k y l - a m m o n i u m thiotungstates
351
of
G. Alonso a, V. Petranovskii b, M. Del Valle r J. Cruz-Reyes r and S. Fuentes d aCentro de Investigacifn en Materiales Avanzados, Chihuahua, Chih., Mfxico. bCentro de Ciencias de la Materia Condensada, Ensenada, B.C. 22800, Mfxico. CFacultad de Ciencias Quimicas, UABC. Tijuana, B.C., Mfxico. dInstituto Mexicano del Petrfleo, Apdo Postal 14-805, C.P. 07730, D.F., Mfxico Abstract
Tungsten disulfide unsupported catalysts obtained by in situ decomposition of tetramethyland tetrabutylammonium thiosalts (TMATT and TBATT) presented better hydrodesulfurization performance than catalysts derived from the ammonium thiosalt (ATT). The reaction rate increased with the size of alkyl group in the precursor, however, no correlation of activity with surface area was observed. Auger analysis revealed that the surface concentration of sulfur and carbon varied with the precursor. The improved performance of WS2 catalysts derived from alkylammonium thiosalts in the HDS of DBT is attributed to the formation of tungsten carbide-sulfide species on the surface. 1. INTRODUCTION A procedure for preparation of unsupported MoS2 and WS2 HDS catalysts [1, 2] by decomposition of ammonium- and tetraalkylammonium thiosalts in the reaction medium, has been reported to give very active sulfides. Catalysts prepared in this way contain certain amounts of carbon, and are described with the general formula MS2.yCz, where 0.01 < y < 0.5 and 0.01 < z < 3.0 and M = Mo or W [1]. A previous work [3], involving tetraalkylammonium thiomolybdate precursors, revealed that change of cation from ammonium to methyl and to butyl increases the surface area and HDS activity of in situ generated MoS2 catalysts. The aim of this work is to investigate the influence of the cation size on the process of in situ formation of WS2 catalysts. The properties of the resulting catalysts are compared with those of catalysts prepared from ATT [4]. The influence of carbon on the catalytic properties of WS2 is discussed. 2. EXPERIMENTAL Tetraalkylammonium thiosalts (R4N)2WS4 were prepared by using an improved version of the method reported by McDonald [5]. More details of the synthetic procedure have been reported separately [6]. The HDS of dibenzothiophene (DBT) was carried out in a Parr Model 4522 highpressure batch reactor. The catalyst precursor (2.0 g), along with the reaction mixture (5 vol. % of DBT in decaline), was placed in the reactor, then pressurized to 3.1 MPa with hydrogen and heated to 623 K with a heating rate of 10 K/min. Characterization of catalysts was performed on samples obtained in situ and recovered after the catalytic tests. The samples were separated from the reaction system by filtration, washed with isopropanol to remove residual hydrocarbons and dried under vacuum before
352 analysis. Specific surface areas were determined by nitrogen adsorption at 77 K using the BET isotherm, in a Gemini 2060 surface area analyzer from Micromeritics. Samples were degassed under flowing argon at 473 K for two hours before nitrogen adsorption. X - r a y diffraction patterns were obtained with a Philips X-Pert MPD diffractometer, using Cu-K~ radiation. The surface composition of the catalysts was determined with a Perkin Elmer PHI 595 scanning Auger electron spectrometer. Measurements of the relative atomic ratios for W, S and C were made without any pretreatment of the sample surfaces. 3. RESULTS The results of surface area measurements are listed in the Table 1. The surface areas of in situ WS2 catalyst formed from ATT are equal to that of the ex situ WS2, indicating that in situ activation does not modify the porous structure of the catalysts. The cation size of the
precursor thiosalts influence surface area significantly but non-monotonously. Table 1. Specific surface areas, initial HDS rate constants and AES surface compositions (normalized relative to W) for in situ-prepared tungsten sulfide catalysts. Data for ex-situ catalyst from ATT are taken from [4]. In situ catalyst S, m2/g k (specific) Surface composition 10 .7 mol / g s S/W C/W ATT 47 8.0 2.25 1.7 TMATT 15 9.0 1.7 0.9 TBATT 57 16.0 1.3 0.7 ex situ ATT 49 5.4 The X ray diffraction patterns of catalysts derived from ATT, TMATT and TMATT are represented in Figure 1. The obtained lines are in agreement with those reported for the poorly crystalline structure of WS2. Estimation of peak halfwidth of the (002) line, for the three catalysts shows that the stacking of layers in the c direction decreases in the order TBATT > TMATT > ATT.
Figure 1. XRD diffraction pattems: a) ATT, b) TMATT and c) TBATT
353
The initial rate constants are reported in Table 1. Data for ATT catalysts from Ref. [4] are included for comparison. The rate constant k is found to increase with the size of the precursor cation R4N+. In situ catalysts prepared from ATT show a higher specific rate constant than ex situ catalysts. The TBATT catalysts presented the highest catalytic activity for the conversion of DBT. Products resulting from the HDS reaction of DBT were biphenyl (BIP), phenylcyclohexane (PCH), dicyclohexane (DCH) and benzene (BEN). Distribution of products depends strongly on precursor composition, that is, on the cation size of the alkylammonium thiosalt (Table 2). ATT catalyst yields DCH as the main product (51%). Both TMATT and TBATT catalysts yield BIP (50 and 46% respectively) as the main product. ATT catalyst also produces benzene by cracking of BIP. The TMATT and TBATT catalysts produce no benzene. The product concentration ratio (PCH + DCH)/(BIP) is taken as the HYD/HDS selectivity for comparing catalyst performance. According to Table 2, catalysts made from TMATT and TBATT have significantly lower HYD/HDS ratios than those of catalysts obtained from ATT. Table 2. Selectivity (%) of in situ prepared tungsten sulfide catalysts, calculated at 14 + 1 % DBT conversion. In situ catalyst derived from ATT TMATT TBATT DCH BIP PCH BEN HYD/HD S DCH/BIP
51.0 32.0 10.5 6.5 1.9 1.6
27.0 50.0 23.0 1.0 0.5
25.0 46.0 29.0 1.2 0.5
Table 1 shows the S/W and C/W atomic ratios on the surface as determined by AES. Both S/W and C/W ratios decrease as the cation size of the precursor thiosalt increases suggesting that the step of decomposition of the precursor is relevant to define the surface species. 4. DISCUSSION The surface area of ex situ and in situ catalysts produced from ATT and TBATT was three times higher than the one of TMATT. However, such tendency is different of that observed for catalytic activity, indicating that the better performance of tetraalkylammonium thiosalts in HDS activity is not related with an increase of contact area. From AES results, the global surface composition for ATT, TMATT and TBATT catalysts is calculated to be W82.25C1.7, W81.7C0.9 and WS1.3C0.7, respectively. While ATT catalysts show excess of both sulfur and carbon, catalysts from TMATT show a sulfur deficiency along with excess carbon. The TBATT catalysts, for their part, exhibit a surface sulfur deficiency that is exactly compensated by the amount of carbon (WS1.3C0.7). These tungsten-carbon-sulfide species may have some structural similarity with tungsten oxycarbides [7, 8]. There is evidence that tungsten oxycarbides with surface composition WC3.300.3 and WC1.300.7 are able to substitute neighboring carbon around W, with oxygen generating new active sites [9, 10]. The total sulfur content measured by Auger electron spectroscopy can be due to both types of sulfur, namely free elemental sulfur (S ~ and divalent stoichiometric sulfur ($2). For ATT
354
catalysts, both types of sulfur are expected because of the decomposition of WS3 to WS2 and So [11]. On the other hand, for the in situ decomposition of TMATT and TBATT, where the formation of WS3 is not observed [12] the amount of sulfur detected by AES consist exclusively of divalent stoichiometric sulfur. Thus, the amount of excess sulfur at the surface of ATT-derived catalysts can be attributed to the remains of elemental sulfur from the decomposition of WS3, while the sulfur deficient surfaces of TMATT- and TBATT-derived catalysts are attributed to the formation of sulfur vacancies. The origin of the carbon detected on the surface of catalysts depends on the type of precursor salt. In ATM-derived catalysts, carbon only comes from the organic media of the reaction. In alkylammonium thiotungstate-derived catalysts it can comes from both the organic media and the precursor. In the first case, the deposited carbon is composed of polymeric arrangements of five- and six-members rings, as well as amorphous carbon from cracking reactions produced on acid sites. In the second case, carbon originates from organic fractions released by the precursors during decomposition. This type of carbon may react with the catalyst during the decomposition step, very probably substituting the sulfur atoms eliminated as H2S, via the following reaction: WS2 + H2yCz ~
WS2_yCz -Jr- yHzS
(1)
Finally, the variation of selectivity among the catalysts can be understood by assuming that the rim-edge model proposed by Daage [13] for MoS2 is valid for WS2. It assumes that both HDS and HYD reactions occur at rim sites while HDS takes place only in edge sites. These geometrical considerations relate the HYD/HDS ratio to the stacking of WS2 layers in particles so that, as the average stacking in the catalyst particles increases, the corresponding HYD/HDS ratio decreases. The X-ray diffraction patterns of these in situ catalysts show a decrease of the (002) diffraction peak width in agreement with the variation of selectivity. Overall, the variations in selectivity relative to layer stacking among the catalysts prepared in this work are consistent with those predicted by the model proposed by Daage. 4. CONCLUSIONS The method of in situ activation of WS2 starting from tetraalkylammonium thiosalts yields catalysts with improved catalytic performance. Modification of the composition of precursor thiosalts influences the surface area. The increased activity of tetraalkylammonium derived catalysts can not be correlated with an increase of surface area. The amount of carbon detected at the surface is related to the decrease of surface sulfur. In TBATT-derived catalysts, which exhibit the best catalytic performance among the prepared sulfides, about a third of the stoichiometric sulfur at the surface has been substituted by carbon. It is proposed that formation of WS• species on the catalyst surface leads to the existence of new surface species, which improve the catalytic activity.
Acknowledgments Fruitful discussions with Dr. N. Bogdanchikova, Dr. L. Morales and Dr. M. Farias during the course of this work are gratefully acknowledged. The authors appreciate the valuable technical assistance of E. Aparicio and G. Soto. The present work has been financially supported by DGAPA-UNAM, through grant No. IN-107696.
355
References
[1] [2] [3] [4] [5] [6] [7] [8] [9] [10]
[11] [12]
R.R. Chianelli and T.A. Pecoraro, US Patent 4,508,847, 1985. T.A. Pecoraro and R.R. Chianelli, US Patent 4,528,089, 1985. G. Alonso, M. Del Valle, J. Cruz, V. Petranovskii, A. Licea-Claverie and S. Fuentes, Catalysis Today, 43 (1998) 117. G. Alonso, M. Del Valle, J. Cruz, V. Petranovskii and S. Fuentes, Catal. Lett., 52 (1998) 55. J.W. MacDonald, G.D. Friesen, L.D. Rosenhein and W.E. Newton, Inorg. Chim. Acta, 72 (1983) 205. G. Alonso, G. Aguire, I.R. Rivero, and S. Fuentes, Inorg. Chim. Acta, 274 (1998) 108. F.H. Ribeiro, M. Boudart, R.A. Dalla Betta and E. Iglesia, J.Catal.,130 (1991) 498. E. Iglesia, J.E. Baumgarten, F. Ribeiro and M. Boudart, J. Catal., 131 (1991) 523. A. Muller, V. Keller, R. Ducros and G. Maire, Catal. Lett., 35 (1995) 65. A. Frennet, G. Leclercq, L. Leclercq, G. Maire, R. Ducros, M Jardinier-Offergeld, F. Bouillon, J.M. Bastin, A. Lofberg, P. Blehen, M. Dufour, M. Kamal, L. Feigenbaum, J.M. Giraudon, V. Keller, P. Wehrer, M. Cheval, F. Garin, P. Kons, P. Delcambe, L. Binst. Preprints and Extended Abstracts of the 10th International Congress on Catalysis, Budapest, 1992, pp. 144 - 146. K. Ramanathan and S.W. Weller, J. Catal., 95 (1985) 249. G. Alonso et al., submitted at Applied Catalysis A. M. Daage and R.R. Chianelli, J. Catal., 149 (1994) 414.
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Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
357
Synthesis, characterization and HDS activity of CoMo / Ah03 catalysts prepared by two ways (impregnation of a sol-gel alumina and complete sol-gel synthesis) F. Dumeignil and J. Grimblot Universit~ des Sciences et Technologies de l~flle / Labomtoire de Catalyse H~t~mg~ne et Homog~ne /B~timent C3 UPRESA 8010 / 59655 Villeneuve dL4s(xtC~dex/France
Abstract Co~Mo-Alg)a catalysts obtained by impregnation of a sol-gel (SG) alumina or by Co and Mo introduction during the S-G procedure are compared with an industrial cat~yst. Due to the presence of weakly sulfidable oxides, the samples entirely prepared by the S-G procedure have not improved conversions for HDS unlike the impregnated ones which particular texture permited high HDS conversion ofthiophene. 1. INTRODUCTION In Europe will be impesed new severe rules concerning the S level One of the less expensive solution is the development of new catalysts able to decrease S level in FCC feeds, representing about 40 % of gazoline and contain 90 % of the total S. We chose two ways of synthesis. The obtained catalysts are compared using an HDS test. 2. EXPERIMENTAL
2.1. Catalysts preparations The first two series were synthesized by impregnation of a S-G altunina [1, 2] with solutions of Co nitrate and ammonium heptamolybdate (AHM). 10 and 14 wt% Mo were used with r = Co/(Co+Mo) up to 0.6. Three series were prepared with introduction of the Co and Mo d u m ~ the S-G alumina synthesis. Two of them (20 and 30 wt% Mo) were prepared with Co nitrate in the hydrolysis water and AHM with the complexant of aluminium p~r. r was taken up to 0.6. The last series was synthesized with various Co p ~ r s (Co acetate, oxalate, steamte, citrate and carbonate) with r = 0.35 and 25 wt% Mo. All the solids were calcined in the same way than the altunina [1].
2.2. Characterization methods and catalytic test Bulk atomic compositions were determined by the "CNRS, Vernaison, France". X-my
358
diffraction (XRD), and laser Raman spectroscopy (IA~) were used to investigate the structure of the samples. Specific surface areas (SSA) were measured by BET method. XPS permitted to obtain information about atomic composition using two models : homogeneous binary mixed oxides [3] and Moulijn and Kerkhov [4]. An HDS test, described elsewhere [1], was performed. 3. RF~ULTS
3.1. Composition The impregnation technique permitted a good control of the composition. The Mo composition of S-G samples was lower than the expected one. In all cases, the C content was 0.3 wt%, the organic precursors being eliminated during the calcination. 3.2. Specific surface area (SSA)
Figure 1. SSA of S-G samples.
Figure 2. SSA of impregnated samples.
The S-G samples with only Mo have higher SSA (Fig. 1) than that of the alumina (- 450 m2.gl). This was att14buted to the lattice distorsion due to the Mo introduction [1]. When adding Co, SSA decrease linearly. SSA of impregnated samples (Fig. 2) are larger for the 10 wt% Mo. The corrected SSA of 10 wt% Mo series is - 100 m2.gI too high. It is suggested that this series has a CoMoO phase which develops its own SS/~ The SSA of the catalysts are not influenced by the nature of the COprecursor used. 3.3. XRD, LRS and XPS XRD and LRS on S-G samples containing only Mo detected no MoO3 [1] whereas Co induces CoMoO4 and MOO3. LRS permitted to distinct [a] and [b]CoMoO~ the former being stable at high temperatures [5, 6]. The slow cooling of the solid and the S-G procedure permitted to stabilize it at room temperature but grinding the samples led to [b]CoMoO~ Co acetate was the only pmcm~r giving sample free of any oxide feature. Co citrate and carbonate gave Co304. Impregnated samples were free of any undesirable weakly suUidable oxides. XPS showed a slight sta4ace Co enrichment and Mo distribution was found to be bad
359 except for 10 wt% Mo impregnated series. Indeed, for other series,Mo distribution was affec~d by introduction of Co which induced migration of it at the sta~ace of the grains. On the 30 w% Mo series, big surface crystallites were detected. Then, for the 10 wt% Mo series, a slight Mo surface enrichment, constant with the Co loading, confirmed the peculiar texture of this series. 3.4. Catalytic test
30 _~_ .. _.. 25~," ".~
•
E] " , / -t
20 ~', I ~
9 15 _~ . "~
Z0ne0f hydrogenating
10i !
,
I
catalysts
- II
,
-li
I i~
0
T,
- --~1 I ~
-,'
9
10
~
20
I S o l - g e l (CoMo)
l::! Sol-gel (Mo) -- Co screening i l Impregnated (CoMo) O Impregnated (Mo)
9
0
<>Reference catalyst
30
40
Thiophene conversion (%)
50
F~ttre 3. Butane se]ectiviWas a ftmction ofthiophene conversion. Fig. 3 shows that the best HDS performances are obtained for impregnated CoMo catalysts (conversion superior of 60 % to that of an industrial P-doped catalyst, some catalysts exhibited good performances, one of them being prepared with Co acetate. The dotted line corresponds to catalysts having the HDS/HYD of the reference catalyst (P~F). This ratio is an important factor as large hydrogenating properties lead to saturated hydrocarbons lowering the octane number. Then, it is expec~d to keep acceptable RON after HDS when HYD capability of the catalyst is low. Our best impregnated catalysts are situated on this line and their very good HDS activity would not be therefore detrimental to RON. In the zone of hydmgenathlg catalysts we find Mo / Al~3 catalysts and S-G catalysts prepared from Co dtmte and carbonate. 4. DISCUSSION 4.1. Sol-gel catalysts The low HDS performances are partly due to the precipitation of oxides during the preparation (XRD of dried samples). The SSA after tests are significantly decreased, due to the lattice perturbation (Co and Mo intercalation), and contribute to the bad HDS results but Co acetate seems to be promising as its performance is comparable to that ofthe REF. On Fig. 3, we showed a zone of hydrogenating catalysts with Mo / Al~3 and catalysts prepared from Co citrate and carbonate. All of them should not contain C(~pmmoted sims. This suggestion is
360
in good agreement with the XRD patterns which showed the presence of Co30, This is consistent with a weak Co-Mo interaction. 4.2. Impregnated catalysts These catalysts gave very high HDS conversions in good agreement with absence of undesirable oxides. The 10 wt% Mo series seems to have a peculiar morphology. Indeed, SSA of the oxide pmcxa~rs are consistent with the presence of a weakly linked CoMoO phase. This is confirmed by the SSA at~r test which are constant for the 14 wt% Mo series whilst it decreases with the Co loading for the 10 wt% Mo one. We suppase that on the 14 wt % Mo series, some Co are lost in the almnina lattice (C~) to reinforce. For the I0 wt% Mo series, Co upon its progressive int~xtuction, interacts with Mo to form a CoMoO-like independant giving a weakly linked CoMoS phase, like the CoMoS([I) phase [7, 8]. Then, difference between the two series could be due to the presence of two diiferent CoMoS phases, CoMoS(l])-like and CoMoS(I)phases, the former being more active than the latter. CONCLUSION Complete S-G preparation did not give catalysts with enhanced HDS activity because of undesirable oxides but the use of Co acetate is promising. Impregnation of CoMo solution led to sulfided catalysts with very high HDS performances and good HDS/HYD ratio. This can be explained by the presence of a CoMoS(I]) phase for low Mo loadings and a CoMoS(1) phase, certainly very divided with small MoS2 slabs [9], in the case of higher Mo loadings.
Acknowledgements This work was fimded by EEC through the Joule-Thermie HI program. 5. REFERENCES 1 L. Le Bihan, C. Mauchauss6, L. Duhamel, J. Grimblot, E. Payen, J. Sol-Gel Sci. TechnoL, 2, 837 (1994) 2 L. Lebihan, Doctoral thesis, Lille, France (1997) 3 J. Grimblot, L'Analyse de Surface des Solides, Masson, Paris (1995) 4 F.P.J.M. Kerkhot~ J./~ Moulijn, J. Phys. Chem., 83, 1612 (1979) 5 E. Payen, Doctoral thesis, Lille, France (1983) 6 E. Payen, M. C. Dhamelincourt, P. Dhamelincourt, J. Grimblot and J. P. Bonnelle, Applied Spectroscopy, 36, 30 (1982) 7 R. Candia, O. Sorensen, J. Villadsen, N. Y. Tops~e, B. S. Clausen and H. Topsoe, Bull. Soc. Chim. Belg., 9311~ 763 (1984) 8 H. Topsoe and B. S. Clausen, Appl. Catal., 25, 273 (1986) 9 S. Kasz~lan, H. Toulhoat, J. Grimblot and J. P. Bonnelle, Appk Catal., 13, 127 (1984)
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
361
On the use of P~i)Mon heteropolyanions for the preparation of Alumina support~ HI~ catalysts A Gritx)val~,P. Bhncha~ a E. Payen~*,M. Foumier~,J J_, Duboisb,J. P, B~lard b a ~ m i D i r e de Catalyse H 6 ~ v g ~ e et Homog~e, URA CNRS 402, Universit6 des Sciences et Technologies de Lille, B~t~C3, 5~x%5Y'flleneuve d'Ascq, C~ex, France ~entre de ~ e m h e s ELF-Antar France, Solaize, France Abstract Alumina supported hydnxiesulfiafmtion (HDS) catalysts were ptepaxed by impregnation with Co or Ni substituted phosphomolytxtate solutions. Their ~ d e n d e s in HDS of thiophene were evaluated and compared with catalysts obtained by impregnation with classical ammonium heptamolytxiate solutions. 1. Introduction
HDS catalysts are obtained by ~ t i o n of an oxidic preoxt~r prepared by impregmtion of a ? alumina with impregnating solutions containing the elements to be deposite& The most c o m m o n ones a r e obtained b y dissolving a m m o n i u m heptamol3dxtate (AHM) in presence of phosphoric acid and Co or Ni nitrate, the presence of P~ioK)~ anions in these solutions has been proposed [1]. The use of ~ (~ouO40 or N ~ o ~ 4 0 ) [2] and Anderson salts [3, 4] as starting material was also reporte& We recently p ~ [5] to use Co or Ni salts of reduced hetempolymolytxtate ~ that allows to increase the Ca~o atomic ratio up to .29 and .33 for respectively phospho and silico HPC. This value is however still below the optimum ratio determined in the classical preparations. In this work we relx~ on the preparation of catalysts by impregnation of an alumina with non reduced ~ H I ~ having a higher ~ o ratio. Indeed it is possible to prepare PYMou H I ~ where Y is a Co or a Ni aton~ The synthesis of the starting material is presented and the chamcterisations of the impregnating solutions and of the oxidic precurmrs are disctms~ ~iophene HDS is used to evaluate the catalysts and the restflts are ~ by reference to the natm~ of the staface oxomolybdate phase of the oxidic pmcxu~r.
362
2. Experin~ntal ~he P-C~Mo based oxidic p ~ r s were prepared by the incipient wetness impregnation method of~'-Alg)a extrudates (Kq/k 250 m2/g and Vp: 0.68 cma/g) with the different impregnating solutions. These latter ones were prepared in order to obtain a Mo loading of 14 wt % as MoO~ The impregnated extrudates were dried at 383K overnight and then calcined either in air or nitrogen for 4h at 773 K. The nomenclature of bulk compounds and catalysts is directly derived from the H ] ~ use& The substituted Keggin anion, i.e. ~ i ) M o 1 1 0 ~ - I ~ will be referred hemaiter as PYMon where Y is Co or Ni_ For comparison purposes a Cx~oP catalyst was prepared with the conventional impregnating solution with AHM, Hat~4 and ~ O a ) ~ at P and Co loading (2 wt% P205 and 3 wt% ~ ) c o ~ n d i n g to the optimum for this type of prepm~tiorL Starting materials and oxidic precursors were c h a m ~ with various physical techniques namely laser Raman ~ and mp liquid state NMR specCms~pies. The experimental conditions have been reported elsewhere [5]. After activation under a HaS/H2 (l(Yg0) mixture at 673 K for 2h, catalytic activities for thiophene HDS were measured at atmospheric pmsstm~in a flow-typereactor at 3(D~ 3. Results 3.1. The bulk PYMo11 ~ l t s ~lle ~wlmoniuln salt of the VWMo11anion were prepared aomrding to litemtme data [6]. The IR ~ of these solids showed that the keggin structnre is maintained upon the substitntiorr These substimt~ H I ~ have also been e h a m ~ by 31pliquid state NMR speeCm3seopy.The peak of the Keggin HaPMo= is ~ e d at -3.1ppm The ~ of this line observed at 668 ppm and 538 ppm for respectively the Co and the Ni substituted H t ~ oonfinns that the Co and the Ni are included in the smacmm (Paramagnetic effect of Co or Ni). The Co or Ni salts of this ~ O l l anion were obtained through a two s u a v e substitutions a c o o ~ the followingequations: PYMoI10~I(NI-I4)6"["6 [N(CHa)~CI-]- ~ P Y U o 1 1 0 ~ i ( i ( C S 3 ) 4 ) 6 -[- 6 [NH4+,CI-] PYUoliO4~-I(i(CS3)4)6 -b 3 [Y~, 2(CIO4~]-~ (3 Y~, PYMo110~Ia) + 6 CIO~N(CI-h)4) This method allows to obtain directly the impregnating solutions containing th$ions to be deposited without foreign counterions (NO3"andNH~present in classical solutions). We checked by IR that the exchange was complete by analysis of the solutions that can therefore be used direly for the impregnation of the alumina. These aqueous solutions all exhibit the same Raman featmes characteristic of the P~Mol, aniorL A typical Raman specCnun is shown in figure l& The introduction of Co or Ni induces a down ~ of the vs and was vibrational modes of the M(~Ot and MoOMo bonds of the Keggin unit~ This is in agreement with the data of C. RtxrJficm"lli~ielt~eff et al [7] who showed that the extraction of a Mo atom of a PMoL4)40~ ion, ~ the lactmary
363 VMollO~ 7-aD_ion~induces a down shif~of the IR bands c h a m ~
of these modes.
3.2. C 3 m r a c t e ~ t i o n of the catalysts ~he Raman ~ of the dried and calcined under N2 supported ~ O l l salt (fig lb,c) exhibit a main line at 930 cm! This line chamcteSses a not yet well identified ~ e s . If the caldnation is perfurmed under air the spectrum ~ the well described surface polymolybdate (fig ld) with the main line at 952 cm-1 [8]. Upon drying under air an overlapping of these two lines is observed (fig.le). So a decomposition of the starth~g anion, oonfirmed by the line at 1100 cm1 c h a m ~ of a aluminophosphate entities, has c~'urmd on the alumina. This decomposition depends on the atmosphere of calcinatiorL The Raman spectra of the ~ i M o 1 1 or NiaVNiMoll based solids exhibit the same line at 930 CII1"1on the underlying line at 952 cm-1Ofthe clas~] polymolybdate (fig10. XPS analysis of the calcined solids shows a good dispersion of the oxomolybdate phase. This dispersion is preserved upon sulfidation and the Mo3d and Co2p or Ni2p XI~ binding energy differences (550.2 and 625 eV) chamaerise respectively the well-known ~ o S or NiMoS phase [9]. 952
972
~,,t~"
930
952~
IlOOA,
b)
11oolk
c)
~2oo "=,,
~(cm-a)
9
soo j
|
.
4~o
.
.
.
930.'~,.,.~~.~
. . . . . . . . . ,t
1100/9~30
"
=~26o
VCcm'D
e
s~o
e)
|:-
460 "
-
~gure 1: R am an spectra of PYMo11 based catalysts: a) PCoMoll solution, ~ O l l supported: b) dried under N~ e) calcined under N~ d) calcined under Air, e) dried under air, f) ~ N i M o l x o r NiaPNiMou s u p p o ~ calcined under Ng. 3.3. Catalytic activity Table I shows the df~ciencies of the prepared catalysts. The synergy effect between cobalt or nickel and molybdenum is also observed with the substituted hetempolymolybdate based catalysts. Upon calcina~on under air the activity of the ~MOll solid is similar to the one of the reference CoMoP solid in spite of the lower ~o atomic ratio. But upon calcination under N2 an increase is observed that can be
364 ~ted tDthe presence on the N2 calcined oxidic precurmr of a new surfa~ oxomolybdate phase ~ in Raman spectmsmpy by the line at 930 cm-~. The lack of counterions that allows a better intEa~cfion between the promotDr and the oxomolybdate entities, should also ~ t e ~ the activity impmvement~ Similar improvement was also observed with s u p ~ silidum based H t ~ ( ~ C a M o l l ) , the r precursor of which presents the same ~ oxomolybdate phase.
Catalyst H.eMo (A) ~"I4)~Oll
a~ ratio 0.088
% Convemion 12 20 39
Catalyst
(N) CorNel(N)
% PromotDz/Mo at. ratio Conversion
0.36 ~PNfiV~I(N) 0.36 CoMoP 29 0.36 Table 1: 2hiophene oanver~_on (D) = Dried; (A) = calcination under air;, (N) = calcination under n i ~
0.36 0.36 0.36 0.4
33 26 39 30
4 Conclusion
This study presents an original way of synthesis of HDS catalysts by using Co (Ni) substituted Keggin Ht~. 1WLxedNi and Co promoted catalysts can thus be obtaine& The activity of these catalysts could be ~ t e d to the nature of the surface oxomolybdate phase of the ~ precursor and/or m the lack of foreign counterions in the impregnating solutions. Works are now in progress to dearly identify the new ~ oxomolybdate phase. References
1. JJ~_tLVan VeerL, P_h_JfVLHendriks, tLP, Andrea, E.J.G2~ Romers, A.E. Wilson, J.Phys. Chem. 94 (1990) 5282. 2. /k Spozhakin~ S. Damyanova, V. Sharkova, D. Shopov, T. Yrieva, Pr~. VI ~ Int~ Syrup. Heterogeneous Catalysis, Sofia, part 1, (1987) 503. 3. S. Damyanova,/~ Spozhakina, D. Shopov, Appl. CataL 48 (1989) 177. 4. AaM_Maitra, N.W. Cant, D.L Wrimm; AppL CataL 48 (1989) 187. 5. /~ Griboval, P. Blanchard, E. Payen, 1~ Fournier, J.L Dubois, Stu& Surf. ScL CataL 1{}6(1997) 181. 6. 1~ Leyrie, ~ Fournier, P, Massart, C.I~ Acad. Sc., Paris, t 273 (1971) 1569. 7. C. Rocchiccioli-Deltchef~ 1~ Thouvenot, J. Chem. Research, (1977), 46. 8. L Le Bihan, P. Blanchard, E. Payen, J Grimblot, 1~ Fournier, J. Chem. Soc. Faraday Trans. 94 (1998), 937. 9. R. Candia, B.S. Clausen, H. Topsoe, J. Catal, 77 (1982),564.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
365
H Y D R O T R E A T I N G WITH MIXED F e - N i S U L P H I D E S
P. Betancourt 1, C. E. Scott 1, J. Goldwasser 1, F. Gonzalez-Jimen6z 1, P. B. Embaid 1, R. Hubaut 2, A. Rives 2. 1. Universidad Central de Venezuela, Centro de Cathlisis, Petr61eo y Petroquimica, Apartado Postal 47102. Los Chaguaramos. Caracas. Venezuela. E-mail: [email protected] 2. Universit6 des Sciences et Technologies de Lille, Laboratoire de Catalyse H6t6rog6ne et Homog6ne, URA CNRS N ~ 402, B~timent C3, 59655. Villeneuve d'Ascq. France.
Abstract
Unsupported Fe-Ni sulphide catalysts have been synthesised and characterised by X-ray diffraction, XPS and 57Fe MSssbauer spectroscopies, and thermo-reduction measurements. Thiophene hydrodesulfuration (HDS) and Vanadyl octethylporphyrin hydrodeporphyrination (HDP) activities were also measured. A 4.5-fold increase in HDP activity is observed in comparison to the activities displayed by the corresponding amount of iron and nickel sulphides. This synergetic effect is ascribed to the presence of pentlandite ((Fe, Ni)gSs). 1. I N T R O D U C T I O N Different petroleum fractions contain undesirable elements such as S, N and metals (Ni, V) which have to be removed. An important body of work has been devoted to the study of the hydrodeporphyrination-hydrodemetallizaation of petroleum residues. In a typical Venezuelan oil residue the metals (mainly in porphyrinic structures) are concentrated after distillation. Generally a cheap scavenger catalyst (iron sulphide) is used in order to protect a second valuable catalyst. Pyrrhotite type iron sulphides have been proposed to be the active in HDM of heavy oils [1,2], with subsequent formation of mixed Fe-V sulphides. In previous works [3,4], we have shown the existence of synergetic effect for unsupported Fe-V sulphides. In this work, the ternary FeNiS system was examined and the catalytic properties were determined for model reactions characteristic of hydrotreatment (i. e. HDS of thiophene and HDP of Vanadyl octethylporphyrin (VOOEP)).
2. E X P E R I M E N T A L
The aim was to produce a series of intimately dispersed mixed sulphides covering the range 0 - 100 % Fe. Aqueous solution of Fe(NO3)3 (purity 99.9%, Aldrich) and Ni(NO3)2 (>98%, Strem Chemicals), were added drop-wise to
366
aqueous ammonium sulphide (20%), and vigorously stirred. The concentrations of both solutions have the appropriate Fe/Ni stoichiometry of the final sulphide to be obtained. The solution was filtered off and the solid product dried a room temperature. The samples were analysed by the "Service Central d'Analyses du CNRS". Specific surface areas of the catalysts were determined by nitrogen adsorption at 77K (BET method), using a Quantasob Jr. (Quantachrome), The samples were outgassed 30 min. at 150~ The thermo-reduction was carried out gravimetrically in purified H J A r with a Sartorius S3DV electrobalance. X-Ray difractograms were recorded using a Siemens D-5000 (30 kV, Cu K~ radiation), on fresh sulphide samples. X-Ray Photoelectron spectra were obtained with an AEI ES 200B spectrometer equipped with an A1 anode (A1 Ka 1486.6 eV, 300 W). The C ls peak was taken as reference at 285 eV. 57Fe MSssbauer spectra were recorded at room temperature in a triangular symmetric mode spectrometer. MSssbauer results were computer fitted, with Lorentzian lineshapes. Thiophene HDS was carried out in a continuos flow reactor, under high pressure, outfitted with a gas cromatograph (Intersmat IGC 131) equipped with a flame ionisation detector for on-line analysis of thiophene and hydrocarbon products. The only products observed were n-butane, l-butene, trans-, cis-2butene and tetrahydrothiophene (THT). Catalyst samples (~200 mg), diluted in SiC, were presulfided in-situ. HDP of VOOEP, was carried out in a high pressure continuous flow system. VOOEP in decaline (Jansen Chimica) containing 2% of dimethyldisulphide, (DMDS, J a n s e n Chimica) to maintain a sulphiding atmosphere, was used as a liquid feed. The concentrations of the porphyrins and their hydrogenated intermediates were calculated from UV-visible spectra applying Beer's law. The experimental conditions have been described in detail elsewhere [3]. Conversions and product distributions were measured after 15 h (steady state).
3. R E S U L T S AND D I S C U S S I O N The atomic composition and specific surface areas of the catalysts, before reaction, are summarised in Table 1. The specific areas of the pure and mixed sulphides vary between 6 to 17 m2g -1, the larger being that of Fe-Ni-S 0.75. For the mixed compounds, there is not larger variation. Bulk MoS2, prepared for comparison, has a surface area of 16 m2g -1. A s u m m a r y of phase identifications made from diffraction patterns of the various catalysts is also given in table 1. Pure iron sulphide corresponds to pyrrhotite as indicated by XRD data. The XRD analysis showed t h a t the nickel sulphide consist of well-formed crystallites of NiS (millerite). The sulphur content of the pure samples is in agreement with this result. The XRD patterns of the mixed FeNi sulphides, reveal a mixture of phases (see Table 1). One of these phases has four sharp peaks for a d-spacing (20) of 3.03 (29.455); 2.90 (30.808) ; 1.931 (47.020) ; 1.775 (51.440), showing the presence of a Pentlandite phase. Our preparation method, for FeNi sulphides, give truly mixed compounds. XPS
367
analysis shows the presence of well sulphided phases, in a g r e e m e n t with above results. Non oxidic phases were observed. The reduction of Ni and Fe are observed, by TPR, to be complete at around 220~ and 387~ respectively. For mixed sulphides, a new peak was observed at 281.8~ We propose t h a t this peak, at higher reduction t e m p e r a t u r e (-281.8~ is indicative of a FeNiS phase, probably Pentlandite. The specific activity in HDS (table 2) shows t h a t pure nickel sulphide (NiS) converts thiophene 1.5 times more t h a n MoS2. On the mixed Fe-Ni sulphides a weak synergy is observed with a m a x i m u m for the FeNiS-0.23, however, the a m o u n t of t e t r a h y d r o t h i o p h e n e decreases with the iron content (NiS exhibited only THT as reaction product). This result shows t h a t a p p a r e n t l y iron in mixed catalysts is responsible for the hydrogenolysis activity, while the nickel has a hydrogenating function in the HDS reaction. The HDP activity sharply increases with Fe, reaches a m a x i m u m for the FeNiS-0.74 catalyst, and then decreases to the value observed for pure iron sulphide. The m a x i m u m of the HDP activity is ca. 4.5 times higher t h a n the sum of the activities displayed by the corresponding a m o u n t of iron and nickel sulphides, clearly pointing out a synergetic effect.
Table 1. Atomic Composition and Surface Areao ~ F e N i sulphides: Catalyst % Fe Surface Bulk Composition Phase Reference .... (atomi c)......Ar ea (m2g-1)..........................(Atom!c) .................................................................................................................... Fe NiS-0 0 5.6 NiS 1.1 Mille rite FeNiS-0.23 23 11.05 F e 0 . s N i 0 . ~ T S Pentlandite, Millerite Pentlandite, Millerite FeNiS-0.49 49 13.85 Fe0.27Ni0.27S Pyrrhotite Pentlandite, Fe NiS -0.74 74 17.43 Fe o.96Nio.33S Pyrrhotite Fe NiS- 1 100 6.38 Fe7Ss Pyrrhotite Table 2.
Specific activities of the Fe__Nisulphides ........................................ c a t a l y s t s "................................................ HDS THT yield HDP Reference % Pent. a 10 s mol/g.s 10 s mol/g-s 109 mol/g.s FeNiS-0 59.7 59.7 1.9 FeNiS-0.23 13.8 64.9 48.7 2.2 FeNiS-0.49 14.7 45.0 17.9 3.8 FeNiS-0.74 59.2 35.5 8.5 5.0 FeNiS-1 0.8 0.02 3.0 MoS2 38.4 0.1 a. % of iron as pentlandite in catalysts determined by MSssbauer Spectroscopy
368
MSssbauer spectroscopy allows us to quantify the proportion of the total iron present in the form of pentlandite (table 2), it reveals that HDP activity increase with the amounts of pentlandite in mixed sulphides. Being the FeNiS-0.75 catalyst the one with higher pentlandite content. The two-site idea of hydrogenation and subsequent hydrogenolysis of porphyrin appears to be generally accepted today [5]. We propose that pentlandite phase has a particular balance of hydrogenating- hydrogenolysing sites, which makes it highly active for HDP of VOOEP, but not for HDS of thiophene. There is important to point out that the hydrogenation step in HDP is (at least in part) a thermal phenomena [6,7], which suggests that the main step is the hydrogenolysis of the porphyrin (N-V bond cleavage), and this role could be assumed by iron sites located in pentlandite phase. The interest of these results from an industrial point of view is that scavenger catalysts could ensure a considerable amount of demetallation in addition to some hydrodesulfuration. 4. CONCLUSION This work has shown that the reactivity and morphology of unsupported FeNi sulphide catalysts are critically dependent on the composition. In particular Pentlandite, (Fe, Ni)gSs, appears to be highly active for HDP of VOOEP but not for HDS of thiophene.
5. A C K N O W L E D G E M E N T S
We are grateful to PICS 324 and CONICIT (projects QF-15 and $1-2698).
6. R E F E R E N C E S
1. Gonzalez-Jimenez F., Constant H., Iraldi R., Jaimes E. and Rosa-B. M., Hyp. Int., 28 (1986) 927. 2. Gonzalez-Jimenez F., Bazin D., Dexpert H., Villain F., Constant H. and RosaB. M., Physica B158 (1989) 215. 3. Scott C. E., Embaid B. P., Gonzalez-Jimenez F., Hubaut R., and Grimblot J., J. Catal., 166 (1997) 333. 4. Betancourt P., PhD. Thesis, Universidad Central de Venezuela (1998). 5. Bonn~ R. L. C., van Steederen P. and Moulijn J. A., ACS, Prepr. Div. Fuel Chemistry, 36(4) (1991) 1853. 6. Rankel L. A., ACS, Prepr. Div. Petr. Chemistry, 26 (1981) 689. 7. F. Vandeneckoutte, R. Hubaut, S. Pietrzyck, T. Des Courri~res and J. Grimblot, react. Kinet. Catal. Lett., 45-2(1991)191.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
369
TPR and NO adsorption studies of Mo, CoMo and NiMo catalysts supported o n A1203-TiO2 mixed oxides L. Cedefio a, J. Ramirez ~, A. L6pez-Agudo b, M. Vrinat c and R. L6pez Cordero d aUNICAT, Fac. de Q. UNAM, Cd. Universitaria 04510 Mexico D. F., M~xico. bInstituto de Cat~_lisis y Petroleoquimica, CSIC, Madrid, Espafia. cInstitut de Recherches sur la Catalyse, ViUeurbanne, France. dCentro de Investigaciones del PetrSleo, La Habana, Cuba.
Abstract The effect of the support composition on the surface structure of Mo, CoMo and NiMo catalysts supported on TiO2-A120~ mixed oxides was characterized by TPR and FT-IR spectroscopy of adsorbed NO. The incorporation of titania led essentially to an increase in the fraction of easily reduced octahedral polymeric Mo species, this being more pronounced for the promoted catalyst series. Differences in reducibility between the Ni- and Co-promoted catalysts were observed only for the rich-titania based catalysts. From the NO adsorption results, lower extents of sulfidation for both Mo and promoter (Co or Ni) were observed in the intermediate-titania based catalysts. Possible correlation between the NO adsorption results and those previously found for HDS activity are discussed. 1. INTRODUCTION In our previous works the HDS of thiophene and dibenzothiophene over Mo, CoMo and NiMo catalysts supported on AI~03-TiO2 mixed oxides was studied (1). It was found that for unpromoted Mo catalysts the incorporation of titania to the alumina support increased the catalytic activity per Mo atom in the HDS of dibenzothiophene at high pressure (4.5x 10~ Pa), and to a lesser extent in the HDS of thiophene at atmospheric pressure. However, for Ni or Co promoted catalysts, the incorporation of titania to the alumina support led to drops in the HDS activity, for rich-alumina catalysts. It was only in the rich-titanium samples where the activity resulted increased with respect to that for the alumina-supported catalyst. Such changes in activity were ascribed to the loss of Co, or Ni, promoters in the alumina matrix in form of inactive CoAl204, as result of the preferential octahedral site occupation of Ti in the alumina matrix. Nickel being less reactive to alumina than Co showed this effect to a lesser extent. However, not much was said about the state of Mo in the catalysts since the change of support did not allow a clear evaluation of its dispersion by XPS, and the diffuse reflectance UV-VIS characterizations of the
370
state of Mo were obscured by the presence of TiO2. Also, no characterization of the state of Ni was made in the above samples. The aim of the present work was a further characterization of the Mo, CoMo and NiMo catalysts supported on AIg03-TiO2 mixed oxides with temperature-programmed reduction (TPR) and FTIR spectroscopy of NO adsorption in order to obtain a better insight of the catalytic activity of the mentioned catalysts. 2. E X P E R I M E N T A L
The Mo, CoMo and NiMo catalysts used in the present study were the same as those described previously (1). They were prepared by impregnating Ti(x)-A1 mixed oxides [x= molar relation (TiO2 x100)/(TiO2 + A12Os) = 0, 20, 50, 80, 90, 95 and 100], obtained by coprecipitation of A1 and Ti isopropoxides, with appropriate solutions of ammonium heptamolybdate, nickel or cobalt nitrate. The surface concentration of Mo was fixed at 2.8 Mo atoms per square nanometer of support surface and the Ni (or Co)/(Mo+ Ni(or Co)) atomic ratio was fixed at 0.3. TPR of calcined catalysts were conducted using a flow of 25 cc/min of an Ar/H2 mixture (70 vol. % H2) and a temperature rate increase of 10 K/min from r.t. to 1273 K. The infrared studies of NO adsorbed were conducted in a Nicolet 510 FTIR spectrometer. Thin wafers (10 mg/cm 2) of catalyst sample were sulfided with a 15 % vol. H2S in H2 at 673 K in the IR cell prior to NO adsorption. 3. RESULTS AND DISCUSSION The TPR profiles of the oxidic unpromoted Mo catalysts are shown in Figure 1. They exhibited contributions from the reduction of different types of Mo species and also of the TiO2 support. Incorporation of titania to the catalyst produces a gradual shift to lower temperatures of the two main characteristic reduction peaks, at low and high temperature, of supported Mo species (2). It led also to a decrease in the relative proportion of Mo tetrahedral species (high temperature peak) strongly bound to the support, and to an increase in the population of Mo polymeric species and their reducibility. Incorporation of the Co promoter to Mo catalysts favored even more the formation of Mo polymeric species (Fig. 2), and the incorporation of Ni produced more of the highly dispersed octahedral Mo species (Fig. 3). Both promoters induced a decrease in the tetrahedral Mo species, effect that was enhanced with the Ti content. Quantitative results of H2 consumption indicated, in agreement with the above interpretation, that the total reducibility of the catalysts increased with Ti content. However, the catalyst reducibility trend did not coincide with the HDS activity trend, reported in (1), when the Ti content was increased. The reduction of a part of the support masked the reduction of the fraction of polymeric Mo species.
371
w l
~=-
~~=ioo~ 573
Temperature
K
1278
F i g u r e 1. T P R p a t t e r n s of calcin e d Mo/Ti(x)-Al c a t a l y s t s .
578
Temperature
K
1278
F i g u r e 2. T P R p a t t e r n s of calcin e d CoMofri(x)-A1 c a t a l y s t s .
The NO adsorption studies on the sulfided unpromoted Mo catalysts (Figure 4) showed that the two bands of NO adsorbed on Mo (about 1790 and 1700 cm -1) shifted to higher wavenumbers for intermediate Ti contents. These shifts are related to different extents of sulfidation of the Mo species, as the amount of Ti varies, due to differences in their interaction with the support. In spite of these differences, it was found that the variation of total NO adsorbed follows a trend similar to that for the thiophene HDS catalytic activity with Ti content.
~ - - - O
-
~--O 578
Temperature
K
12;73
F i g u r e 3. T P R p a t t e r n s of calcin e d NiMo/Ti(x)-Al c a t a l y s t s .
I X=0 ,~
-
-
-
w
-
.
-
l I
- - .
1900 Wavenumbers
-
._
,.
(cm-X)1400
F i g u r e 4. IR of NO a d s o r b e d on sulfided Mofri(x)-Al c a t a l y s t s .
372
@
x=O D
i [ I'
~ '
: i . . . . . . .~ . . . .
,
|
_
....
|
x=_loo
x=9~
_
,
19'00 W a v e n u m b e r s (cm -1) 14'00
Figure 5. IR of NO a d s o r b e d on sulfided CoMofri(x)-Al catalysts.
F i g u r e 6. IR of NO a d s o r b e d on sulfided NiMofri(x)-A1 catalysts.
In the case of the promoted catalysts, the presence of a band (or shoulder) at high frequency, due to the adsorption of NO on reduced oxidic Mo species, was evident (Fig. 5 and 6), confirming that in the rich-alumina supported catalysts higher interaction of the promoter with the alumina takes place, in agreement with previous findings (1). The observed changes in the position of the NO adsorption bands associated to the promoter also indicate the presence of less sulfided Co and Ni species at intermediate Ti contents. In both cases, a plot of the thiophene and DBT HDS catalytic activity (not shown) versus the frequency shift of the NO band associated to the promoter, showed that the greater the shift to higher wavenumbers, the lower the catalytic HDS activity. This indicates that for the Co or Ni-promoted catalysts it is the state of the promoter which determines the relative HDS catalytic activity. 4. ACKNOWLEDGMENTS The financial support of the UE (Contract no. Cli*CT92-0024) is gratefully acknowledged. R. L6pez Cordero thanks to the DGICyT, Ministry of Education and Science, Spain for the sabbatical grant (SAB95-0270). 5. REFERENCES
1 E. Olguin, M. Vrinat, L. Cedefio, J. Ramirez, M. Borque and A. L6pez Agudo, Appl. Catal. A. General 165 (1997) 1. 2 R. LSpez-Cordero, F. J. Gil Llambias and A. LSpez-Agudo, Appl. Catal., 74, (1991) 125.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
373
Preparation and charac~rization of HNaY-alumina supports and their impregnated Mo catalysts T. Klimovaa, D. Solis~,J. ~ a
and/k l ~ e z Agudob
aUNICAT, Facultad de Quflnica, UNAM~ C& Universitafia, M&dm DiF. 04510, Ms bInstituto de Cathlisksy Petroleoquimica, C~SI.C., Sermno 119, 28006 Madrid, Spain
Abstract A ~xies of H N a Y - ~ supports with different amounts of HNaY zeolite was prepared by peptization metho& It was found that the proa~ of zeolite incorporation, as well as the procedure of aqueous impregnation of molybdenum species cause some de~uminization of HNaY zeolite framework and loss ofits ctyazd]flu'ty as a result o f ~ o O 4 ) ~ formation. 1. I N T R O D U C T I O N
Hydmtreatment catalysts supported on zeo]itesis a subject of ~ t intexest due to the need of less contaminant fuel. It has been found that the zeolite add centers are capable of h y ~ thiophene and that the presence of add ~.ntexs helps the dispersion ofmetals [1]. ~fly, it has been found that the combination of za~te with conventional catalysts is a promising mute to hidmdesu]fiaization of the most mfmaory compounds to HI)S, substituted DBT's, [2, 3]. In the case of liDS catalysts, the 7_t~te must be incorporated into a matrix, normally alumina, to achieve the necessary mechanical strength. However, this p ~ alters the properties of the support in a manner which has not been deafly studie& It is the object of the present study to contribute to the understanding of the performance of Mo HDS catalysts supported on zeolitealumina (~nposites, by charactenzing the changes in taxtural and chemical properties, that o~xtr when different amounts of HNaY zeolite are incorporated into an alumina matrix and when Mo is deposited on such support 2. E X P E R I M E N T A L
Zeolite-alumina supports with 0, 5, 10, 20 and 100 wt, % HNaY zeolite were prepared using as a s m r ~ g matexials a NaY commen~ za~te (SYAIratio = 2.42) and p s e u d ~ ~ t e Catapal B. Before use, the zeolite NaY was 58 % intemhanged with an ammonium acetate solution to obtain HNaY. The impregnation of Mo to the supports was made by the pore volume method using aqueous solutions of ammonium heptamolybdate (pH = 5.5) to obtain 12 wt. % MoOa loading. Impregnated catalysts were dried (373 K, 24 h) and calcined (773 K, 4 h). Supports were c h ~ by the surface a t ~ pore diameter and pore volume (N2 physisorption), crygalline smaaxtm (XRD), FF-I~ framework SYAl ratio (2"SiMAS NMR) and staface acidity (Py FF-IR). The Mo catalysts were c h ~ by textural properties, XRD, t e m p e r a ~ p m g r m m n e d reduction (TPR), UV-visible (DRS) and infrared (FHR) spectroscopy. H ~ the supports and
374
catalysts will be denoted as HNaY(X)-Alg% and Maq-]NaY(X)-Alg~ resp~vely, where X represents the weight percent of HNaY. 3. RF~ULTS AND DISCUSSION 3.1 Supports
Staface areas and pore volmnes of the prepared supports increase with zeolite content as exp~ed from the contributions of the pure ~ t e and alumina components, indicating that no substantial pore blodmge o(~tus in the ze~te during the preparation procedure (Table 1). Table 1 Textural properties of HNaY(~-Alg~ supports Sample
Surface Area (m2/g)
S~r HNaY(0)-AIX~ HNaY(5)-Al~3 HNaY(10)-AIX~ HNaY(20)-AlgSh HNaY(100)-AIX~
Pore Volume (cm3/g)
~
202 217 231 262 582
0 22 33 76 478
V~
Vm~
0.42 0.42 0.42 0.42 0.31
0 0.008 0.014 0.034 0.222
XRD d e ~ the presence offaujasite in the supports with 10 and 20 wt. % o r a t e . When the zeolite is incorporated to the ahmfina, the positions of X my diffraction lines reveal a small decrease in the zeolite unit cell parameter, ao (from 24.54/k for HNaY ~ t e to 24.34 A for HNaY(20)-AIX~ sample), in "dmating a slight slnLrdmge of the framework, cattsed by some dealumination [4]. This effect was thought to be caused either by the acid medium used in the peptization pl~xlure or during the calcination step in the support preparation. e~
J o
i
,/J~\\.\ ........ ....
0 r~
,
.
.
.
,
.
.
.
,
.
.
.
,
.
.
.
,
800 600 w a v e n u m b e r , cm - I
1200
.
.
..~.,.,
,~
-
p~
c
.,--.--- ...... . / , . / / \
/
,,v--,,,J'--'l""
",,
...... .-......... .,..
"\
'" ...... '--,..,....
j \... "'x.
/1-fl
x,,
~-~
b a
1000
............. : 8 ' o
................. : 1 ~ 0
................. - l t " o
chem ical shift,
Figme 1. Fr-IR spectra of a) HNaY, b) HNaY(20)-AIX~ and c) H N a Y ( I O ) - A h ( ~ .
ppm
.............
Figm~ 2. ~Si-MAS NMR spectra of a) HNaY, b) HNaY(20)-AIX~ and c)
HNaY(10)-Ah~.
375
The deakunination p ~ is also suggested by the IR results that show a ~ to higher waven~ of the strucCme sensitive vibration bands [5] with respect to the pure HNaY ~ t e 1). In line with this, the framework SYAlratio, detexnfined on the basis of ~Si MAS NMR spectral data [6],increases from 3.7 in the pure HNaY zeolite to 4.7 in the zeolite-alumina suppose with 20% HNaY (Figme 2). The SYAIincrease in the zeolite framework is known [7] to affect its addiW leading to a decrease in the n u m t ~ of acid sites but increasing their strength. Indeed, the IR restflts of adsorbed Py confirm that addition of HNaY into an alumina matrix results in a smaller number of Bro'nsted and Lewis acid sites, but of greater strength than those in the pure HNaY sample (Table 2). Table 2 BrSnsted acidity of HNaY(X)-Alg~ samples per garn of support
( ~ m o l By) g.1
Sample
T 150 ~ zeolite support HNaY(100)-AIK~ 0.565 0.565 HNaY(20)-AIR~ 0.442 0.088 tEXlaY(10)-A1gh 0.377 0.037 HNaY(O)-AIK)a 0.000 * Temperature ofPy (~sorption
T 250 ~ zeo]ite support 0.307 0.307 0.248 0.049 0.241 0.024 0.000 -
Fraction of sites (s ~ (T=150 ~ 0.54 0.56 0.63 -
It may be conduded that incorporation of HNaY zeolite into an alumina matrix by peptization results in certain changes in the zeolite properties, that should be taken into account when these matexials are used to deposit a c a t a l ~ y active phase such as Mo(h. 3.2 Mo/zeolite-alumina catalysts The impregnation of Mo on HNaY(~-AIKh supports produces a decrease in BET ~ that is more pronounced when the zeolite content in the support is high (Table 3).
mm_a
Table 3 Texawal properties ofMo/HNaY(X)-A1K)3catalysts Sample
SmTace Area (m2/g)
S~ Mo/HNaY(0)-A1K)3 Mo/HNaY(5)-A]K~ Maq-INaY(10)-Alg:h Mo/HNaY(20)-AIR)3 Mo/ItNaY(100)-AIR~
199 199 204 227 338
0 11 18 55 287
Pore Volume (cm3/g) V~
Vm~w~
S~r decrease caused by Mo deposition (%)
0.35 0.35 0.35 0.34 0.18
0 0.003 0.006 0.024 0.133
1.5 8.3 11.7 13.4 41.9
376
This effect may be due to p m ~ zeolite pore blockage by MoOa crystals during drying and calcination of the catalysts. This is supported by the TPR results that indicate an increase in the intensity of the 500 oCpeak, assadated to the reduction ofMoOa species in the Mo/HNaY(X)-Alg)a catalysts. The DRX results show that the faujasite phase is less crystalline in Mo-containing samples, fact that may be considered as evidence of the destruction of the zeolite framework, as a result of the intexaction with Mo species. This effectcan also contribute to the ~ area drop. The FHR and DRS spectra of MdHNaY(X)-Alg)a (when X , 0) indicate that the proportion of tetmhedrally coordinated Mo species in(xeases with HNaY content. TPR data corroborate this observation. The intea~ty of the peak corresponding to the reduction of MoO4z species (at 870 ~ is higher for catalysts containing ze~te. Earlier [1, 8] it was reported the formation of ~ o O 4 ) ~ crystal]ites in Mo catalysts supported on ~ - 5 and NaY ze~tes. The p ~ c e of ~ o O a ) 3 was attributed to several factors, among them: the strong intexaction between Mo species and nonframework AI cations (the formation of which is enhanced, in our case, by the peptization pmce~ used to ineorporate HNaY zeolite into alumina), and the extraction of AI cations from the ~ t e framework by Mo species. It could be concluded thex~ore, that in the case of the HNaY zeolite, the destructive effect of supported Mo spedes on the zeolite framework is very strong and cannot be prevented by the ineorpomtion ofthe zeolite into the alumina m a t ~ 4. CONCLUSIONS It is concluded that both, the pepfization and the Mo aqueous impregnation processes produce serious changes in the composition and crystalline strucCme of the HNaY ~ t e . These changes can be attributed to the zeolite d e a l ~ t i o n pincer, induced by the contact with aqueous solutions and subsequent calcination, and to the formation of aluminum molybdate with destruction of the zeolite framework, after Mo deposition. The incorporation of the zeolite into an alumina matrix is not ~ e n t to prevent the zeolite dete~oration. There is a mutual m(xiification effect between support and Mo species. The change in support pmpe~es during ~ t e incorporation to the alumina matrix affect the c h ~ of the Mo surface species but also, the Mo deposited species affect drastically the textur& stmcCta~ and chemical properties of the original support. REFERENCES
1 2 3 4 5 6 7 8
W.J.J. Welters, G. Vod)eck, H.W. Zandbergen, LJ~/L van de Ven, E~I. van Oers, J.W. de Haan, V.H.J. de Beex and R.A_Van Santen, J. CataL, 161 (1996)819. M.V.Landau, D. BexgexandM. Herskowitz, J. CataL, 158 (1996) 236. M. Yumoto, K. Usui, K. Watanabe, I~ IdeiandH.Yamazaki, CataL Today, 35(1997)45. D.W.BreckandE.M:Flanigen, Soc. Chem.ln&, (1968)47. E. lVLFlanigen, H. Khatami and H. A~Szymanski, Adv. Chem. Ser., 10 (1971) 201. E. Lippnma, lVLMagi,/k Samoson, lV[ Tannak and G. Engelhardt, J. Am. Chem. Soc., 103 (1981) 4992. JI-I.C.vanHooffandJ.W.Rodofsen, Stu.SmfSd. CataL, 85 (1994) 241. Y.Xu, W.Liu, S.T.Wong, LWangandX Guo, CataL Lett., 40 (1996) 207.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
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MODELING OF NATURE AND STRENGTH OF ACID CENTRES IN ULTRASTABLE ZEOLITES AS A COMPONENT OF HYDROCRACKING CATALYSTS
A.V. Abramova 1, Ye.V. Slivinsky 1, Y.Y. Goldfarb 1, L.Ye. Kitaev 2, A.A. Kubasov z 1 A.V. Topchiev Institute of Petrochemical Synthesis, Russian Academy of Sciences, Leninsky prospect, 29,117912, Moscow, B-71, Russia
2 M.V. Lomonosov Moscow State University, Moscow, Russia
ABSTRACT The methods of X-ray diffraction, IR-spectroscopy and TPD of ammonia had been used for investigation of crystal framework structure, nature of surface hydroxyl groups, number and distribution on strength of ultra stable zeolite Y acid centres and its modification,~: by acid and alkaline treatment. In the hydrocracking process of vacuum gas oil on catalysts on the basis of ultra stable zeolite the influence of modifying ways on activity and selectivity of catalysts had been investigated.
1. INTRODUCTION
Concentration Br6nsted and Lewis acid centres and their distribution on strength are the determining factors influencing selectivity and stability of the hydroprocesses catalysts. Ways of control of an acidity: change of Si/A1 ratio, Na20 contents, presence of nonframework inclusions in zeolite cavities, modifying of zeolites by methods of ion exchange, processing by solutions of acids and alkalis, and also various ways of calcination. One of the perspective directions in creation of new hydro cracking catalysts is to use ultra stable zeolite Y as acid components, having the high Si/A1 ratio, low Na20 contents, and capable long duration time in conditions of reaction and oxidizing regeneration to save a crystal structure and acid characteristics. The conventional dealumination ways - dealumination of the zeolite Y ammonium form by such organic complexing agents as ethylenldiaminotetraacefic acid or acetylacetone, and mineral
378
acids, frequently result to zeolite amorphization [ 1,2]. Recently in the field of ultra stable zeolites synthesis have been developed methods based on direct replacement of aluminium with silicon in a zeolite structure, using such inorganic fluorides and chlorides as SIC14. (NH4)2SiF6 and others as dealumination agents [3-5].
2. EXPERIMENTAL In the present paper the water solutions having different concentrations of (NH4)2SiF6have been used for obtaining of ultra stable zeolite Y as dealumination agent. The necessary effect was achieved by sequential combination of dealuminated zeolite on air at 500~ and exchange of residual Na + on NH4 + ions. Next, zeolite with SIO2/A1203 ratio 6,6 (according to chemical analysis), and residual Na20 contents of 0,02 %, have been further modified by processing with the diluted water solutions of HCI or sodium hydroxide. With using of X-ray diffraction and methods the structure of the prepared samples was investigated. A nature of surface hydroxyl groups, number and distribution on strength of acid centres are established under the IR-spectroscopy and temperature programmed desorption of ammonia data. On the basis of dealuminated or acid and alkaline modified zeolites, NiO and MoO3 containing catalysts of vacuum gasoil hydrocracking were prepared. The process conducted at temperature 380-420~ pressure 10 MPa, volume feed rate 1 hr1, the H2/feedstock ratio 1000:1.
3. RESULTS AND DISCUSSION
The results of tests of the catalyst are presented in Table 1. As can be seen from the given data, the feedstock conversion and the gas and gasoline fraction (<180~ yields increase with increasing reaction temperature. The temperature dependence of the yi eld of diesel fraction (180350~ is influenced appreciably by the zeolite modification method. C f note is the higher yield of gasoline fraction and also the higher content of gas in the hydrocracking products on Cat-3 by comparison with the remaining catalysts. Taking account of the fact that the catalysts differ only in the zeolite modification method, an attempt can be made to explain the relationships obtained on the basis of differences in the structure and properties of zeolites. As can been seen from X-ray data, dealumination of zeolite and alteration of ion exchange and calcination stages results to initial reduction of unit cell parameter up to 24,35 A. After acid and alkaline modifying unit cell parameter grows (up to 24,41-24,48 A) a little.
379
Table 1 Results of tests of catalysts based on an ultra stable zeolite in the hydro cracking of vacuum gas oil. Catalyst Way of zeolite Temperature of Yield (%) of liquid fractions with different boiling modifying conversion, ~ temperatures, ~ gas C!-C_4 <180 180-350 ibp-350 Cat-1 380 1.8 7.1 23.6 30,7 400 19.4 39.8 59,2 420 7.4 26.9 42.3 69,2 440 11.4 36.2 45.5 81,7 380 6.1 40.0 46,1 Cat-2 NaOH 400 1.2 13.0 40.0 53,0 420 10.8 30.0 48.0 78,0 440 20.3 29.6 57.2 86,8 Cat-3 HC1 380 0.7 6.3 34.9 41,3 400 0.8 18.5 35.7 54,2 420 12.3 31.5 33.3 64,8 440 18.4 74.0 14.0 88~
Simultaneously in IR-spectra of a crystal framework the shift of a maximum structural - sensitive absorption band from 583 cm1 for NH4Y to 610-613 cm1 for dealuminated samples and some reduction of specified band frequency for samples processed by acid or alkali solution are observed. The X-ray and IRS data confirm the fact of a zeolite crystal framework dealumination. At the same time, the unit cell parameter size and value of a structural-sensitive band maximum can be influenced by A13*in ion exchanged condition. Acid and alkaline processing can partially extract A13§ from zeolite, thus promoting the marked change of the structural characteristics. The X-ray and IRS data of a crystal framework are necessary for adding by the information on a nature of deluminated zeolite hydroxyl groups, as the specific band set of structural OH-groups is characteristic of the ultra stable form. As have been shown by measurements of IR-spectra, zeolite samples have well resolved band - 3740 cm1 and broad absorption in the field of 3500-3700 cm1, for which it is possible to allocate a maximum- 3600 cm1, and also band shoulders at 3560 and 3620 cm1. A similar set of OH-group absorption bands is close to described in the literature for the ultra stable form of zeolite Y, prepared by hydrothermal process. Thus, the set of the received data allows to make the conclusion about formation ultra stable zeolite during dealumination and subsequent treatments; however, the zeolite structure was imposed by some features of acid or alkaline modifying. More distinctly influence of modifying is shown by consideration of acid properties of these samples.
380
The acid properties of zeolites were investigated by the matching of experimental and computational TPD curves and determination of total of acid centres and function of their distribution on activation energies of ammonia desorption. The concentration of acid centres has made 296, 236 and 197 gmol/g for USY-1, USY-2 (processing USY-1 by a solution of alkali) and USY-3 (processing USY-1 by a solution of an acid) zeolites, accordingly, previously calcinated at temperature 500~ Thus, the total number of acid centres i.e. number of centres holding ammonia, at modifying decreases, especially, in case of acid processing. It can be connected with possible partial amorphization of samples, that is appreciable on weak fall of structural- sensitive band intensity - 610 cml in a IR-spectrum, and also with additional dealumination, and removal of ion exchanged aluminium, that results on the one hand in disappearance of a part of centres, and with another - to appearance of stronger centres. The latter finds the reflection in increase of values minimum and maximum activation energies of ammonia desorption after acid and alkaline treatments. Analyzing the TPD data of absorbed ammonia one can conclude, that the initial ultra stable zeolite modifying results in full removal of weak centres with E < 99 kJ/mole, appreciable reduction of medium strength centre quantity with 99 <E < 129 and strong centres E > 129 kJ/mole. The established distinctions in acid properties of zeolite allow to interpete the marked above differences in behaviour of hydrocracking catalysts, though it is necessary to take into account, that the deposition of a nickel and molybdenum compounds can influence acid properties of a zeolite component. The increase in the proportion of medium-strength centres leads at 420~ in growth of an yield of gasoline and gas. Hence, for achievement of a maximum yield of gasoline the hydro cracking process of vacuum gas oil is expedient for conducting at higher temperatures on catalyst, which zeolite component contains mainly acid centres of medium strength. In this case optimum properties has the sample, modified by acid treatment. For increase of diesel fraction yield at high temperatures it is expedient to use the zeolite with the greater share of strong centres mainly of electron-acceptor type. In the carried out research to this requirement satisfies catalyst including zeolite, washed out by alkali solution.
REFERENCES
1. G. Kerr, J. Phys. Chem. 72 (1968) 2594. 2. L.F. Edward, R.V.C. Lovat, J. Chem. Soc. Faraday Trans.1.83. No. 5 (1987) 1531. 3. H.K. Beyer, I.M. Belenykaja, F. Hange, et al., J. Chem. Soc., Faraday Trans.1. 81.No. 11 (1985) 2889. 4. G.W. Skeels, D.W.Breck, Proc. 6th Int.Conf. on Zeolites, Reno, 1984.87. 5. D.W.Breck, G.W. Skeels, A.C.S. Symp. Ser. Washiington.V.218.1983.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
381
H y d r o g e n a t i o n h e a v y oil r e s i d u e s u n d e r 6 M P a p r e s s u r e in m o t o r fuels a n d f e e d s t o c k f o r c a t a l y t i c c r a c k i n g A.S.Maloletnev ~ and U.P.Suvorov b a-Hydrogenation and Catalysis Science Division, Fossil Fuel Institute, Leninsky Prospect 29, 117910 Moscow, Russia b- Hydrogenation of heavy oil residues Department, Institute Oil mid Chemical Synthesis, Leninsky Prospect 29, 117910 Moscow, Russia
Abstract The main task for petroleum refining industry, at present stage is the increase of depth of oil processing with the aim to intensify motor fuels production. One of the ways to solve this problem is to involve heavy, oil residues into processing. In Russia a process for hydrogenation of heavy, oil residues, including high S, V, Ni contents ones, into distillates and metal concentrates under the pressure Hz to 6 MPa had been created on Pilot plant ST-5 ( capacity 25 t heavy oil residues/day ).The design module unit with 1,6 million tonnes of heaD' oil residues capacity for industrial plant had been developed. 1.EXPERIMENTAL Now existed foreign processes (Fig.l) for increasing the depth of oil refinery to 80-85% as a rule are carried out under 15-20 MPa, equipment for which is very expensive and such equipment does not produced in Russia. That is why in accordance with Energetic strategy of Russia on the period to 2005 for increasing the depth of oil refine~', receiving additional quantities of motor fuels and feed stock for catalytic cracking in Fossil Fuel Institute ( F F I ) the process of hydrogenation recovery of heavy oil residues under low hydrogen pressure is worked out (6 MPa).The point of Russian technology for heavy, oil residues hydrogenation is emulsion production from ammonium paramolybdate (NHa)~MoTO24. 4H~O with heavy oil residues water solution in an apparatus of specific design. Then the mixture of produced emulsion with hot H:, contact with feedstock sulfurous compounds and gas will lead to ammonium paramolybdate conversion into MoSz. Medium particle dimensions of the above MoSz are comparable with oil residues molecules dimensions.
382
Figure 1.Licensors of Fir'.eft-Bed Residuum Hydroprocessing Capaci .ty .(BPD) of Operating Units (Worldwide. Total - 35 Operating Units; 1,42 Million BPD Capacit3;) These conversions in the system give the opportunity to reach higher efficiency of Mo catalyst utilization and to lower hydrogen pressure to 6 MPa.Fonned microparticles are comparable with average macro molecular dimension that gives the opportunity to get essentialy higher degree of catalyst utilization in compm'ison with catalyst application in tablets or as finely divided suspension and to use microquantities of the catalyst (less than 0,1 wt.% from feedstock).The laboratory. and Pilot plant investigations with the heav3~ oil residues from Moscow refinea3' plant ( density at 20~ - 0,978 g/cm .~, content fractions with b.p.< 520~ - 18,4 Wt.%, elemental content, Wt.%: C - 85,57; H - 10,94; S - 2,64; N - 0,51; metal content, g/tonne: V-180; Ni-90; cocking capacity-9,4 Wt.%) showed (Tablel), that degree of conversion feedstock in distillate products = 90-95%.The developed technology had been tested both at laboratory FFI mid Pilot plant ST-5 conditions with the powea" up to 25 t/day of heavy oil residues. The process material balance at the Pilot plant ST-5 is given in Table 2.The process takes place at the pressure to 6 MPa, reaction zone temperature to 425-440~ and LHb-~~ 1,0-2,0 h-~.Catalytic additive content in the feedstock is equal to 0,01-0,08 wt.%. The feedstock and catalyst along with circulating hydrogen-containing gas under the pressure to 6 MPa are heated in the pipe furnace. Gasoline fraction ( density 0,769-0,788 g/cm3; iodine number 64-75; sulfur content 0,3-0,4 wt.%; aromatic hydrocarbon content 18-23 wt.%; oc tane number 80-82 ) is used as a saleable motor fuel component aftea" hydrotreatment and catalytic reforming. Diesel fraction ( density 0,850-0,864 g/cm3; iodine number 30-45; sulfur content 0,75-0,80 wt%) after hydrotreatment is used as a saleable diesel fuel component.
383
Table 1 Matea'ial balance hydrogenation of heavy oil residues at 450~ (wt.~:;) at laboratory unit ( yolume reactor 300 cm 3) Fossil Fuel Institute . . . . . Indices ............ Conditions and products Pressure, MPa 6,0 6,0 10.0 LHSV, h -1 1,0 2,0 1,0 Ratio Hz: feedstock, llkg 500 1000 500 INPUT' l.Feedstock 97,7 97,7 97,7 2.Mo catalyst (water solution) 2,3 2,3 2,3 3.Hydrogen from gas on reactions 1,4 1,1 1,5 TOTAL: 101,4 101,I 101,5 OUTPUT" l.Gas (C1.C4 + HzS) 8,0 6,9 7,8 2.Fractions with b.p.,~ C5-180 12,9 10,5 16,4 180-360 22,9 34,4 35,2 360-520 15,0 18,3 15,2 > 520 38,2 ~,5 23,5 3.Water + losses 4,4 ,.,95 3,4 TOTAL: 101,4 101,I 101,5 Fraction with b.p. 350-520~ ( density 0,940-0,950 g/t,nn3; sulfur content 1,30,1,38 wt.%; coking capacity" 0,1 wt.%; content V and Ni < 1,0 wt.'u can be used as a feedstock for catalytic cracking unit. The residue of the process ( fraction with b.p.>520~ ) qualitively practically does not differ from feedstock with the exception of more higher hydrogen content ( on 0,3-0,5 wt.~ ) and is used as a recycle and for metal concentrates extraction. V and Ni are practically completely extracted from the process products.The catalysts introduced into the process is reeirculating system. The catalyst losses are re stored by micro additives.The calculations economic data module unit with annual capacity" to 1,6 million tormes of hea~,aj oil residues for industrial plant are given in Table 3. The results of comparison testil~r the advantages of the hea~,~- oil residues hydrogenation process under 6 MPa in comparison with visbreaking, thus under the industrial realization of the process the structttral ch~uages in assortment of saleable products will rake place. The output of gasolines ( normal and extra trade marks) ,and diesel fuels will increase in 4,4 times. The cost of ,annual output will increase in 1,7 times in comparison with analogous; obtained saleable products
384
will provide profit 18,8 millions US$ and remunerative production. Table 2 Material balance of heavy oil residues hydrogenation at the Pilot plant ST-5 _ Products Indices (wt:%) INPUT' Heaw oil residues 100,0 Recycle ( with b.p.> 520~ 53,85 Hydrogen-containing gas ( 96 % Hz ) 2,79 TOTAL: 156,64 OUTPUT" Hydrogen- containing gas (75 % Hz) 1,00 Gas C1-C4 7,08 H2S 1,00 Fractions with b.p.,~ Cs-180 19,95 180-360 42,02 360-520 30,74 Recycle ( fraction with b.p.> 520~ 53,85 Residue of the process 1,00 TOTAL: 156,64 Table 3 Main economic data module unit hydrogenation of heaw oil residues .for petroleum refinery industrial Plant Data Process hyProcess Indices drogenation visbreakunder 6MPa ing 1600,0 1600,0 Capacity module unit per year, thousands tormes 885,0 198,6 Production of motor fuels per year, thousands tonnes 238,6 87 o including: gasoline 646,4 111,4 diesel fuel 560,0 Production feedstock for catalytic cracking per year, thousands tonnes 98051 22070 Annual demand of energy means, tonnes fuel 193,6 85,0 Calculated value ofconstraction, millions US$ 211,5 1" o Annual production costs, millions US$ 18,8 Conventional profit, millions US$ $ .t...,
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
385
INFLUENCE OF THE NATURE OF THE METAL AND OF THE ACIDITY OF THE SUPPORT ON HYDROCRACKING REACTIONS
Julie-Anne Porta, Joi~l Despr~s and Franqois Garin Laboratoire d'Etudes de la R6activit6 Catalytique, des Surfaces et Interfaces (LERCSI) UMR 7515 du CNRS - ECPM- Universit6 Louis Pasteur 25 rue Becquerel, 67087 Strasbourg Cedex 2 - France e-mail: [email protected] ; Tel: 0033 (0)3 88 13 69 44 ; Fax: 0033 (0)3 88 13 69 68
Abstract Contact reactions of nC6 on sulphated zirconia supported 0.2wt% of Pt or Pd or Ir exclusively give isomerization reactions while from n-C7 only hydrocracking reactions occur. During the C-C bond rupture, n-C6 will give a primary carbenium ion plus i-C4, when with nC7 and upper, at least secondary carbenium ions are formed plus i-C4. Moreover the first step of the mechanisms of hydrocracking involves a "metal-proton adduct" in an additive reaction with the hydrocarbon giving a carbenium ion plus hydrogen. 1. INTRODUCTION The catalytic behaviour of solid acids supported metallic particles are not an easy task to understand. We are going to monitor their comportment in alkane reforming reactions. Several points have to be taken first into account such as: - i) which temperature ranges the metallic particles and the acidic support are the more reactive and -ii) which bonds between C-H and C-C bonds are first ruptured. From a thermodynamic point of view the bond dissociation enthalpies for C-H is stronger than for C-C single bond, 418 KJmo1-1 and 348 KJmo1-1 respectively. Table 1 gives the average range of temperatures in which such bond dissociations are observed. Table 1 Approximate temperatures (~
for the activation of various bonds [ 1]
'Catalysts
C-H bond rupture
C-C bond rupture
Metalllic catalysts
between-80~ to 20~
For Ir catalysts : from 150~ For Pt catalysts : from 220~ For Pd catalysts : from 280~ between 200~ to 400~
Acidic catalysts
between 20~ to 250~
i
Second, geometric and/or electronic changes are claimed to be due to changes in the particle size [2,3]. However, these two possibilities are not necessarily independent; it is possible to observe a modification in the surface topology and simultaneously a change in the electronic structure with increasing particle size. Another factor which might influence both electronic and geometric properties is the metal-support interaction, more specifically the type of bonding between the particle and the surface of the support and also its own acidity strength. Third, such catalysts, solid acids supported metallic particles, are bifunctional catalysts, which term stands for a heterogeneous catalyst that exposes two types of active sites, e.g., acid sites
386
and ensembles of transition metals. At that point several questions may be asked concerning the reforming reactions, -i) Do we may expect any variations in activity or selectivity in function of the metal used, Pt, Pd or Ir ?, - ii) What is the influence of the acidity of the support in such reactions ?, and - iii) What is the bifunctional activity of such systems?. To try to answer these questions a comparison will be done between the catalytic behaviour of neutral y-alumina supported metals (Pt, or Pd, or Ir) which catalysts will be named "metallic catalysts" and the sulphated zirconia supported metals (Pt, or Pd or Ir) named "acidic catalysts" and noted Pt-ZS, Pd-ZS and Ir-ZS. 2. EXPERIMENTAL All the experiments were performed in a plug flow catalytic reactor working under atmospheric pressure. For the "metallic catalysts", the metal content was of 10wt% on yA1203, and labelled hydrocarbons were used to follow the various mechanisms involved. For the "acidic catalyst" a sol-gel method was used already explained in Ref.[4]. The hydrolysis is perfomed in acid medium, the addition of sulphuric acid, dropwise into the mixture of zirconium alkoxide and n-propanol, leads in one step to sulphate alcogel, which after drying gives the sulphated zirconia sample. Then 0.2wt% of the metals is impregnated on the ZrO2/SO42 support from Pt(NH3)4(NO3)2, Pd(NO3)2 and H2IrC16. The hydrogen reduction is performed at 350~ for 1 h. under a H2 flow of 40 cm3.min-1. The hydrocarbons used in this study are: 2-methylpentane (2MP), n-hexane (n-C6), n-heptane (nC7), branched heptanes and n-octane (n-C8). 3. RESULTS 3.1. Behaviour of "Metallic catalysts". The explanation of particle-size effect by the invocation of electronic factors is reiforced by the results of the comparative study made with the various supported metal catalysts of similar dispersion. In Table 2, the distributions of the C6 products and the locations of the carbon-13 in the 3-methylpentane isomers are given for the isomerization of 2-methyl-213C pentane on Pt, Pd and Ir. From these data the percentages of bond-shift and cyclic mechanisms (selective and non selective) can be deduced [5].
Table 2: Reforming of 2-methyl-213C pentane on ),-A1203 supported 10wt%Pt, or Pd or Ir. Met.
Particle T~ size A
Pt
90
% Select. hydrocracking in isomers reactions
254 51
Isomerization Bond Shift
mechanisms Selective a non select a.
statistical C-C 84 7 9 bond ruptures Pd 90 270 30 demethylation 12 0 88 predominates Ir 80 160 30 deethylation 0 100 0 predominates a Selective cyclic mechanism does not give n-C6 from methylcyclopentane hydrogenolysis, when the non selective cyclic mechanism does. (PH2 = Patm and Puc = 5 Torr).
387
The mechanisms involved on such metallic "monofunctional" catalysts take place via •-alkyl, rc-olefinic, carbene and carbyne species. The common intermediate being first an additive entity between adsorbed hydrogen and the hydrocarbon, similar to agostic species [6,7]. 3.2. On sulphated zirconia supported Pt or Pd or Ir the results are as follows:
- i) at 200~ under the same experimental conditions as above, from n-C6, only isomerization reactions take place. The total rate, i.e. isomerisation rate, decreases by a factor of five from Pt to Ir, - ii) starting with n-hexene-1 as reactant, the rate obtained on Pt catalyst is similar to the one get with n-C6, a n d - iii) at 200~ with n-C7 and n-C8 as reactants, only hydrocraking reactions take place.The results, obtained at 200~ are reported in Table 3 which follows. i
ii iii
Reactant
n-C6
% Catalysts Sisom.
Rate x 102
Pt-ZS Pd-ZS Ir-ZS
50 15 10
i
98 97 99
ii
i
i
i
i
n-C7
n-C8
2MP/ 3MP
% Sisom
Rate x 102
C3/ i-C4
% Sisom
Rate x 102
C3+iC5 / iC4
1.6 1.7 1.5
2 1 2
80 80 40
1.1 1.1 1.1
1 1 2
110 80 60
1 1 1
ii
i
ii ii
%Sisom: Selectivity in isomers; Rates are expressed in ~mol(g.s) 1'' Observations: - No cyclic molecules are formed. - From n-C6, the ratio 2MP/3MP is equal to 1.6 + 0.1 on the three catalysts.- From n-C7 and n-C8, only one Carbon-Carbon bond rupture occurs, i.e. the mol fractions are equal: [C3] - [i-C4] and [C3 + i-C5] = [i-C4]. - The values of the various ratio reported in Table 3 are independent of the reaction temperature between 140~ upto 200~ How these hydrocracking reactions proceed: Do we have first isomerization then cracking or the reverse ? We studied the influence of the structure of the alkane on the reactivity at 200~ on Pt-ZS; the results are reported in Table 4. Table 4" Influence of the structure of the alkane on the reactivity at 200~ on Pt-ZS i
Reactants
% Select. in isomers
Rate. 102 lamol(g.s) -~
C3 / i-C4
i
2,4 - dimethylpentane 5 2 - methylhexane 6 3 - methylhexane 5 2,3 - dimethylpentane 4 n - helatane 2
510 245 195 155 80
1.1 1.1 1.1 1.1 1.1
Such results suggest that isomerization reaction takes place before hydrocracking reactions. Whatever the reactant is the same value is obtained for the ratio propane over isobutane. 4. D I S C U S S I O N First, to explain the different behaviour between n-C6, and lower hydrocarbons, on one hand, and n-C7, and upper hydrocarbons, on the other hand, we have to take into account the fact that with n-Hexane we shall get a primary carbenium ion during the Carbon-Carbon bond
388
rupture plus i-C4, when, with n-C7, and upper, at least secondary carbenium ions are formed plus i-C4. Second, in the classical model of bifunctional catalysis, proposed by Mills et al. [8], isomerization of n-alkane is assumed to start with the adsorption of the reactant molecule on a metal site, where it is dehydrogenated to an olefin. This can move to an acid site, where it is adsorbed as a secondary carbenium ion which can isomerize to a tertiary carbenium ion. When this species decomposes into a surface proton and a branched olefin, that molecule is adsorbed at a metal site, where it is hydrogenated to a branched alkane molecule. From our results, we can see that the three catalysts give similar results. Such observation suggests that an adduct entity as [Metal-H] +, already mentioned by the group of Sachtler [9,10], is responsible of the catalytic selectivity. All reaction steps can be realized during one single residence of the molecule. No endothermic decomposition of adsorbed carbenium ions into protons and desorbed olefin molecules is require in this case. If it were the case we should have found different selectivity versus the metal involved. We have already shown the importance of the electronic factor when considering the dominant isomerization mechanism on various metal catalysts. It underlines that the intermediate species have different electronic requirements on platinum, palladium or iridium which explains that bond shift mechanism predominates on large platinum particles while non-selective and selective cyclic mechanisms are favoured on palladium and iridium respectively. With the sulphated zirconia supported Pt or Pd or Ir, these metals seem to be used only to activate either the C-H bond or the proton. I suggest that the latter possibility is more probable, then a "metal-proton adduct" [ 10] is formed followed by an additive intermediate which is initiated with the hydrocarbon and the "metal-proton adduct". The protolysis reaction occurs, liberates hydrogen and a carbenium is formed. The kinetic model already proposed by the group of Frennet [ 11 ] can be applied for both reactions occuring on "metallic" or "acidic" catalysts and can explain the negative and positive values of the orders versus hydrogen. 5. R EF ER EN C ES
9
10 11
F.G. Gault, Gazz. Chim. Ital., 109 (1979) 255 F.G. Gault, Adv. Catal., 30 (1981) 1 G. Maire and F. Garin, in Catalysis, Science and Technology, J.R. Anderson and M. Boudart, Eds., Springer-Verlag, Berlin, 6 (1984) 161 D. Tichit, B. Coq, H. Armendariz and F. Figu6ras, Catal. Lett. 38 (1996) 109 F.G. Gault, V. Amir Ebrahimi, F. Garin, P. Parayre, F. Weisang, Bull. Soc. Chim. Belg., 88 (1979) 475 F. Garin and G. Maire, Acc. Chem. Research, 22 (1989) 100 F. Garin and G. Maire, J. Mol. Catal., 52 (1989) 147 G.A. Mills, H. Heinemann, T.H. Millikan and A.G. Oblad, Ind. Eng. Chem. 45 (1953) 134 X. Bai and W.M.H. Sachtler, J. Catal., 129 (1991) 121 T.J. McCarthy, G.D. Lei and W.M.H. Sachtler, J. Catal. 159 (1996) 90 A. Frennet, G. Lienard, A. Crucq, L. Degals, J. Catal. 53 (1978) 150
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
389
Hydrogenation of heavy oil using carbon-supported catalysts Atsushi Segawa, Katsuya Watanabe, Yukio Shibata and Toshikazu Yoneda Petroleum Energy Center, Advanced Catalysts Research Laboratory, KSP D 12F-1237, 3-2-1, Sakado, Takatsu-ku, Kawasaki-city, Kanagawa, 213-0012, Japan
ABSTRACT The activity of carbon-supported catalysts was investigated for a slurry phase system. In this paper, some active carbons were used as carbon-supports. We prepared NiMo/active carbon and Fe/active carbon catalysts. These catalysts were used for hydrogenation of heavy oil at 693K. At the same time, we examined 1methylnaphthalene (1-MN) hydrogenation activity and methylcyclohexane (MCH) dehydrogenation activity for our catalysts and supports. We discussed the correlation of asphaltene conversion with 1-MN and MCH reactivity. There was little correlation between them. It is found that active carbons contributed to decrease the coke formation. 1. INTRODUCTION To meet an increasing demand for conversion of heavy oils with high efficiency to transportation fuels, we are studying high-performance catalysts. And now, we are investigating carbon-supported catalysts for a slurry phase system. Properties of carbon supports, such as surface area, porosity and functional groups are very important for the catalytic activity [1]. For example, HNO3 treatment of carbon supports is known to improve catalytic activity by the introduction of oxygen functional groups to the catalyst surface [2,3 ]. Active carbons have MCH dehydrogenation activity more or less. That means active carbons have the ability to abstract hydrogen from hydrocarbons. This time, free radicals are formed on active carbons. The free radical should initiate the cracking reaction [4]. Developing active carbon catalysts, we have to pay attention to the ability of hydrogen transfer. In this work, we discussed the correlation of asphaltene conversion with 1-MN and MCH reactivity. Furthermore, we explained the coke formation in active carbon catalysts under sufficient hydrogen atmospheres.
390 2.EXPERIMENTAL
Arabian-heavy vacuum residue (AH-VR) was used as heavy oil. The properties of AH-VR are given in Table 1. Reactions were conducted in 140ml-batch reactors. AHVR of 10.0g was charged into the reactor. Reaction temperature, time and initial hydrogen pressure were 693 K, 2h and 10MPa (cold charge), respectively.
Table 1 Properties of AH-VR VR" Asphaltene wt% wt% 90.4 16.0 a bp.>793 K
Sulfur wt% 5.32
Nitrogen wt% 0.46
Vanadium wtppm 171
Nickel wtppm 61
We used five kinds of active carbons (A, B, C, D and E). The characteristics of active carbons are given in Table 2. Table 2 Characteristics of active carbons A B SBETa mZ/g 830 1500 PV b ml/g 0.38 0.76 Dc nm 10 10 Ash wt% 6.0 4.4 aBET surface area bpore volume
C D 1540 900 0.78 0.40 3 9 0.8 8.1 CAveragepore
E 1100 0.48 1 2.3 diameter
The catalysts were prepared by an incipient wetness impregnation method to obtain NiMo and Fe loadings of 10wt%. NiMo catalysts and Fe catalysts were added to AHVR to be 1000 and 8000ppm on metal base, respectively. Ni/(Ni+Mo) was 0.30 in molar ratio. After reactions, we measured VR conversion, asphaltene conversion, coke yield and activities of hydrodesulfurization (HDS) and hydrodenitrogenation (HDN). And we checked the basic characteristics of our catalysts and supports. For catalysts, 1-MN hydrogenation activity was examined. These reactions were conducted in 30mlbatch reactor at 613K. For carbon-supports, we measured MCH dehydrogenation activity. Pulse reactors were used for these reactions at 773K under nitrogen.
3.RESULTS AND DISCUSSION
MCH dehydrogenation activity for active carbons and 1-MN hydrogenation activity for NiMo/active carbon catalysts were given in Table 3 and 4. The results of AH-VR hydroconversion were shown in Table 5 and 6.
391
Table 3 MCH dehydrogenation activity for active carbons Active carbon A B C MCH conv. % 23 69 31
D 45
E 46
Table 4 1-MN hydrogenation activity for NiMo/active carbon catalysts NiMo/active carbon A B C D E 1-MN conv. % 18 75 37 46 49 Table 5 Hydroconversion of AH-VR by NiMo/active carbon Active carbon A B C VR conv. % 61 60 57 Asphaltene conv. % 44 34 HDS % 37 57 54 HDN % 0 5 6 Gas C1-C5 wt% 5.5 5.7 5.3 Coke wt% 3.3 3.0 0.2 Table 6 Hydroconversion of AH-VR by Fe/active carbon Active carbon A B C VR conv. % 56 50 53 Asphaltene conv. % 86 89 78 HDS % 67 48 49 HDN % 32 27 21 Gas C1-C5 wt% 5.4 5.7 5.2 Coke wt% 1.4 1.7 1.4 ano metal (only active carbon)
D 69 48 45 22 5.5 3.9
E 69
D 59 50 51 14 6.1 0.8
E 62 42 47 17 7.0 0.3
44 13 6.8 3.9
no-cat 66 30 15 0 7.3 10.7
A" 66 21 17 0 7.4 6.4
MCH dehydrogenation activity of active carbons is increased in the order of B, E, D, C and A from Table 3. 1-MN hydrogenation activity of NiMo/active carbon catalysts is increased in the same order from Table 4. If the metals on supports have sufficient hydrogenation activity under enough hydrogen atmospheres, the active carbon which has higher MCH dehydrogenation activity will give better hydrogenation catalysts. From Table 5 and 6, VR conversion is always about 60%. VR conversion would be dominated by the reaction temperature. However, we can find the influence of hydrogenation activity of the catalysts. Asphaltene conversion will be dominated by the catalytic hydrogenation activity and the surface properties of a carbon-support. Fe/A and Fe/B show excellent asphaltene conversion (86-89%) and small coke yield. In these catalysts, VR conversion is 50-56%. On the other hand, VR conversion is over 60% when the catalysts show low asphaltene conversion. It is due to the coke formation that VR conversion seems high.
392
Higher 1-MN hydrogenation activity catalysts seem to show excellent asphaltene conversion. For NiMo/A and NiMo/B in Table 5, 1-MN hydrogenation activity and asphaltene conversion do not always have correlation. The surface properties of catalysts will be more important than 1-MN hydrogenation activity. From Table 6, MCH dehydrogenation activity and asphaltene conversion do not always have correlation, too. The reason why Fe/A and Fe/B show excellent asphaltene conversion is probably that their average pore diameter is large enough to react with asphaltene. From Table 6, it is clear that the addition of the active carbon A reduced the coke from 10.7 to 6.4wt%. It is probably because the active carbon A works as a vehicle for hydrogen transfer to coke precursor under sufficient hydrogen atmospheres. Furthermore, we observed that the loading of Fe on the active carbon A fairly contributed to decrease the coke formation. 4.CONCLUSION The main findings of this work can be summarized as follows: (1) MCH dehydrogenation activity and 1-MN hydrogenation activity have good correlation. High MCH dehydrogenation activity of active carbons is very important to develop better hydrogenation catalysts. (2) VR conversion would be dominated by the reaction temperature. It is due to the coke formation that VR conversion seems high. (3) 1-MN hydrogenation activity, MCH dehydrogenation activity and asphaltene conversion do not always have correlation. For high asphaltene conversion, not only hydrogenation ability but also the surface properties of catalysts will be important. (4) The addition of active carbon reduces the coke. Active carbon works as a vehicle for hydrogen transfer to coke precursor under sufficient hydrogen atmospheres.
5.ACKOWLEDGEMENT
This work has been carried out as a research project of the Petroleum Energy Center and subsidized by the Ministry of International Trade and Industry.
6.REFERENCES
1 F.Rodriguez-Reinoso, Carbon 36, No.3 (1998) 159 2 J.M.Solar et al, J.Catal. 129 (1991) 330 3 S.Wang and G.Q.Lu, Carbon 36, No.3 (1998) 283 4 I.Nakamura and K.Fujimoto, Sekiyu Gakkaishi 39, No.3 (1996) 245
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) o 1999ElsevierScienceB.V. All rightsreserved.
Effects of Gaseous and Liquid Components on Rate Desulfurization of Heavy Atmospheric G a s Oil.
393
of Deep
M . V . L a n d a u a, L . V r a d m a n a , M . H e r s k o w i t z a a n d D.Yitzhakib aBlechner Center for Industrial Catalysis and Process Development, Chemical Engineering Department, Ben-Gurion University of the Negev, P.O.Box 653, Beer-Sheva 84105, Israel boil Refineries Ltd., P.O.Box 4, Haifa 31000, Israel Abstract The effects of concentrations of sulfur, nitrogen, bi-(BA) and monoaromatics (MA) in heavy atmospheric gas oiI(HAGO), H2S and ammonia in gas phase on HDS rate at deep desulfuriation stage (Sin 1110-60 ppm) were studied with Co-Mo-Al and Ni-W-Si catalysts using HAGO with FBP of 390~ and initial sulfur content of 1.24 wt.%. The complete elimination of hydrogen sulfide, ammonia, polyaromatics and partial elimination of monoaromatics prior to the deep desulfurization stage increases the overall rate of deep HDS by a factor of about six. 1. I N T R O D U C T I O N Hydrodesulfurization processes were designed to lower sulfur in diesel fuels to 2000 p p m mad than to 500 ppm. Significant modifications of catalysts and process design are required to meet the new 50 ppm standards. This is a specifically difficult task for HAGO feedstocks with FBP > 370~ Estimation [1] based on desulfurization kinetics of HAGO showed that at current operating conditions, the activity of desulfurization catalysts should be increased 3-5 times. This target could be substantially decreased by cleaning the liquid and gas phases before the deep desulfurization stage from nitrogen, aromatic compounds, hydrogen sulfide and ammonia. These components inhibit desulfurization of HAGO at high sulfur conversions (>90%) [2-4]. The scope of this work was to measure effects of inhibitors and determine their impact on the desulfurization rate in the deep range (60-1100 p p m sulfur and 5-100 p p m nitrogen) using a real feedstock r a t h e r than model compounds. 2.
EXPERIMENTAL
The HDS experiments were carried out in an automatic minipilot trickle-bed system. The presulfided commercial Co-Mo-A1 and a proprietary Ni-W-Si catalysts (covering a wide range of desulfurization and hydrogenation activity measured with model compounds dissolved in paraffinic solvent (Table1)), were tested in hydrotreating of HAGO feedstocks. All tests were carried out at standard conditions: total pressure 5.4 MPa, temperature 360oQ H2/oil ratio = 500 NL/L, LHSV of 0.5-20 h-1. Fresh HAGO and hydrotreated HAGO were tested. The biaromatics and nitrogen concentration in HAGO w e r e controlled in addition to hydrotreating by addition of naphthalene and
394
Table 1 Catalysts characteristics (o:dde form) Catalyst *
* 9 9 9 9 9
Co-Mo-A1
Chemical c o m p o s i t i o n , wt.% Co Ni Mo W Support Surface area, m2/g Pore v o l u m e , cm3/g Bulk density, g / c m 3 Average pellets diameter, m m P s e u d o - f i r s t - o r d e r rate c o n s t a n t s , h-l: DBT ttDS 4,6-diMe-DBT HDS Toluene h y d r o g e n a t i o n
Ni-W-Si
6.3 18.4 A1203 280 0.49 0.75 1.3
16.6 26.8 SiO2 180 0.37 1.00 1.7
486 272 0.6
362 422 5.8
quinoline. The c h a r a c t e r i s t i c s of f l e s h a n d nine h y d r o t r e a t e d HAGO used in these tests are shown in Table 2. Model c o m p o u n d s (DBT and 4,6-diMe-DBT) and p r o d u c t s of their h y d r o c o n v e r s i o n were m e a s u r e d by GC a n d GC-MS. The gas oil c o m p o s i t i o n was d e t e r m i n e d by GC-MS a n d HPLC (aromatic c o m p o u n d s ) . Organic sulfur a n d n i t r o g e n in HAGO were a n a l y z e d by ASTM 13-4045 and ASTM D-4629 m e t h o d s , respectively, a n d the c o m p o s i t i o n of effluent gas ( h y d r o g e n sulfide a n d a m m o n i a ) - by on-line GC. Table 2 Compositions of HAGO feedstocks Feedstock Fresh HAGO HT HAGO-1 I-IT HAC,(>2 HT HACK3-3 HT HACX>4 HT HAC,O-5 HT HACK)-6 HT HAGO-7 HT HAG(N8 HT HACK)-9
3. RESULTS AND
Sulfur content, ppm 12400 5000 2000 1100 400 200 176 118 115 60
Nitrogen content, ppm 302 250 160 100 30 10 <10 <10 10 <10
Aromatics, wt.%: monobi20 20 20 20 21 23 27 20 23.5 11
7.5 7.5 7.5 7.3 5.5 4.0 2.8 -
DISCUSSION
By m o v i n g the feedstock along the catalysts l a y e r t o g e t h e r with lowering the sulfur c o n t e n t occurs a decrease of c o n c e n t r a t i o n s of n i t r o g e n a n d
395
biaromatics, increase the concentration of m o n o a r o m a t i c s in liquid and increase the c o n c e n t r a t i o n of hydrogen sulfide and a m m o n i a in the gas phase. This is illustrated in Fig.1 for HDS of flesh HAGO at standard conditions and IMSV = 3 h "1 with commercial Co-Mo-A1 catalyst and in Table 2. These results were obtained by m e a s u r e m e n t s the compositions of liquid and gas products after h y d r o t r e a t i n g at different LHSV. Close picture was m e a s u r e d with Ni-WSi catalyst that displayed higher HDN and aromatics h y d r o g e n a t i o n efficiency. 10000o ~
.
.
.
.
.
.
.
.
.
.
.
1.8
12
10000
1.0
~
..~
o6 } 04
~
~
.__~-
100 m
0
.
20
40
.
.
.
-
60
Dimensionless reactor length
80
~ O0
100
%
Figure 1. Distribution of sulfur in HAGO and H2S/NH3 in gas phase catalysts layer
along the
The desulfurization rates were calculated by fitting integral m e a s u r e m e n t s of sulfur c o n c e n t r a t i o n as a function of (LHSV)-Z to a power low expression. As expected, the reaction o r d e r on sulfur was 1.65 for flesh HAGO. It decreased gradualy to unity at sulfur concentrations below 0.11 wt.%. The concentrations of H2S and NH3 already reach close-to-maximal value at the deep desulfurization point that could be defined as ~ 1000 p p m sulfur content (Fig.l). A gradual decline in the desulfurization rate was r e c o r d e d with increasing sulfur conversion (decreasing the initial sulfur concentration in HAGO) as shown in Fig.2. It implies that the reactivity of the l u m p e d sulfur components in the liquid phase d e c r e a s e d with increasing conversion. Carrying out the experiments in two stages at the same conditions, with the same catalyst and HAGO sulfur concentrations, but using pure h y d r o g e n at the second stage, increased the HDS rate - m o r e at higher sulfur conversions (Fig.2). That is a result of excluding the inhibition of HDS by H2S and a m m o n i a in gas phase. In o r d e r to separate the effects of H2S and a m m o n i a the gas oil desulfurized to residual sulfur content of 0.11wt.%(HT HAGO-3) was mixed with dimethyldisulfide to get the same H2S content in the gas phase as that present at the same sulfur conversion with fresh HAGO. Comparison of the pseudofirst-order rate constants m e a s u r e d at the same initial sulfur concentration (0.11 wt.%) working in one stage with fresh HAGO and working with HT HAGO3 (two stages) with and without DMDS addition is shown in Fig.3. Separation of both poisons before deep desulfurization stage increases the HDS rate by a factor of 2.75 while addition of DMDS decreased the HDS rate only by a factor of 2.3. It means that contribution of H2S to the overall decrease of HDS rate at deep desulfurization stage comprise 85% while contribution of a m m o n i a - only 15%. The nitrogen concentration in HAGO at deep HDS stage became <100 ppm Table 2). Decreasing of its concentration to 10 p p m (Fig.l, Table 2) or addition of 10 p p m nitrogen in form of quinoline to HT HAGO-3 did not change the HAGO HDS rates. It means that the impact of residual nitrogen in the liqiud feedstock at deep HDS stage is negligible. No significant differences in the
396
I--"- ,,.o.,,.g. ,~ I
"o
i,.
I0
1
~0.1 m
0.1
0.01
. . . . . . . . . . 100
~
9
0.01
' ..... - . . . .
1000 S ~ . ppm
104
.
0
,
-
.
.
,
0.1
. . . . . .
0.2
,
.
.
.
.
0.3 Sl:~ce time. hOurS
j
.
.
0.4
.
,
.
.
0.5
.
0.S
Figure 3. Effects of H2S/NH3 separation and DMDS addition on HDS of HT HAGO-3
Figure 2. Effect of sulfur conversion on HAGO HDS rate
effects of HzS and a m m o n i a on Co-Mo-A1 and Ni-W-Si catalysts were detected. The effects of bi- and mono-aromatics contents in HAGO on the pseudo-firsto r d e r HDS rate constants at deep HDS stage are shown in Figs.4 and 5. Ni-W-Si
10
~
Open flgure~, naphthalene addition to HT HAGO [ Solid figures: real aromatics in HT HAGO [ I S in = 1 1 5 - 2 0 0 ppm I
MA = 2o-a7% N-, ~0 ~,~,m I
CI A
"
Co..Mo..AI | J Ni--W-SI I I
"'~='
II
] J I I
Sin = 6 0 - 1 7 6
22
ppm;
[
"
N~'w'si'I
10ppm
I
=
Co-Mo.-AI I
t. . . . .
J
20 "
8
,y, 7
5
t..... 0
,
1
.....
,'..-
,
2
3
...1
....
~-
4. 5 Biaromatics in the f e e d . w t . %
-.,
..... 6
Figure 4. Effect of biaromatics on HAGO HDS rate constants
~.. 7
t
10
~
5
9 10
9' -
1
-
-
-
i . . . . . 15
-
.
-
.
, 20
. . . .
Monoaromatics in the feed, w L %
i
. . . .
25
Figure 5. Effect of m o n o a r o m a t i c s on HAGO HDS rate constants
display higher HDS activity, and in this case increasing the b i a r o m a t i c s content in HAGO decreases the HDS rate about linearly. Co-Mo-A1 catalyst is less sensitive to biaromatics at its high concentrations that explains no effect of biaromatics on gas oil HDS rate detected in [3] by addition of polyaromatics to partially d e a r o m a t i z e d feedstock. Co-Mo-A1 catalyst is also m u c h less sensitive to m o n o a r o m a t i c s (Fig.5), while with Ni-W-Si catalyst the HAGO HDS r a t e increased by a factor of 2.5 after removing about 50% of m o n o a r o m a t i c s existed in fresh HAGO. Regulation the composition of liquid and gas phase prior the deep desulfurization stage is a viable option for i m p r o v e m e n t the HDS p e r f o r m a n c e .
4 REFERENCES i. M.V.Landau, Catal.Today, 36(1997) 393. 2. E.Lecrenay and l.Mochida, Stud.Surf.Sci.Catal., 106(1997) 333. 3. F.van Looij, P.van der Laan, W.H.J.Stork, D.J.DiCamillo and J.Swain,Appl.Catal, 170 (1998) I. 4. LVradman, M.V.Ixndau and M.Herskowitz, Catal.Today, 48 (I 999) 41.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Fromentand P. Grange (Editors) 91999 Elsevier Science B.V. All rights reserved.
397
Catalytic functionalities of TiO2 based SiO2, A1203, ZrO2 mixed oxide hydroprocessing catalysts M. S. Rana, B. N. Srinivas, S. K. Maity, G. Murali Dhar and T. S. R. Prasada Rao Catalysis Division, Indian Institute of Petroleum, Dehradun-248 005, INDIA ABSTRACT Thiophene HDS and cyclohexene HYD activities on three mixed oxides supported Mo and promoted by Co or Ni are presented. A comparative study indicated that Ti containing mixed oxide supports show outstanding activity for HDS. The trends of variation of activities indicated that the two functionalities originate from different set of catalytic sites on molybdenum sulfide phase. 1.
INTRODUCTION
Sulfided molybdenum and tungsten based catalysts supported o n A1203 are well known in hydroprocessing of petroleum fractions [ 1-4]. Various types of reactions catalysed by these materials are hydrogenation [HYD], hydrogenalysis [HDS] and hydrocracking. These functionalities are not only required for an effective catalyst but also should be appropriately balanced. To be able to prepare effective well balanced catalysts it is necessary to gain knowledge on the origin of different functions of the catalysts and its dependence on nature of the support, composition, pretreatment etc. variables. It is well known that support plays an important role in hydroprocessing reactions. Support effects have been studied by many authors including our group. We have studied hydrotreating reactions on molybdenum supported on A1203 [5], SIO2-A1203 [6], TiO2-A1203 [7], sepiolite [8], carbon [9] and W supported on SIO2-A1203 [10], ZrO2 [11], TiO2-A1203 [12] etc. systems. TiO2 based systems have been acknowledged to be very active supports for HDS and related reactions. However because of the difficulty in preparing the support and catalyst with high surface area and its retention at high temperatures, coupled with unsuitable mechanical properties, stood in the way of advancing their application further. In order to overcome these limitations TiO2 based mixed oxide supports are being tried in recent years [ 13,14,15]. In this investigation a comparative study of catalytic functionalities of molybdenum supported on various TiO2, ZrO2, SiO2, A1203 based mixed oxides will be presented. For want of space, characterization details will not be given here in this report, and the emphasis is on a comparative analysis of catalytic functions of catalysts prepared by using mixed oxides, single oxides of the elements present in the mixed oxides, as supports for promoted and unprompted molybdenum active component. 2.
EXPERIMENTAL
The mixed oxide supports of varying compositions were prepared by homogeneous precipitation technique from dilute solutions containing respective ions using urea hydrolysis at 90~ for 6h and calcined at 550~ for 5h. [8,12]. The impregnation of the active component Mo and the promoters Co and Ni were carried out by incipient wetness method. The catalysts as well as supports were characterized for surface area by BET method, X-ray diffraction, volumetric oxygen chemisorption at -77~ Thiophene HDS and cyclohexene HYD activities were evaluated at 400~ on a catalyst sulfided in CS2-hydrogen mixture in a fixed bed reactor operating at atmospheric pressure and interfaced with a gas chromatograph using a six way sampling valve for on line product analysis. Details about support, catalyst preparation, oxygen chemisorption and activity studies were given else where [4,11].
398
3.
R E S U L T S AND D I S C U S S I O N
In this investigation the results of a comparative study of catalytic functionalities hydrogenalysis of thiophene and hydrogenation of cyclohexene on molybdenum supported on three mixed oxides based on TiO2 viz., TiO2-ZrO2 (65/35), TiO2-SiO2 (12/88) TiO2-A1203 (50/50) are presented. In order to better appreciate the support effect, studies on Mo supported on SiO2, TiO2, A1203 and ZrO2 is also included. The same reactions on Co, Ni promoted analogs on all the supports is also discussed, in order to understand the role of the promoter and origin of the functionalities. The over all objective is to understand the role of support in altering the two catalytic fucntionalities, hydrogenalysis and hydrogenation, and also role played by Ti ion in mixed oxides in modifying these functionalities. All the catalysts studied in this investigation contained 8wt.% molybdenum except in the case of TiO2-ZrO2 where 12wt.% Mo is used. The reason for this being, variation of molybdenum loading on these supports indicated that the maximum in activity is obtained at this wt.% of molybdenum loading. The supports are of good enough surface area to spread molybdenum as a monolayer at the Mo loading studied and XRD and oxygen chemisorption measurements indicated that molybdenum is well dispersed on all the samples presented here. First of all a comparison of the catalytic activities for the two functionalities on Mo/TiO2-ZrO2, Mo/TiO2-SiO2, Mo/TiO2-A1203 as a function of Ti/M ratio where M, is Zr, or Si, or A1 will be discussed. The relevant data on catalytic activities is presented in Fig.1. It can be seen that in SiO2-TiO2 system, HDS activity is maximum at 12 mol% TiO2 in the support. Hydrogenation activity showed only marginal difference as a function of TiO2 content. It is interesting to note that at lower TiO2 content HDS/HYD ratio favour of HDS and gradually shifts towards hydrogenation at higher TiO2 contents, showing clearly that support influences the functionalities of Mo independently. Let us discuss the variation TiO2 content in TiO2-A1203 system. Among the three compositions of the support studied, TiO2-A1203 (50:50) molar ratio showed highest activity for both the functionalities. At higher A1203 content in the support, HDS is higher than hydrogenation and at higher TiO2 contents the reverse namely higher hydrogenation is observed indicating that the support is influencing the catalytic sites in an independent way. In the case of TiOz-ZrO2 where four TZ compositions were studied, it can be seen that the maximum in activity is obtained at 65% TiO2 for both HDS and hydrogenation. HDS undergoes profound changes as a function of TiO2 content, compared hydrogenation which shows less pronounced differences. Among the three best support compositions viz. Mo/TiOz-ZrO2 (65:35), Mo/TiOz-A1203 (50:50) and Mo/TiO2-SiO2 (12/88) the highest activity for HDS is obtained on TiO2-ZrO2 supported catalyst. Highest activity for hydrogenation is obtained on Mo/TiOz-AI203 supported catalysts. The order of decreasing activity for HDS is Mo/TiOz-ZrO2, Mo/TiOz-AI203, Mo/TiO2-SiO2. The trend of variation of activity for HYD is Mo/TiO2-AlzO3>Mo/TiOz-ZrOz>Mo/TiO2-SiO2. The activity data on ZrO2, TiO2, A1203, SiO2 for the two functionalities are shown in the same figure. It can be seen that in the case of Mo/TiOz-AI203 and Mo/TiOz-ZrO2 the activities obtained are considerably higher than constituent single oxides indicating that a synergistic action of two constituents of the support induced in the supported phase an activity increase, probably due to metal support interaction. It is interesting that in TiO2-SiO2 the similar synergy is observed in the case of HDS but however hydrogenation on mixed oxide supported catalysts is lower than both the Mo catalysts derived from SiO2 or TiO2. This also another instances where HDS and hydrogenation are independently varied via support effect. Promotional effect of Co and Ni on 8wt.% molybdenum supported on ~/ -A1203, TiO2, SiO2, SiOz-TiO2 (12/88), AlzO3-TiO2 (50:50) for thiophene hydrodesulfurization and cyclohexene. From the data shown in Fig. 2 and Table-1 can be seen that the promotional effect indeed is present in all the systems studied. However the extent of promotion is specific to each support and the functionality. There are striking differences in promotional effects for each functionality, for example SiO2 supported Mo exhibits huge promotional effect for hydrogenation where as hydrodesulfurization is only increased to a small extent. In the case of A1203 support cobalt exhibited large promotional effect on HDS. As is well documented in literature on 7-A1203 supported catalysts, Co and Ni show similar promotional effect on
399
the mixed oxide systems studied in this investigation. It is interesting to note that there is large promotional effect for both for HDS and hydrogenation in the case of TiO2-SiO2 supported molybdenum system. However on Mo/TiOz-AI203 the promotional effect is not large in the case of both Co and Ni. A general observation is that on pure TiO2 and TiO2-A1203 supported systems the promotional effects for both the functionalities is only small, in agreement with the results reported earlier on Mo/TiO2 [ 16]. 20
x
15
o
10
,-
5
0
0
E
DS
EIHDS
x 40 ~" t~
A
9
O'0
TS (24/76)
TS TS (38/62)(64/36)
0
x
r-]~
12~
A
*" 100 o
i
i
,
TA TA TA (75/25) (50/50) (25/75)
&
60
$
I HYD
o 40
~.
v 50 e" o
o TZ(85/15) TZ (65/35) TZ(50/50) rz(25/75)
'
[3 HDS
0 .
3302
D
m HYD
30
9 10
rtl
8% Mo loading
x 50
C
~
AI203
m.
g
B
8%Mo/Ti:AI
~. 2o TS (12/88)
m
50
9HYD 8%Mo/Ti:Si
. S~O2
.
.
TS(12/88)
. TIO2
. TA (50/50)
t
I
.
AI203-
TZ (65/35)
ZrO2
Fig. 1 (A, B, C, D) Support composition vs catalytic activity
Table-1. Promotional effects on different supports. Supported Catalysts
....
S.A. (m2g-')
# Catalysts Activity
HDS
HYD
3%Co18%Mo Promoted I-IDS HYD
8% M o
% Increase
3%Ni 8%Mo Promoted I-IDS HYD
HDS
HYD
HDS
I-IYD
3%Co
3%Ni
*SiO2 TS (12/88) TA (50/50)
340 312 129
5.6 11.39 26
17107 8.19 33.1
11.8 38.4 33
60.0 32.2 44
10.63 37.42 38.4
69.7 39.4 40.2
110.7 237.1 26.9
251.7 239 32.9
89.8 228.5 49.2
308.2 381 21.45
A1203 TiO2 TZ(65/35)
153 59
11.2 6.4 50.6
25.4 12.4 11.8
26.8 8.9
24.7 22
19.1 13
29.1 21.7
139.2 39
77.4
70.5 103.1
14.5 75
* 4 % M o loading, # Rate (mol h -1 g-1 cat.) x 103
400
=
80
r-IHOS
IIHYD
~ 60 I
~:~ 50 .E: 40 ~9 30 20
o
DHOS IIHYD
F
]
~rJ 40 ~ 30 "~ 20 10 3~
t( S )
3~
~( T S )
3~
~( T )
3~
i (T A )
3~
~( A )
3 %C o ( S )
3~ o (TS)
3 % 0 o (T )
3~ o (TA)
3 %C o ( A )
Fig. 2 (E, F). Effect of promoters on catalytic functionalities 4.
CONCLUSIONS
These studies on variation of TiO2 based mixed oxides indicated that introduction of Ti into all the supports studied, causes beneficial affect on thiophene hydrogenolysis activity while hydrogenation did not show such outstanding increase. The promotional affect of Co or Ni on Mo supported on Ti containing mixed oxide supports is lower expect in the case of TiO2:SiO2 supported system. The activity variation as a function of the support and promoters indicated that thiophene HDS and cyclohexene HYD originate from different set of sites.
REFERENCES
.
10. 11. 12. 13. 14. 15. 16.
H. Topsoe, B. S. Clausen and F.E. Massoth in "Catalysis Science and Technology" (J. R. Anderson and M. Boudart, eds.), 11, Spring-Verlag, New York, (1996) D. D. Whitehurst, T. Isoda and I. Mochida, Advance in Catalysis, 42, (1998) 345. F. E. Massoth and G. Murali Dhar, Proceedings, Climax of Molybdenum (H. F. Barry and P. C. H. Mitchell eds.), Climax Molybdenum Co., Ann Arbor, Michigan, (1982) p.343. G. Murali Dhar, H. Ramakrishna and T. S. R. Prasada Rao, Catal. Lett. 22, (1993) 351. G. Murali Dhar, F. E. Massoth and J. Shabtai, J. of Catal. 85, (1994) 44. F. E. Massoth, G. Murali Dhar and J. Shabtai, J. of Catal. 85 (1994) 53. B. N. Srinivas, S. K. Maity, M. S. Rana, G. Murali Dhar and T. S. R. Prasada Rao, Presented in IPCAT 1, Cape Town, South Africa, 26-28 Janurary (1998). S. K. Maity, B. N. Srinivas, V. V. D. N. Prasad, Anand Singh G. Murali Dhar and T. S. R. Prasada Rao, Studies in Surface Science and Catal. 113, (1998) 579. K. V. R. Chary, H. Ramakrishna and G. Murali Dhar, J. of Mol. Catal. 68, (1991) L25. K. S. P. Rao and G. Murali Dhar, J. of Catal., 115, (1989) 277. K. S. P. Rao, H. Ramakrishna and G. Murali Dhar, J. of Catal., 133, (1989) 146. B. N. Srinivas, S. K. Maity,, V. V. D. N. Prasad, M. S. Rana, Manoj Kumar, G. Murali Dhar and T. S. R. Prasada Rao, Studies in Surface Science and Catal. 113, (1998) 497. K. Segawa, M. Katsuta and F. Kameda, Catal Today, 29, (1996) 215. Zhaobin, W., Qin, X., Xiexian, G., Sham, E. L., Grange, P., Delmon, B. Appl. Catal., 63, (1990) 305. Zhaobin, W., Qin, X., Xiexian, G., Sham, E. L., Grange, P., Delmon, B. 1991Appl. Catal., 75, (1991) 179. M. Breysse, J. L. Portefalx and M. Virnat, Catal. Today, 10, (1991) 489
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
401
H y d r o d e s u l f u r i z a t i o n o f dibenzothiophene over N i - M o / ( P ) T i - H M S catalysts T. Halachev, J.A. de los Reyes, C. Araujo, L. Dimitrov 1 and G.Cordoba Universidad A. Metropolitana-Iztapalapa, Apdo. 55534 M6xico, D.F. phone 52(5)724-4648 fax 52 (5) 724-4900 email [email protected] lInstitute of Catalysis of the Bulgarian Academy of Sciences, Sofia-1113, Bulgaria
Abstract
NiMo hydrotreating catalysts supported on the newly synthesized mesoporous material TiHMS have been tested in the reaction of HDS of DBT and proved to be superior to the conventional alumina supported catalysts. A discussion on the relation between the catalytic activity measurements and the physicochemical characterization of the samples is presented. INTRODUCTION The ever increasing environmental restrictions imposed on sulfur and nitrogen contents_in refined oil fractions and the growing use of heavier crudes present a difficult and challenging task for the fundamental research in catalysis and the application of its results under real industrial conditions. One of the possible ways to face these problems is the use of new mesoporous supports having large surface area and narrow pore size distribution, which allows to treat heavy oil fractions. In the present communication we report results on the catalytic activity of NiMo/(P)Ti-HMS catalysts in the hydrodesulfurization (HDS) of dibenzothiophene (DBT), using the recently synthesized mesoporous material Ti-HMS [1,2] as a support. The effect of the phosphorus content on the catalytic activity is also studied. EXPERIMENTAL Catalysts preparation The Ti-HMS support (SBET-"800 m2/g, pore volume 1.7 cm3/g, average pore diameter 85A) was modified by the addition of H3PO4 in order to obtain samples with 0, 0.5, 1.0 and 1.5 wt.% P205. The catalysts were prepared by the incipient wetness method (Mo being introduced first) and contain 4.6 wt % Ni and 10.6 wt % Mo. The solids were dried at 120~ overnight and then calcined at 400~ for 2 h. The samples are denominated as sample 1; 2; 3 and 4 according to the increasing phosphorous concentration. An additional sample 5 ,of an industrial catalyst (CDS-R2, Catalysts & Chemical IND-Co.LTD), was included for comparison in the catalytic activity measurements. Catalytic activity measurements Prior to catalytic activity measurements, the catalysts were activated for 1 h in a H2]H2S (10%) flow at 1.1 ml/min at 400~ The activated catalysts were transferred in an inert atmosphere to a batch reactor (Nikko Koatsu) containing DBT and hexadecane as a solvent. The reactor was charged with H2 to a pressure of 70 atm and the reaction was carried out for 8 hrs at 300~ and vigorous stirring (1000 rpm). The reaction products were analyzed at each
402 hour by GC. First order kinetics was used to evaluate the reactions constants and the initial reaction rate. Catalysts characterization The supports and the catalysts were characterized by nitrogen adsorption size and total pore volume) XRD, Raman, DRS and TPR.
(SBET,average pore
RESULTS AND DISCUSSION Fig 1 shows the observed degrees of conversion of DBT on the catalysts. One can note that the catalysts with 1.5% wt P205 manifests a very high degree of conversion- 96%, which is obtained at a relatively low reaction temperature. Besides, the catalysts have a good hydrogenation activity as can be observed in Fig.2. The catalytic activity and selectivity curves indicate that the effect of phosphorous is not linear. The phosphorous free sample has higher activity than the low phosphorous content samples. However, the highest activity and selectivity is manifested by the sample with the highest phosphorous concentration. Besides, among the phosphorous containing samples the HDS and HYD activities increase with the increase in phosphorous content. The Ti-HMS catalysts have higher activity than the conventional one.
100
1 | , Sample 1 0.8 ~ + Sample 2 / ._ Sample 3 0.6 -_
80
60 r/)
g
_04
40_
o
~
20_ r
01 0
~ Sample 4 o Sample5
. . . . 120 -",--,m'2~40tmin)360
480
Figure 1. Conversion of DBT vs. time.
0.2 0 , 0
_
~
~
, ........
120
240 360 480 Time (min) Figure 2.Selecfwity to cyclohexylbenzene vs. time.
Similar effect of phosphorous was observed for catalysts supported on alumina [3]. This effect is rather interesting and needs additional studies. Besides, it is not clear yet up to what concentration the effect of phosphorous would be favorable. One could suggest that with the increase in phosphorous concentration the surface acidity of the Ti-HMS support would increase, hence the interaction between the support and the active phases would change, thus giving rise to different surface species. Indeed, the study of the surface acidity of the phosphorous containing supports [4] by the reaction of dehydration/dehydrogenation of 2propanol indicates that it increases vs. phosphorous content. However, quite unexpectedly, the supports with 1.0 wt.% and 1.5 wt.% of P205 manifested the presence of basic sites. The generation of basic sites can be explained with the structural changes caused by the presence
403
800 700 600 500 _ 400
oo
300 200
-
100
-
-!
0 1
i
i
i
3
5
7
9
20
Figure 3. X-ray patterns of the supports. of phosphorous, evidenced by the increasing loss of "crystallization" as revealed by the XRD patterns (Fig.3). Only the peak typical for the support [1 ] is observed. Fig.3 also contains the X-ray pattern of an additional sample containing 2wt.% P20 5, which we included to show more expressively the changes (amorphization) in the support in the presence of increasing amounts of phosphorous 9 These changes are also evidenced by the decrease of the surface area by about 30 mZ/g and by the gradual decrease of the total pore volume and average pore diameter to 1 9 cm3/g and to 76A, correspondingly. However, all these changes do not coincide with the trends observed in the catalytic activity and selectivity curves. We should also note that no X-ray detectable phases of Ni and Mo were observed, which indicates that highly dispersed catalyst have been prepared. The Raman spectra (not shown) of the treated supports and of the catalysts did not contain the peak at about 134 cm -~ typical for crystalline titania particles. The possible changes in the local surrounding of the Ni and Mo ions were studied by DRS (Fig.4). The spectra of all samples look alike and reveal the presence of Mo6+(Oh) ions at 334nm and of Niz+(Oh) ions at 445nm (shoulder). No changes in the positions of the peaks vs. phosphorous content are observed. Besides, contrary to similar samples supported on alumina, no N1 (Td) aons were registered. 9 2+
9
1600 . . . . . . . /'h /ff~ 0.8 , ///-,,\\ ..~ 0.6 ~
0.4
<
0.2
r~
0
1. Sample 1 2. Sample 2 3. Sample 3
. . . . .
~. Sample ~
2. Sample 2 3. Sample 3
~,~ 1200_ 800 400
I
I
200 300 400 500 600 700 800 Wave lenght (nm) Figure 4. DRS spectra of the samples
0
........
20
i
I
................
i .....
l
220 420 620 820 Temperature ~ Fig.5. TPR profdes of the samples.
404
A chemical reaction approach (TPR) proved to be a more sensitive technique for registering different surface species. The TPR profiles of the supports (not shown) practically coincide with the base line of the equipment, showing that no reduction of the pure and phosphorous treated supports takes place. However, the TPR profiles (Fig.5) of the catalysts give a strong evidence for the appearance of new nickel species [5] with the increase of phosphorous concentration, which are reduced at approx. 670~ Also the increasing amount of phosphorous gives rise to a more pronounced shoulder, related to the reduction of nickel species, at approx. 490~ We should also note that the presence of phosphorous leads to a shift to lower temperatures of the reduction peak (at approx. 400~176 of the molybdenum species [5]. This shift indicates that the surface species in the presence of phosphorous are easier to reduce, which could be possibly related to the catalytic activity results. It is difficult to identify unambiguously the surface species present and their direct relation to the catalytic activity and selectivity at this stage of investigations. Probably a number of factors, such as high surface area of the support, large pore volume and pore diameter, surface acidity changes, high dispersion, specific morphology and enhanced reducibility of the supported species are involved in the observed high catalytic activity and promoting effect of phosphorous. CONCLUSIONS The newly synthesized mesoporous material Ti-HMS is a highly effective support for Ni and Mo containing hydrotreatment catalysts, leading to higher activity than the conventional alumina supported catalysts. The addition of phosphorous to the support has a promoting effect on the catalytic activity. No Ni2+(Td) species are observed on the Ti-HMS support. The presence of phosphorous results in better reducibility (sulfurization) of the supported metal species. REFERENCES
1. 2. 3. 4. 5.
P. Tanev, M. Chibwe and T.J. Pinnavaia, Nature 368 (1994) 321. T. Halachev, R. Nava and L. Dimitrov, Appl. Catal. A: General 169 (1998) 111. P. Atanasova and T. Halachev, Appl. Catal., 48 (1989) 295. T. Halachev, T. Viveros, G. Perez and L. Dimitrov, Materials Research Society, in press. P. Atanasova, R. Lopez Cordero, L. Mintchev, T. Halachev and A. L. Agudo, Appl. Catal., A: General 159 (1997) 269.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F. Fromentand P. Grange(Editors) 91999ElsevierScienceB.V. All rightsreserved.
405
The p r e p a r a t i o n o f h y d r o c r a c k i n g catalysts using m e s o p o r o u s a l u m i n o s i l i c a t e s M C M - 4 1 - influence o f the preparation conditions on the catalytic b e h a v i o u r A. Klemt a, A. Taouli a, H. Koch b and W. Reschetilowski a a
Institute for Industrial Chemistry, University of Technology Dresden, D-01062 Dresden*
b Aventis Research & Technologies GmbH & Co KG, D-65926 Frankfurt/Main
This paper reports different methods of modifying A1-MCM-41 materials and the characterization of their acidic properties using TPD of ammonia. The NiMo-impregnated samples were tested in the hydrocracking of n-decane. The results indicate that the kind of preparation method has impact on the catalytic properties. The increase of the amount of strong acidic sites of the NiMo/MCM-41 materials was reached using the proposed pretreatment. This led to the decrease of the hydrogenolysis in the hydrocracking of n-decane.
1. INTRODUCTION The hydrocracking of high boiling hydrocarbons is a flexible way of upgrading the different fractions of crude oil destillation. During the various processes different, catalysts are used to enhance the conversion rate and to influence the selectivity of the process. The accessibility, especially for bulky molecules, can be improved by the use of mesoporous aluminosilicates of the MCM-41 type [1]. The variation of the acidic properties using different pretreatments is important for the use in the cracking processes.
2. EXPERIMENTAL 2.1. Synthesis and modification The MCM-41 material with the Si/A1 ratio-19 was synthesized according to the procedure by Schmidt et. al. [2]. The template removal was carried out using two different methods. Sample A was only calcined (540 ~ 10 1/h air, 12 h) and then ion-exchanged in 0.1 N aqueous NHnNO3 solution. Samples B and C were extracted and ion-exchanged simultaneously [3]. After deammoniation the H-A1-MCM-41 material was impregnated using an aqueous solution of ammonium heptamolybdate and nickel nitrate and calcined after every impregnation. The obtained samples contain 10wt.-% MoO3 and 2.3 wt.-% NiO. The different ways of modifying the parent MCM-41 material are summarized in Table 1.
406
Table 1 Pretreatment and modification methods for MCM-41 materials conventional method template removal to thermal / oxidation Na-MCM-41
...........
ion exchange to NH4-MCM-41
3 x 0.1 N aqueous NH4NO3 solution, 70 ~
transfer in H-MCM-41
thermal
impregnation
optimized method
extraction with ethanol / acid mixture, calcination
MCM-41 material + (NH4)6M07024 and Ni(NO3)2 solution, calcination
2.2. Characterization The determination of the nitrogen adsorption isotherms was carried out at 77 K using a SORPTOMATIC 1900 (FISONS). The TPAD investigations were performed out using a UHV apparatus [4] which enabled also the investigation of the nature of the acidic sites using in-situ FT-IR. The QMS was used to detect the desorbed species during the TPD of ammonia. The hydrocracking of n-decane was carried out in a flow reactor at 400 ~ under 1.5 MPa total pressure (? (H2) - 15 1 h -1, WHSV = 19 h -1 ).
3. RESULTS AND DISCUSSION The results of the physicochemical characterization are shown in Table 2. XRD measurements and the nitrogen adsorption isotherms still showed the typical mesoporous MCM-41 structure. The removal of the template by extraction (samples B and C) reduces the number of silanol groups and the thickness of the pore walls [3]. The extracted H forms of the MCM-41 materials showed a larger specific surface area, higher porosity and greater pore diameter as compared to the calcined sample A. The more acidic extraction media dealuminated the material significantly. The MCM-41 structure remained intact. The pretreatment stabilized the materials with regard to the impregnation using aqueous solutions of metal salts. BET surface area and the pore diameter of all template-free materials were reduced, especially in the case of sample A. The MCM-41 structure of sample A is largely destroyed. The impregnation using the incipient wetness method led to the same result. The amount of ammonia adsorbed, i.e. the sum of the acidity per gram, of extracted and ion-exchanged sample B, was greater than that of the conventionally modified sample A. The specific surface acidity also increased slightly. However, the types of acidic sites and their strenghts were not changed by this pretreatment (Figure 1). In contrast the strong dealumination of the sample C reduced the specific surface acidity as well as the amount of weak acidic sites. The impregnation changed the surface acidity significantly. The amount of strong acidic sites was reduced. The extracted sample B contained a larger amount of strong acidic sites (Figure 2) compared to sample A.
407
Table 2 Physicochemical properties of the prepared MCM-41 samples sample A sample B removal of the template calcination extraction EtOH/NH4NO3 H NiMo H NiMo BET surface area (m2 g-l) 770 390 960 620 pore volume (cm 3 g-l) 0.5 0.3 0.8 0.4 BJH pore diameter (nm) 1.8 1.5 2.2 1.8 A1 (wt.-%) 2.76 n.d. 2.68 n.d. TPAD (gmol NH3 g-l) 610 720 1080 1120 desorption < 550 ~ 500 718 910 1080 desorption > 550 ~ 110 2 170 40 (%) (18.5) (0.3) (15.7) (3.6) TPAD (#mol NH3 m "z)
0.8
1.8
1.1
sample C extraction EtOH/HNO3 H NiMo 1200 1000 1.1 1.0 2.1 1.9 1.95 n.d. 650 1090 520 1060 130 30 (20.0) (2.8)
1.8
0.5
1.1
100500
'2:
l:i
9 6O 4O
;
2;
4; time
6;
(rain)
8;
~
\
9149
/
0
60
9
.~
5 9
;,
;
80
_ _ . m sample A - - = - - sample B
30.
20.
~
40
4O
sample B
- - - o - - NiMo/USY P
v
~
20
Figure 2. TPAD profiles of the different modified NiMo/MCM-41 samples.
i
~9
E
200~
time (min)
ram--
15
300 B
100
Figure 1. TPAD profiles of the different modified H-MCM-41 samples.
9
,.-..
o
1 400~"
400~
;
-~
O.
~ ~o--,~ 9
;
;
10.
1'0
number of C atoms
Figure 3. Products yield v s . C-number for the hydrocracking n-decane at 400 ~
~.~:~m~..._~_..~ ~ ~
~
~,
;
;
-;
;
;
1'o
number of C atoms
Figure 4. Products yield v s . C-number for the hydrocracking n-decane at 400 ~
408
The conversion rates for the hydrocracking of n-decane using different NiMo-containing MCM-41 materials were comparable. This catalyst samples led to a characteristic product distribution which was different from one of the NiMo/USY (Figure 3). However, the observed selectivities were comparable to the selectivity of a commercial NiMo/amorphous aluminosilicate (NiMo/AAS). Both catalysts reinforced the yield of C7-C9 hydrocarbons. The main reaction products yielded by the zeolitic catalyst were C3C6 hydrocarbons. The extracted sample B showed lower methane production compared to the conventionally modified sample A and even produced more C8-C9as well as i-C10 hydrocarbons (Figure 4).
4. CONCLUSION The template removal by extraction stabilized the mesoporous structure. The number and the density of the acidic sites of the H forms were increased by the extractive pretreatment. The increased dealumination reduced this effect. The investigated NiMo-containing MCM-41 catalysts achieved conversion rates comparable to the rates of commercial hydrocracking catalyst NiMo/AAS under the chosen conditions. The selectivity of these catalysts was comparable but different from that of the zeolite based hydrocracking catalyst. The amorphous structure of the pore walls could be responsible for this catalytic behaviour. The increase of the relative amount of strong acidic sites, caused by the extractive template removal, and the induced electron deficiency state of the metal particles led to a decrease of hydrogenolysis during the hydrocracking of n-decane.
REFERENCES 1. A. Corma, A. Martinez, V. Martinez-Soria, J.B. Mont6n, J. Catal. 153 (1995) 25-31. 2. R. Schmidt, D. Akporiaye, M. St6cker, and O. H. Ellestad, Stud. Surf. Sci. Catal., 84 (1994) 61. 3. S. Hitz, R. Prins, J. Catal., 168 (1997) 194-206. 4. A. Liepold, K. Roos, R. Reschetilowski, A.P. Esculcas, J. Rocha, A. Philippou and M.W. Anderson, J. Chem. Soc., Faraday Trans., 92 (1996) 4623.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
409
Selective hydrodesulfurization technology of cracked gasoline for gasoline pool in 2005 M i n g f e n g Li, H o n g N i e , Y a h u a Shi and D a d o n g Li R e s e a r c h In s t i t u t e ER.China
of
Petroleum
Processing
(R IPP),
SINOPEC,
Beijing
100083,
Abstract The distribution of sulfur and hydrocarbon of commercial FCC gasoline has been analyzed. The effect of isomerization on octane number is also investigated. Based on the above results, two kinds of hydrotreating processes are chosen for producing low sulfur FCC gasoline with minimal octane loss in order to meet the requirement of European gasoline in 2000 and 2005, respectively. 1. INDRODUCTION It is well known that air pollution is a serious environmental problem. A major source of air pollution worldwide is the exhaust from fuel combustion. Experimental data show that reducing sulfur content in gasoline may be one of the best ways to improve automobile emissions [~1.The sulfur content in gasoline will be limited to less than 130ppm and 150ppm in the United States and Europe in 2000, respectively. More restrictive standard requires less than 30ppm of sulfur in gasoline at California and 50ppm sulfur for the Europe in 2 0 0 5 [2'31 . The components of the gasoline pool in a refinery vary from case to case. FCC gasoline in gasoline pool contributes about 70-~75% in China comparing with 40 percent in the United States. On the other hand, approximate 90-98% of the sulfur in the pool are from FCC gasoline. So it is necessary to pay more attention to study the desulfurization of FCC gasoline with minimal octane loss. There are several technologies reported for producing low sulfur FCC gasoline [4-14]. From them, prehydrotreating technology for FCC feed may be one of the most acceptable technologies. But considering the capital investment and other factors, alternative post-treating process seems the most likely schedule for achieving sulfur reduction in the pool. However, the olefin hydrogenation (HYD) causing octane performance of the hydrofinished gasoline to be severely degraded is the main disadvantage of this option. In this paper, on the base of analyzing the distribution and components of sulfur and PONA of commercial FCC gasoline, two kinds of post-hydrotreating processes were chosen for producing low sulfur FCC gasoline with minimal octane loss in order to meeting the requirement of European gasoline in 2000 and 2005, respectively.
2. EXPERIMENTS 2.1 Preparation of catalyst There are three kinds of hydrotreating catalysts used in the present work. Among them, RS catalyst is a commercial catalyst, which is used in naphtha hydrodesulfurization (HDS). Both E-1 and E-2 are hydrotreating catalysts with different acidity. Prior to use, all catalyst are sulfide& The experiments were conducted in an isothermal, cocurrent, up-flow, fixed-bed reactor containing 50ml of the catalyst in the central zone.
2.2 Thermodynamic equilibrium and calculation of octane number In order to study the necessity of the skeletal isomerization of the paraffin and olefin, the thermodynamic
410 equilibrium of C5--~C7paraffins and C5, C6 olefins are calculated [Js]. The equilibrium temperature is selected at 600K because this temperature is mostly used in HDS reaction, at which HDS could be larger than 95%. The octane number of isomers at thermodynamic equilibrium is calculated based on individual hydrocarbon' s octane number. Meanwhile, for olefin isomers, the effect of HYD is taken into account. Octane number of a typical FCC gasoline' s hydrocarbon components is also calculated for comparison.
3. RESULT AND DISCUSSION 3.1 Sulfur and hydrocarbon distribution in commercial FCC gasoline A typical sample of FCC gasoline has been analyzed. The sample is divided into several fractions. Each part is detected for its S, olefin and aromatics. Figure 1 shows the distribution of sulfur, aromatics and olefin in different FCC gasoline fractions. It indicates that sulfur and aromatics concentrate on heavy fraction and olefin concentrates on light fraction.
3.2 Hydrotreating of heavy FCC gasoline fraction According to the distribution of FCC gasoline components, a two step process is chosen. First, total FCC gasoline is split into light and heavy Figure 1. Compound distribution of streams. Then, the heavy FCC gasoline would commercial FCC gasoline subject to catalytic hydrodesulfurization, and whether or not the light FCC gasoline is treated in a process for extraction of mercaptan sulfur depend on the mercaptan content in it. This process is very popular, but the choice of catalyst for hydrotreating heavy FCC gasoline is the key. The catalyst should have following features: High HDS activity, high selectivity of HDS/HYD and low aromatic saturation. RS series catalysts developed by RIPP are chosen in our experiments. The research of the relationship between olefin saturation and octane loss indicates that if the total olefin saturation were less than 10%, the octane loss would be less than 1.0. Some test results are illustrated in Table 2. It provides the flexibility to treat cracked gasoline with varying boiling range, sulfur content and to meet different product requirements. Table 2. Hydrotreating result of FCC gasoline heavy fraction Feed I
Full Boiling Range FCC Gsao. Heavy FCC Gaso. Treating conditions (1) (2) Light FCC Gaso. Lt. FCC + Desulf. Hv. FCC (1) Lt. FCC + Desulf. Hv. FCC (2)
Feed II
RON
MON
91 88.3
79 78.6
Sulfur, ppm 1471 1921
87.0 84.8 94.2 90.1 89.6
77.7 77.1 81.5 78.6 78.3
608 142 787 687 308
RON
MON
91.2
79.3
Sulfur, ppm 425
88.5 87.6
77.5 77.1
45 30
411 From the results shown in table 2, it is found that, after the heavy fraction of FCC gasoline feed I is hydrotreated with RS catalyst, sulfur content decreases significantly, only 142ppm detected at treating condition (2). With the light straight run re-blended, the overall sulfur content become 308ppm, and the octane loss becomes tolerable, only 1.4 research and 0.7 motor octane number lost. Assuming that the FCC gasoline represents 1/3 of the pool, the sulfur in the overall pool would be about l l0ppm as the FCC naphtha contributes to 95% of sulfur, which will meet the requirement of European gasoline regulation in 2000. For feed II, sulfur content in product is less than 50ppm directly, which will meet the requirement of European gasoline regulation in 2005. However, it is implied from the above results that the hydrotreating method of heavy FCC gasoline fraction .is limited. Only when the sulfur content in feed is not very high (less than 1500ppm) or the requirement of product sulfur is not very strict (about 300ppm), the loss of road octane number could be controlled within 1.0, sometimes less than 0.5. Otherwise, in order to reduce the sulfur content, the cut-point temperature should be decreased, that must lead to high octane number loss. 3.3 Effect of isomerization on octane number of FCC gasoline Hydrotreating of heavy FCC gasoline fraction could only decrease but not clear up the octane number loss. By studying the components of paraffin and olefin in commercial FCC naphtha, it is found that most of them is in linear or monomethyl style. We know that only olefin and branched paraffin have high octane number. During hydrogenation reaction, linear and monomethyl olefin are saturated to low octane number alkane causing remarkable octane number loss. Isomerization can convert linear alkenes and alkanes into branched alkenes and alkanes having the same carbon numbers. In this section, effects of isomerization on octane number of C:-C7 paraffins and C5, C6 olefins are studied ( Table3, Table 4 ). From the above investigation, it is concluded that isomerization of olefins and paraffins makes the octane number of the hydrocarbon mixture increase significantly, especially for long chain olefins or paraffins, which is beneficial to compensate to octane number loss caused by olefin saturation during FCC gasoline or its fractions HDS. But we still learned that although isomerization is effective, comparing with untreated feed, the product' s octane number still decreased even after isomerization if the HYD is larger to some degree. That' s to say the new isomerization/HDS bifunctional catalyst should have characteristics of isomerization and selective hydrodesulfurization.
Table 3. Octane number a change of paraffin at thermodynamic equilibrium C~ paraffins ON of feed 82.3 ON chan~e of isomer at equilibrium +0.9 a Octane number means (MON+RON)/2 Table 4. Octane number change of olefin at thermodynamic equilibrium C5 olefins 50 % Feed HYD ON change of feed at different 89.4 -3.8 HYD ON change of isomer at equilibrium with different HYD +1.6 -1.6 comparing with feed A ON +1.6 +2.2
(calculation value) C6 paraffins 70.3 +2.4
C 7 paraffins
45.4 +19.6
(calculation value) C 6 olefins
100 % HYD
Feed
50 % HYD
100% HYD
-12.1
90.6
-16
.-32.2
-4.2
-3.1
-11.5
-20.3
+7.9
-3.1
+4.5
+11.9
3.4 Performance of isomer selective hydrotreating catalyst A new kind of hydrotreating catalyst E-1 is prepared. Using this catalyst in full range FCC gasoline hydrodesulfurization, at 95% HDS, the loss of (RON+MON)/2 is only 1.6, and the olefin saturation is larger
412
than 70%. If cracked gasoline contributes at 1/3 of gasoline pool, a 30ppm sulfur and 4v% olefin level in gasoline pool meeting the maximum sulfur and olefin specification of CARB phase II could be achieved by this technology, and also could meet the requirement of European gasoline regulation in 2005. Table 5 Performance of E catalyst hydrotreating full range cracked gasoline Feed III Catalyst
E-1
E-2
Feed Sulfur(ppm)
1473
Olefin, wt.%
26.0
RON/MON
91.6/79.3
Product Sulfur(ppm)
91
41
Olefin, wt.%
6.9
3.4
RON/MON
89.5/78.3
77.6/71.3
4. CONCLUSION 1 2 3 4
By studying the distribution of FCC gasoline, it is found that sulfur and aromatics concentrate on heavy fraction and olefin concentrates on light fraction. The two step process containing heavy FCC gasoline fraction hydrodesulfurization is proven to be effective for producing low sulfur gasoline meeting the requirement of European in 2000 and 2005. By studying the thermodynamic equilibrium of paraffin and olefin in commercial FCC naphtha at 600K, it is found that isomerization makes the octane number increase significantly, especially for long chain olefins or paraffins. A new kind of catalyst containing isomerization and selective hydrogenation functions is prepared. Comparing with traditional hydrotreating catalyst, the novel one has less octane loss at closed HDS.
5. R E F E R E N C E 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15
L.D. Krenzke, J.E. Kennedy, K. Baron, NPRA AM-96-67 C. Patrick, Oil & Gas Journal, 228 (1996) 21 Hart' s European Fuel News, 13:2 (1998) 1 P.H. Desai, D.A. Keyworth, M.Y. Asim, NPRA AM-92-44 J.-L. Nocca, R.M. Gialella, J., Cosyns, NPRA AM-95-50 C. Sudhakar, M.R. Cesar, Heinrich R.A., US Patent No. 5 525 211 G.K. Simon, P.D. Naperville, Pablo, D.H., US Patent No. 5 348 928 L.D. Fletcher, S.M. Sarli, S. Shih, US Patent No. 5 503 734 M.S. Sarli, D.L. Fletcher, T.L. Hilbert, NPRA AM-94-39 D.C. Martindale, G.J. Amos, K. Baron, NPRA AM-97-25 R Forte, US Patent No. 5 582 714 L.T. Nemeth, US Patent No. 5 360 536 D. Hearn, G.R. Gildert, H.M. Putman, US Patent No. 5 595 634 R.M. Gialella, J.W. Andrews, J. Cosyns, NPRA AM-92-21 Adalbert Farkas (eds.), Physical Chemistry of the Hydrocarbons, CHAPTER 9, Academic Publishers, New York, 1950.
Press Inc.,
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
413
H Y D R O T R E A T I N G CATALYSTS ON ALUMINA, T I T A N I A OR Z I R C O N I A FROM ETHANOL/WATER SOLUTIONS OF HETEROPOLYACIDS
Luis Pizzio, Patricia V~zquez, Carmen C~ceres and Mirta Blanco
Centro de Investigaci6n y Desarrollo en Procesos Catallticos -CINDECAUniversidad Nacional de La Plata, CONICET, 47 N~ 1900- La Plata, ARGENTINA.
Abstract Supported P N i M o ( ~ catalysts for hydrodesulfurization and hydrodenitrogenation were prepared. Heteropolyacids were used as precursors and supported on alumina, titania and zirconia. The tungstophosphoric and molybdophosphoric acids showed a stable Keggin-type structure when the supports were zirconia and titania. However, the heteropolyacids on alumina were degraded during the drying and/or calcination of the impregnated samples. The catalysts prepared on alumina had the highest activity, in the order of commercial ones. The catalysts on titania were more active than the samples supported on zirconia.
1. I N T R O D U C T I O N
More restrictive environmental legislation and the need to utilize low quality hydrocarbon feedstocks have focused increased attention on the catalytic processes involved in converting these feeds into clean burning fuels. From this point, a new generation of hydrotreating catalysts were obtained, in which the nature of the support and of the active phase have been varied [1, 2]. The intensive research [3] during the last years on the promoting effect of P in Mo(W)Ni/alumina hydrotreating catalysts resulted in a number of explanations on the mechanism of its effect on the activity of these catalysts. Differences in P effect on hydrodenitrogenation (HDN) activity seem to be mostly related to the model compound and reaction conditions applied [4], while the conclusions of P effect in the hydrodesulfurization (HDS) catalysts show that it is due to different method of preparation, different structure and dispersion of active species in the catalysts [5]. Therefore, as phosphorus is regarded as one of the most effective additives on these catalysts, heteropolyacids (HPA) with Mo(W) and P present in their Keggintype structure have been used as precursors of these catalysts and so dopant and active phase are incorporated simultaneously [6].
414
The HPA catalytic activity is directly related to the nature and subsequent stability of adsorbed species in the impregnation stage. So, it was studied the nature of the species present in solution and on the solid for the tungstophosphoric acid (TPA) from ethanol-water solutions and different supports, as silica, titania, carbon and different aluminas [7]. For the molybdophosphoric acid (MPA) it was studied, among other variables, the behavior of species so much from solutions in organic solvents as in water and for alumina as for different supports [8]. In this work, the preparation of new hydrotreating catalysts, for HDS as well as HDN, is shown. MPA and TPA were supported on alumina, titania and zirconia, using solutions with ethanol/water as solvent. The characterization of the species present in the impregnating solutions as well as in the obtained catalysts, previous and after to its sulphurization, was carried out. The pyridine HDN and thiophene HDS activities were measured.
2. EXPERIMENTAL Supports. The supports used were TiO2 anatase (Ti) Riedel-de H~ien (SBET: 9.8 m2/g), A1203 Spheralite (A1) (SBET: 282 m2/g) and ZrO2 (Zr) (SBET: 181 m2/g). Zirconium hydroxide gel was prepared from an aqueous zirconium oxychloride solution 0.07 M (Fluka) by adding drop wise an ammonium hydroxide solution (Merk, 25 % ammonia) up to pH=9.5. The hydrogel was aged at room temperature for 72 h and then filtered and thoroughly washed with deionized water. It was dried in air at 110 ~ for 24 h and finally calcined at 310 ~ during 24 h. Preparation Of Catalysts. PMo(W)Ni/support catalysts with a content of 0.08 mol Mo(W)/g and 0.035 mol Ni/g, were prepared by means of pore filling two step impregnation: first, MPA (H3PMo1204o.xH20, Merck) or TPA (H3PW1204o.xH20, Fluka) were impregnated on different supports by employing ethanol-water 50 % v/v (e-w). This samples are denoted MPA(TPA)-support. In the second step and after drying the solids at room temperature, these were impregnated with e-w solution of nickel nitrate. Solids thus obtained, MPA(TPA)-Ni-support, were dried at room temperature for 24 h and then calcined at 350 ~ for I h. UV-visible spectroscopy. Spectra of solutions, in the range of 200-600 nm, were obtained with a Varian Super Scan 3 double beam UV-visible spectrophotometer with built-in recorder, using quartz cells of 0.5 mm optical path. Nuclear m a g n e t i c r e s o n a n c e spectroscopy. The solutions were analyzed by 31p NMR. A Bruker MSL-300 equipment linked to a "SOLIDCYC.DC" pulse program was utilized, using 5 ~s pulses, a repetition time of 10 s and a frequency of 121.496 MHz for 31p at room temperature (24 ~ being the resolution of 3.052 Hz per point. Phosphoric acid 85% was employed as external reference. Fourier transform infrared spectroscopy. A Bruker IFS 66 equipment, pellets in BrK and a measuring range of 400-1500 cm -1 were used to obtain the FT-IR
415
spectra of the supports, bulk MPA (TPA), and supported MPA (TPA) samples dried at room temperature and calcined at different temperatures. Diffuse reflectance spectroscopy. The solid samples were studied with a UVvisible Varian Super Scan 3 equipment, fitted with a diffuse reflectance chamber with inner surface of BaSO4. DRS spectra of supports, bulk MPA (TPA) and dried at room temperature and calcined MPA (TPA) alumina supported samples were recorded in the range of 200-600 nm. X - r a y d i f f r a c t i o n . The XRD patterns of the supports and samples dried at room temperature were obtained by using a Phillips PW-1714 diffractometer with builtin recorder. In this measurements, the following conditions were used: CuKa radiation (1.5417 A), nickel filter, 30 mA and 40 kV in the high voltage source, scanning angle (20) from 5 to 60 ~ and scanning rate of 1~per minute. Temperature P r o g r a m m e d Reduction. The TPR patterns were performed in a home-made equipment up to 1000 ~ using a H2 flow of 60 cm3/min. Catalytic activities. The activity measures in HDS and HDN reactions were accomplished in a fixed bed reactor at 3 MPa, 280 ~ a H2/HC molar ratio:5 and LHSV:53 h -1. The catalysts were sulphurized in situ at 280 ~ The liquid feed was a mixture of ciclohexane-pyridine (2000 ppm N2) for HDN and ciclohexanethiophene (6000 ppm S) for HDN.
3. R E S U L T S AND D I S C U S S I O N
3.1. Characterization of the solutions UV-visible spectroscopy. UV-visible
spectra of the TPA and MPA impregnating solution in e-w before the contact with the supports show bands at 265 and 310 nm, which are the characteristic charge transfer bands of the [PW12040] 3- and [PMo12040] 3- heteropolyanions, respectively. Reduction of HPA proceeds without substantial change of their structure with addition of a certain characteristic number of electrons. On reduction they are coloured mainly blue, producing the so called heteropoly blues (HPB) which have a broad absorption band at about 700 nm [9]. It was observed color variations with time both in the original and final MPA solutions; they are initially yellow and become greenish clue to HPB formation. Nuclear m a g n e t i c resonance spectroscopy. For TPA and MPA solutions, the 31p NMR spectra recorded before the contact with the different supports show a line at a chemical shift of-15.1 and -3.6 ppm [6] which have been assigned to the [PW12040] 3- and [PMo12040]3- species.
3.2. Characterization of the solids Fourier transform infrared spectroscopy. Each M3013 triad in the Keggin structure has a common oxygen vertex connected to the central heteroatom X-Oa. The other three classes of symmetric-equivalent oxygens in the a isomer of Keggin
416
structure are: M-Oh-M, connecting two M3013 units by corner sharing; M-Or connecting two M3013 units by edge sharing and terminal oxygens Od-M. The maxima of the stretching modes of vibration for IR absorption bands of MPA (TPA) are: P-Oa, 1070 (1081) cm-1; M=Od, 965 (982) cml; M-Oh-M, 870 (888) cm -1 and M-Oc-M, 790 (793) cm -1 . The FT-IR spectra of MPA(TPA)-Ti dried samples (Figure 1) show the characteristic bands of MPA and TPA ^ / acids, though some of them overlap with those of the supports. The characteristic bands of MPA and TPA also appear in FT-IR spectra of MPA(TPA)-Zr dried samples (Fig. 1). UJ Therefore, the spectra indicate that Z the species present in the dried samples are the [PW1204o]3- and .N [PMo12040] 3- anions. The spectra of TPA and MPA I--supported on both TiO2 and ZrO2 calcined at 350 ~ show similar characteristics to those of dried samples. The study of the catalysts by FT-IR allowed us to verify that the species present in the solid coincided with those in the impregnating solution. I 1400 1200 1000 800 600 400 On the other hand, the FT-IR study shows that both the [PW12040]3- and WAVENUMBER (cm-~) [PMo12040] 3- anion do not undergo irreversible degradation during the thermal treatment until 350 ~ Figure 1. FT-IR spectra of MPA(TPA)Ti(Zr) dried samples, supports, MPA and TPA acids.
Diffuse reflectance spectroscopy. The charge transfer absorption spectra of most non-reduced polyanions appear in the 200 - 500 nm region, and consist of bands which may be ascribed to oxygen-to-metal transfers [10]. For MPA, Mo octahedral exhibits two absorption bands at 220 and 260 nm and another band at higher wavelength. For TPA, the spectrum presents a band at 220-225 nm and another broad band that extends from 260 to 390 nm. The spectrum of MPA-alumina dried sample (Figure 2) indicates the presence of the undegraded Keggin phase. MPA degradation becomes evident in the spectrum of the sample calcined at 350 ~
417
The DRS spectrum of TPA-alumina dried sample (Figure 2) presents a band at 212 nm and another broad band that extends from 260 to 350 nm. This is quite different to the bulk TPA spectrum.
~.
TPA-AI-C TPA-AI-D
TPA =
I
200
,
I
300
,
I
400
A~20;]
500
600
WAVELENGHT (nm)
Figure
2.- DRS of TPA and MPA on alumina
I I
l MPA
s
~
MPA-Zr
i
10
20
:30
2O
rPA-Zr
40
50
60
F i g u r e 3.- DRX patterns of TPA and MPA on zirconia.
The intensity increase of the band at 212 nm can be attributed to the presence of WO4 e- species on the solid, resulting from the possible depolymerization of [PW12Oa0]3during the interaction with these support. The shrinkage of the broad band is the result of the depolymerization of the [PW1204o] 3anion into [PW11039] 7- species, it was confirmed by 31p NMR-MAS [7]. The spectrum of TPA-alumina sample calcined at 350 ~ is similar to t h a t of the dried sample.
X-ray diffraction. The XRD pattern of TPA zirconia based catalyst presents lines corresponding to the phase H3PW1204o.6H20 (Figure 3). When the sample is calcined at 180 ~ the intensity of those lines decrease, due to the dehydration of the phase H3PW12040.6H20 and the subsequent formation of the H3PWleO4o phase. In the XRD pattern of sample calcined at 350 ~ neither the lines corresponding to H3PW12040 nor H3PWleO40.6HeO phases are present; this shows t h a t the H3PWleO40 is present as a noncrystalline phase.
418
On the other hand, the XRD patterns of the other dried catalysts are similar to those of the corresponding supports, so neither the diffraction lines of TPA or MPA nor those of other related crystalline phases were detected.
Temperature Programmed Reduction. TPR tests were carried out on sulfided samples of the catalysts, in order to study the interaction between the active components and the support. TPR patterns of the catalysts prepared from TPA and MPA on alumina, titania and zirconia present a band showing a maximum at 240-300 ~ This would indicate that sulfided catalysts contain sulfided molybdenum and nickel species interacting with the support [11], which are more easily reducible than the species present in the oxidic precursor [3,12]. It must be emphasized that the TPR spectra of the prepared catalysts present differences in this zone. Samples supported on titania and zirconia (TPA(MPA)-Ni-Ti(Zr)) exhibit less intense peaks compared with the samples on alumina, as shown in Figure 4, and so a decrease in the number of reducible sites. One possible reason is a weaker interaction of molybdenum(tungsten)nickel sulfide phase and the support, and a greater sulfurization of this phase. Another possible cause is the lower formation of this sulfide phase, the NiMo(W)S structure, which has been regarded as the active phase for HDS and HDN reactions [13]. However, the 12 extent of NiMo(W)S phase 11 formation is a consequence of the 10 dispersion of oxidic precursors of 9 this phase, which in turn, depends MP~-AI on the support surface area and the concentration of surface OH groups. 6 It is also important to take into account the distribution of these ~4 groups in surface supports. In the ~ 3 case of titania they are uniformly 2 rR~gl-li IP/M~-Zr distributed and in parallel rows 1 on the alumina [14]. N 0
Figure 4. TPR peak area.
3.3. Catalytic activities. The activity results of PNiW(Mo) catalysts supported on A1, Ti and Zr, obtained from HPA, are shown in Figure 5 for the HDS reaction, and in Figure 6 for the HDN one. The catalysts prepared on A1 have the highest activity, in the order of commercial ones. These figures also show that the TPA(MPA)-NiTi catalysts are more active t h a n the samples supported on Zr.
419 30
.
.
.
.
.
.
MP~ 44:
25
T P ~ -r +4 .~ §
~m 2:
.
1PAM-~
MPA-N~
i::_}
Z
50
15
:LT:
m > Z 9 3O (.9
TPA-IWri
:LT:
MPA-I~Zr
>
4+4
~
~
=
.....
I"PA-I~Zr
t!i
0
Figure 5. Catalytic activities for the HDS reaction.
~
,lili i
-I~ M P A ~ f i
":+-)~
~MPA44-7.r
Ci~
I, i,
Figure 6. Catalytic activities for the HDN reaction.
The TPR results obtained enabled to suppose that "NiMoCW)S" phase is formed after a standard sulphidation, no matter what the carrier is. The samples on alumina are the most active, probably as a consequence of the fact that this support has the greatest surface area. This can be 60 - i - H [ ~ MPA .' observed in the Figure 7, in which the HDS and HDN activities are correlated with the support surface area. As above-mentioned, when the surface "" " " ... ,,. .. " ~ I area is increased, the possibility of a ................ 10 ~G . . . . . . . . . . . . . . . . . . . . . . . . . . . . . . - I - " " better dispersion is greater. Moreover, 0 I the formation of the NiMo(W)S active 0 50 100 150 2[]0 250 300 phase of the presursors is greater. SURFAO:: AREA [rn2/g] Notwithstanding, the sample supported on zirconia does not have catalytic values between the Figure 7. HDN and HDS activities in function active corresponding values of the samples of support surface area. TPA(MPA)-Ni-A1 and TPA(MPA)-Ni-Ti.
S'
,
,
,
,
i
Probably, this is due to the fact that the precursor dispersion is controlled in the samples on zirconia by another factor different of surface area. For example, it was proposed that the MoS2 slabs sit vertically on the alumina carrier, whereas they lie parallel to the plane of zirconia [15].
420
The results obtained show the influence on the HDS and HDN catalytic activity of the differences in the genesis, dispersion, stability and morphology of the NiMo(W)S phase, depending on the support. On the other hand, as a result of the study carried out on the species present in the impregnating solutions and in the calcined catalysts, it was concluded that the HPA show a stable Keggin-type structure when the supports are zirconia and titania. However, on alumina the HPA are degraded during the drying and/or calcination of the impregnated samples. In this case, the active phase would probably be more easily formed
ACKNOWLEDGEMENTS The authors thank L.Osiglio, G. VaUe and D. Pefia for their contribution.
REFERENCES
1 2 3 4 5 6 7 8 9 10 11 12 13
14 15
R. Navarro, B. Pawelec, J. Fierro and P. Vasudevan, Appl. Catal. A: General 148 (1996) 23. M. Breyse, J. Portefaix and M. Vrinat, Catal. Today 10 (1991) 489. P. Atanasova, R. LSpez Cordero, L. Mintchev, T. Halachev and A. L. Agudo, Appl. Catal. A: General 159 (1997) 269. H. Topsoe, B.C. Clausen, N-Y. Topsoe and P. Zeuthen, Stud. Surf. Sci. Catal. 53 (1990) 77. M. Jain and R. Prins, Bull. Soc. Chim. Belg. 104 (1995) 231. L. Pizzio, P.V~zquez, M. Gonzgdez, M. Blanco, C. C~ceres and H. Thomas, Stud. Surf. Sci. Catal. 106 (1997) 535. L. Pizzio, C. Chceres and M. Blanco, Appl. Catal. A: General 167 (1998) 283. M. Castillo, P. V~zquez, M. Blanco and C. C~ceres, J. Chem. Soc. Faraday Trans. 92(17) (1996) 3239. E. Papaconstantinou, Chem. Soc. Rev. 18 (1989) 1. M. Pope, Heteropoly and Isopoly Oxometallates, Springer, Berlin, 1983. J. Laine, F. Severino and M. Labady, Journal of Catalysis, 147 (1994) 355. L. Pizzio, P. V~zquez, C. C~ceres and M. Blanco, unpublished data. H. Topsoe, R. Candia, N-Y Topsoe and B.S. Clausen, Bull. Soc. Chim. Belg., 93 (1984) 783. C. V. C~ceres, J. L. G. Fierro, J. Lhzaro, A. LSpez Agudo and J. Soria, Journal of Catalysis, 122 (1990) 113. F. Maug6, J. C. Duchet, J. C. Lavalley, S. Houssenbay, E. Payen, J. Grimblot and S. Kasztelan, Catalysis Today, 10 (1991) 561.
Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
A XANES temperature-programmed NiMo/SiO2 hydrotreating catalysts
421
sulphidation
study
of m o d i f i e d
R. Cattaneo, T. Shido and R. Prins Laboratory for Technical Chemistry, Federal Institute of Technology (ETH), CH-8092 Zurich, Switzerland Abstract The thiophene HDS activity of NiMo/SiO2 catalysts, prepared by the addition of chelating ligands, was correlated with the sulphidation behaviour of Ni and Mo as studied by means of X-ray Absorption Near-Edge Structure spectroscopy. It was concluded that the improvement in the catalytic activity, as a result of the addition of the chelating ligands, is due in part to the sulphidation of Mo at lower temperatures, but mainly to the effects of the ligands on Ni. 1. I N T R O D U C T I O N Supported nickel-molybdenum sulphide catalysts are used extensively in the hydrotreatement of petroleum. The addition of chelating ligands, such as nitrilotriacetic acid (NTA), to the impregnation solution has a beneficial effect on the activity of NiMo catalysts. Temperature-programmed techniques are often selected for the characterisation of the structure of hydrodesulphurisation catalysts, since these catalysts are always exposed to mixtures of H2 and H2S during presulphiding in order to convert the oxidic precursor into the sulphidic, active catalyst. In this study, the sulphidation of Ni and Mo was investigated by measuring in situ X-ray Absorption Near-Edge Structure (XANES) spectra of SiO2-supported catalysts, modified by the addition of chelating ligands. 2. E X P E R I M E N T A L 2.1. S a m p l e s All catalysts were prepared by pore volume impregnation with a solution containing MOO3, Ni(NO3)2-6(H20) and the organic ligand, as described in [1]. Several ligands were tested. The support was SiO2 (C560 Chemie Uetikon), with a particle size of 125 to 250 ~m, a BET surface of 565 m2/g, a BET pore volume of 0.83 ml/g, dried overnight at 120~ prior to impregnation. After impregnation, the powder was dried at 120~ overnight. No calcination was carried out at higher temperatures. (Metal loadings: 7 wt% Mo, 1.3 wt% Ni, molar ratio Ni:Mo=0.3:l). Presulphiding and thiophene HDS reaction occurred at atmospheric pressure in an apparatus that was a modified version of the flow system described in [2]. The precursors were sulphided at 400~ (heating rate 6~ for 2 h in a mixture of 10% H2S in H2 (Messer Griesheim 3.0). The activity of all the catalysts was tested
422
in the hydrodesulphurization of thiophene at 400~ The feed, consisting of 3% thiophene in H2, was obtained by bubbling H2 through a series of four thiophene s a t u r a t o r s cooled to 2~ The product s t r e a m was analyzed on line with a HP5890 gas chromatograph. 2.2. X A N E S The XANES data were collected at the beam line X1 at HASYLAB (Desy, H a m b u r g ) [3]. The samples were pressed into self-supporting wafers and sulphided in situ in an EXAFS cell [4]. The sulphidation conditions were the same as mentioned above. A spectrum was collected every 3 min. Series of spectra were collected for the Mo and Ni K-edges. The spectra at the edge jump were fitted by a linear function of the m e a s u r e d oxidic and sulphided samples to e s t i m a t e the ratio of sulphided Ni and Mo during the sulphidation. The following formula was used: (XANES observed) - fl 9 (XANES of oxidic state) + f2 9 (XANES of sulphided state) where fl and f2 are the fractions of the oxidic and sulphided states, respectively. Both fl and f~. were treated as free p a r a m e t e r s in this analysis. The reference for the oxidic state was the fresh sample, for the sulphided state, the sample sulphided at 400~ for 30 min.
,!
400 ~
3. R E S U L T S
3.2. Ni K - e d g e X A N E S Fig. 1 presents a series of XANES spectra m e a s u r e d during the sulphidation of the catalyst prepared without ligands. It is easy to distinguish sulphide and oxide environments by m e a s u r i n g the Ni K-edge XANES spectra, because they show a strong whiteline when atoms of the second row are present in the first coordination sphere of Ni. Due to the non-ionic character of the NiS bond, this feature is absent for sulphided systems [5]. In order to quantify the effect of the various ligands on the sulphidation t e m p e r a t u r e of Ni, all collected XANES were simulated as described in the experimental section. Fig. 2 shows the resulting fractions of Ni in the sulphided state. These plots show t h a t EN shifts the sulphidation of Ni to lower t e m p e r a t u r e s . In the presence of EN, Ni starts to be sulphided at room
9
~
o
291 269 245 223 202 181 161 141 119 96 75 46 fresh
<
9
8320
8340
8360
8380
X-rays energy [eV]
Figure 1. Ni K-edge XANES spectra m e a s u r e d during the sulphidation of NiMo/SiOz (no ligand) plotted as a function of the sulphidation t e m p e r a t u r e .
423
temperature. The same is true for the catalyst prepared with the smallest amount of NTA (NTA:Ni=0.25). Higher concentrations of NTA and EDTA cause a delay in the sulphidation of Ni. 100 3.3. Mo K - e d g e XANES 90 .v..~ Mo K-edge XANES spectra, Z 8o measured during the sulphidation ,-~ 7o m/e"' 9 no hgand ~ 60 x EN:Nl=2.33 of the catalyst prepared without 50 x NTA:Nl=0.25 ligands, are presented in Fig. 3. ~:~ 40 o NTA:Ni=0.5 Mo K-edge XANES has been sub~ 3o + NTA:NI=I.5 20 o NTA:Ni= 3 ject of several studies, according to - EDTA:Ni=I which the assignment of the first T 0 0 100 200 300 400 three components in the near-edge spectrum of oxidic molybdenum Sulphidation Temperature [~ are due to l s - + 4 d , l s - + 5 s and ls--+ 5p excitations [6]. The first Figure 3. Sulphidation profiles of Ni in transition is allowed in a tetradifferent SiO2-supported NiMo catalysts, hedral field, but is forbidden in an according to the XANES measurements. octahedral one. This selection rule breaks down when the octahedral symmetry is distorted. The shapes of the nearedge absorption functions for all the samples can be divided, qualitatively, into three types, all of which correlate with the extent of reduction. In the first type (Fig. 3, 41-194~ the shape is that of the oxide catalysts and the energy of the edge is typical for Mo(VI). As sulphiding progresses, the pre-edge feature fades and the edge moves towards lower energies 400~ (Fig. 3, 209-296~ Finally, after extensive 40O treatment, the near edge absorption attains a 364 346 shape typical for Mo(IV) in MoS2. 331 314 296 In the case of the Mo K-edge, too, all 279 9 263 spectra were simulated by means of a linear 246 226 209 regression. The results are shown in Fig. 4, 194 9 172 from which it is clear that all the catalysts 154 141 125 tend to shift the sulphidation of Mo toward < 107 90 72 lower temperatures. In particular, for the ~ catalyst with NTA:Ni=3, the sulphidation of ~= Mo is drastically altered by the presence of the ligand. Another feature is that, after a first sulphidation (40-140~ Mo remains unchanged according to XANES, and is 19980 20010 20040 20070 further sulphided at temperatures above X-ray energy [eV] 200~ This step-wise sulphidation suggests Figure 3. Mo K-edge XANES the presence of an intermediate product such spectra measured during the as a molybdenum oxysulphide [ M o O 4 . n S n ] 2sulphidation of NiMo/SiO2 (no ligand) plotted as a function of the sulphidation temperature. 9~
~es~
424
4. D I S C U S S I O N
EDTA, NTA and EN can form complexes with nickel. In a previous study, we demonstrated by means of EXAFS, UV-VIS and R a m a n how the presence of such ligands in the catalyst precursors prevents the formation of nickel silicates [1]. NTA was found to have a much stronger effect t h a n EN. According to the results presented here, EN favours the sulphidation of Ni, while a sufficient a m o u n t of NTA (NTA:Ni>0.25) and EDTA retards, it. This effect is correlated with the stability of the formed Ni species. All the ligands have a similar effect on the sulphidation of Mo, a shift of the sulphidation towards lower t e m p e r a t u r e s and an increase in the degree of sulphidation at intermediate temperatures (100-200~ Most ligands form more stable complexes with Ni t h a n with Mo and, only when all Ni is complexed, they can interact with Mo [1]. A comparison with catalytic activity data shows that, in the most active catalysts, 100 nickel is sulphided at the 9o highest temperatures. This delay in sulphidation favours z 80 9 no l i g a n d the formation of active sites, ~ 70 + EN:Ni=2.33 since there is a larger tempe~9 60 rature range for the forma- ~.~ 50 9NTA:Ni=0.5 j" 40 tion of MoS2 crystallites. The .~" 9NTA:Ni=I fact t h a t Mo starts to be ~= 30 9NTA:Ni= 3 20 sulphided at lower temperao 10 ~y/~/"/' - E D T A: N i=1 tures has the same conse0 50 100 150 200 250 300 350 400 quence. In both cases, Ni has the possibility to arrange on Sulphidation T e m p e r a t u r e [~ better ordered MoS2 crystallites. Figure 4. Sulphidation profiles of Mo in different SiO2-supported NiMo catalysts, according to the XANES measurements. 5. C O N C L U S I O N S
jF
//!
The correlation of catalytic activity data with the t e m p e r a t u r e - p r o g r a m m e d sulphidation followed by XANES showed that the activity improves as a result of the earlier sulphidation of Mo for all tested ligands. For ligands such as NTA and EDTA, the retardation of the sulphidation of Ni also has a positive effect on the activity of the obtained catalysts. In the case of Mo, the XANES features suggested the presence of an intermediate sulphidation product. 6. R E F E R E N C E S .
2. 3.
Cattaneo, R., Shido, T., and Prins, R., J. Catal. 185 (1999), in press. Medici, L., and Prins, R., J. Catal. 163, 38 (1996). TrSger, L., Synchr. Rad. News 6, 11 (1997).
425
.
.
6.
Kampers, F. W. H., Maas, T. M. J., Grondelle, J. v., Brinkgreve, P., and Koningsberger, D. C., Rev. Sci. Instrum. 60, 2635 (1989). Louwers, S. P. A., and Prins, R., J. Catal. 133, 94 (1992). Chiu, N.-S., Bauer, S. H., and Johnson, M. F. L., J. Catal. 89, 226 (1984).
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Hydrotreatment and Hydrocracking of Oil Fractions B. Delmon, G.F. Froment and P. Grange (Editors) 9 1999 Elsevier Science B.V. All rights reserved.
HY
427
zeolite-based catalysts for hydrocracking h e a v y oils
K. Honna a, K. Sato b, Y. Araki a, Y. Miki a, N. Matsubayashi b, and H. Shimada a* aTsukuba-branch, Advanced Catalyst Research Laboratory, Petroleum Energy Center, 1-1 Higashi, Tsukuba, Ibaraki 305-8565 Japan bNational Institute of Materials and Chemical Research, 1-1 Higashi, Tsukuba, Ibaraki 305-8565 Japan
Abstract Hydrocracking of Arabian-Heavy atmospheric residue was carried out over Yzeolite supported catalysts. The results indicated that the mesoporosity of zeolite played an important role in the increase in hydrocracking heavy fractions. In addition, it was found that the distribution of the acid sites and hydrogen supply from the sulfide to the acid sites were the important factors to maximize the hydrocracking functions of zeolites. 1. INTRODUCTION There has been an increasing demand for the production of high quality middle distillates from heavy crude oils. To satisfy this demand with economical competitiveness, innovative improvements of hydrocracking (HCK) technology including catalyst developments are needed [1]. The final goal of our project is to develop a bifunctional catalyst that is suitable for hydrocracking heavy oils with residue. For this purpose, we selected Y-zeolites as solid-acid catalysts and investigated their performances in the HCK of different kinds of feedstocks [2]. As the results, it was found that the H CK function of HY zeolites did not work properly for the heavy feedstocks with residue. In the present paper, special attention is paid to i) the effects of mesoporosity of zeolite and ii) the role of hydrogenation (HYD) activity during the reactions of these heavy feedstocks. 2. EXPERIMENTAL The H CK catalysts used in the present study were four kinds of NiW (NiO = 3.5%, WO3 = 24.0%) and NiMo (NiO = 1.7%, MoO3 = 6.7%) sulfide catalysts "Corresponding author
428 supported on USY and HY zeolites. NiW/USY possessed a high HYD activity and a small n u m b e r of acid sites corresponding to a high Si/A1 ratio. The same USY support was used for the p r e p a r a t i o n of NiMo/USY. Two HY supports, HY(L) and HY(S), h a d a similar Si/A1 ratio ( ~ 5 ) but different mesoporosity. As a result, NiMo/HY(L) h a d a higher mesoporosity t h a n NiMo/HY(S) as described in the footnotes of Table 1. The feedstock used was Arabian-Heavy atmospheric residue (AH-AR) with 7.7% of asphaltene. All the reactions were conducted using a batchtype autoclave with an inner volume of 140 cm 3. After charging an autoclave with 10 g of AH-AR, 1 mmol of catalyst as total metal (for example, 0.096 g as Mo) and 9.8 MPa of hydrogen, the reaction was carried out at 410 ~ for 3 h.
3. R E S U L T S A N D D I S C U S S I O N Table 1 shows the product distribution over the catalysts. Molybdenum dithiocarbamate (MoDTC), an oil-soluble catalyst, exhibited almost the same performance as NiMo/A1203 for the asphaltene conversion, although MoDTC had no Ni promoter. Since MoDTC decomposes into MoS2 below 350 ~ highly dispersed fine particles of MoS2 formed during the reaction are supposed to function quite effectively for the heavy fractions. The higher yield of 343 ~ fraction over NiMo/A1203 t h a n MoDTC may be a t t r i b u t e d to the weak acidity of Table 1 Reaction of Arabian heavy atmospheric residue (AH-AR) Product Distribution (~ %) Gas C ~ - 3 4 3 343-520 520-565 565 + Feed 2.9 51.5 11.0 26.9 MoDTC 1.9 16.2 41.7 11.8 28.2 NiMo/A1203 1.7 21.3 38.9 7.8 30.2 NiW/USY 4.9 36.5 25.6 8.2 18.8 NiMo/USY NiMo/HY(L) NiMo/HY(S) NiW/USY:
13.5 16.2 19.6
44.3 40.1 10.7
9.3 13.0 21.1
5.7 5.9 9.9
21.5 15.8 30.6
H2.C #r
(%)
Asp. 7.7 0.22 0.07 5.6
2.59 2.80 2.65
5.7 8.9 8.1
4.12 4.65 5.42
Si/AI=8.4, surface area (SA); 453 m2/g, micro pore volume (PV); 0.181 cm3/g, meso PV (D > 2 nm); 0.065 cm3/g. NiMo/USY: Si/AI=8.4, SA; 555 m2/g, micro PV; 0.205 cm3/g, meso PV 0.134 cm3/g. NiMo/HY(L): Si]AI=5.0, SA; 579 m2/g, micro PV; 0.201 cm3/g, meso PV; 0.124 cm3/g. NiMo/HY(S): Si/AI=5.3, SA; 560 m2/g, micro PV; 0.213 cm3/g, meso PV; 0.054 cm3/g. * H2-C; Hydrogen consumption (wt.%/feed). * The preparation methods of HY(L) and HY(S) will be published elsewhere. * The Si/A1 ratios were determined using the lattice parameters obtained by X-ray diffraction. * The higher yields of 343 ~ fractions over the NiMo/zeolite catalysts than NiW/USY were due to the larger amounts of zeolites charged in the former reactions, because the catalyst was charged so that the amount of active metals were the same (1 mmol) in all the reactions.
429
Table 2 HCK of tetralin over the catalysts Conv. .(%) USY 45.3 NiW/USY 48.8 NiMo/USY NiMo/HY(L) NiMo/HY(S)
51.6 49.4 44.9
Product distribution (%) Gas .1 Mono.2 Di.3 1.0 13.1 25.2 1.4 21.0 21.9
Heavy.4 60.7 55.7
2.1 2.3 2.2
49.3 49.6 50.3
25.6 26.2 25.1
23.0 21.9 22.4
* All the reactions were performed in a batch autoclave (50 cm3), containing 5 cm3 of tetralin, 0.3 g of catalyst and H2 with an initial pressure of 5.9 MPa for 1 h at 350 ~ *1: C1-C4, "2: Monocyclic compounds, "3: Dicyclic compounds (decalin, indanes, naphthalene). "4: Heavy compounds (alkyltetralins, tricyclic compounds, etc.)
the A1203 support. NiW/USY produced a larger amount of 343 ~ fraction than Mo/DTC and NiMo/A1203 in spite of the lowest asphaltene conversion. In addition, NiW/USY gave a large amount of gases such as C3 and C4 hydrocarbons. The above results indicate that the HCK function of the catalysts was essential for the production of middle distillates from heavy oils at 410~ At the same time, however, a primary problem of NiW/USY was found to be the insufficient hydrogenation activity evidenced by the low asphaltene conversion. Table 2 more clearly elucidates this problem. The product distribution over NiW/USY was not very different from that over USY without NiW sulfide and both of the catalysts yielded more than 50% of heavy compounds that were produced by retrogressive reactions. This was very likely due to poor hydrogen supply from NiW sulfide to the acid sites on USY. Another problem of NiW/USY would be that the active sites for heavy molecules were limited to the external surface of USY zeolites, because the micropore diameter of Y-zeolites is 0.74 nm that is smaller t h a n the molecular sizes of asphaltenes. As shown in Fig. 1, the loading of NiW sulfide almost completely eliminated the meso-pores which USY zeolites originally possessed. To clearly elucidate the effects of the acidity and porosity of the catalysts on the catalytic performances, the reactions were carried out over the three NiMo catalysts supported on HY zeolite. As compared with NiMo/USY, the NiMo/HY(L) showed a poor performance for asphaltene conversion, although both of the catalysts exhibited similar production of light fractions (Table 1). In the previous paper, the authors reported that the balance between HYD and HCK activities should be properly controlled to activate the HCK function of zeolite [3]. The present results further indicate that the dense distribution of acid sites on NiMo/HY(L) is the cause for the retrogressive reactions. NiMo/HY(L) with high mesoporosity gave a higher yield of C5-343 ~ fraction than NiMo/HY(S) with low mesoporosity in the reaction of AH-AR (Table 1). On
430
0.3 .
.
.
',~--~~ ~ i--,~l--+-: NiW/USY -7~-'-~, ,,,~'-.... ! ~-'-'~'---~1 USY I ''" .
0.2 0 . 1
. . . . .
.
.
~
_
_
.-i F ~ - I - -
~
-
-
4-1-
I
101
iii ' .......
" ~
I- -
.
i-- --f - i -
.
~
.
t--~..~
~- -4 - ~ - . ~ - ~ -
t
I
~
I
i ......
.
I-t--t-t--
--
--
-
-
- , - - , - - , - , - , - , - , , , - - -
--~-..1-
t
~- -
-
I
10 2
-~ -
-
~-~t-
-~ - ~ -
I
t"
~- ~ - ~ -
t " r ' ~ t -,, ..4.
10 3
Pore diameter (A) Fig. 1. Pore size distribution of NiW/USY and USY the contrary, the two catalysts exhibited a sfinilar performance in the reaction of tetralin (Table 2). The different feature between the two reactions was evidently due to the different molecular sizes between tetralin and the heavy fractions in the AH-AR. It should be noted that NiMo/HY(S) and NiMo/HY(L) yielded an almost similar amount of gaseous products (Table 1). This indicates that the gaseous products were mainly produced inside the micropores. In conclusion, the present study has confirmed the importance of i) mesoporosity of zeolites, ii) distribution of the acid sites, and iii) hydrogen supply from sulfide to the acid sites. The mesopores play an essential role in the conversion of 565 ~ § however, insufficient supply of hydrogen to the acid sites inside the mesopores rather results in the/ormation oYaspha]tene. The loading of active sulfide species on zeolites provides HYD functions to zeolites, but conventional loading of the sulfide significantly reduces the mesoporosity as shown in Fig. 1. The mesoporosity of HY zeolites are formed during the ionexchange and dealumination procedures that at the same time control the distribution of the acid sites in zeolites. Furhter, these procedures likely result in the destruction of zeolite framework. Thus, the developments of catalyst preparation methods, such as modification of zeolite and loading of active sulfides on zeolites are needed to improve the catalytic activities of bifunctional catalysts. 4. Acknowledgment This work has been carried out as a research project of the Petroleum Energy Center with the subsidy of the Ministry of International Trade and Industry.
5. References 1 I.E. Maxwell, Catal. Today, 1 (1987) 385. 2 K. Sato, Y. Iwata, T. Yoneda, A. Nishijima, Y. Miki and H. Shimada, Catalysis Today, 45, (1998) 367. 3 H. Shimada, S. Yoshitomi, T. Sato, N. Matsubayashi, M. I m a m u r a , Y. Yoshimura and A. Nishijima, Studies in Surf. Sci. Catal., 106 (1997) 115.
Hydrotreatmentand Hydrocrackingof Oil Fractions B. Delmon,G.F.Fromentand P. Grange(Editors) 91999ElsevierScienceB.V.All rightsreserved.
431
Hydrodesulphurization of Residue-Oil Over Ni-Mo/HY-Zeolite catalyst S. Bhatia, J.K. Heng and A.R. Mohamed School of Chemical Engineering, Universiti Sains Malaysia, Perak Branch Campus, Seri Iskandar, 31750 Tronoh, Perak, Malaysia
Abstract Hydrogenation of atmospheric residue-oil containing 1.32 wt % sulphur was carried out in the presence of sulphided nickel molybdenum on HY-Zeolite (Ni-Mo/HY-Zeolite) catalyst. The study was carried out using a 300 ml high pressure stirred batch reactor at 340~ and 8.4 MPa operating pressure. HY-Zeolite supported Ni-Mo catalyst was prepared by double impregnation of nickel nitrate and ammonium molybdate or by ion exchange method and followed by calcination and sulphidation. The liquid product obtained was analyzed for its viscosity, H/C atomic ratio, weight % of sulphur and nitrogen in order to interpret the catalytic effect on its visbreaking, product selectivity, hydrocracking, hydrodesulphurization (HDS) and hydrodenitrogenation (HDN) activities. The Ni-Mo/HY-Zeolite catalysts prepared by ion exchange showed higher HDS, HDN and hydrogenation selectivity than those prepared by impregnation. 1. INTRODUCTION Petroleum residues upgrading through the catalytic hydrogen addition route is becoming increasingly important in modern petroleum refining schemes. Extensive research studies have been devoted in developing the usage of new support materials, incorporation of additional promoters, modifications and optimization of preparation procedures and increase in metal loading [ 1]. In the present work the performance of a Ni-Mo/HY-Zeolite catalyst in residue-oil hydroprocessing is reported. The results that were monitored included hydrodesulphurization (HDS), hydrodenitrogenation (HDN), atomic ratio of H/C, products distribution and visbreaking activity. 2. EXPERIMENTAL The residue-oil used in this study is low sulphur waxy residue refined by Petronas Petroleum Refinery, Melaka, Malaysia. The properties of the residue are the same as reported earlier [2].
2.1 Catalysts Preparation The Ni-Mo/HY-Zeolite catalysts were prepared by dry impregnation or ion exchange of commercial HY-Zeolite support. The HY-Zeolite support was supplied by Norton Chemical Process Products Corporation, U.S.A. The support was successively impregnated or ion exchanged with ammonium molybdate and nickel nitrate in a constant 0.3 atomic Ni/(Ni + Mo) ratio [3].
432
The catalysts were calcined and reduced before presulphidation. The catalysts were sulphided prior to usage by passing a 5 vol% hydrogen sulphide in hydrogen at 380-415 kPa at 400~
2.2 Experimental set-up and procedure Experiments on residue-oil were conducted in a batch mode. A 300 ml reactor was supplied by M/S Parr Instrument Company, U.S.A, is capable of handling pressure up to 13.9 MPa and a temperature of 350~ A residue-oil sample was mixed with catalyst and placed into the reactor. After purging with nitrogen, the reactor was purged with 2.9 MPa hydrogen gas. The reaction was carried out at 340~ and pressure of 7.7 MPa under constant stirrer speed of 1000 rpm. 2.3 Analysis of products Different analytical procedures were carried out to determine the composition of the feed and products of the hydroprocessing reaction. Viscosity was measured at 50~ using a spindle type viscometer model DV-111 Rheometer supplied by M/S Brookfield Engineering Laboratories, USA. The PE2400 Series II CHNS Element Analyzer supplied by M/S Perkin Elmer Co., USA was used to determine the weight percentage of C, H, N and S in organic samples. ASTM distillation was done to determine the product boiling range distribution. 3. RESULTS AND DISCUSSION
3.1 Comparison of Ni-Mo/HY-Zeolite catalyst prepared under different methods. Ni-Mo/HY-Zeolite catalyst was prepared using 3 different techniques. (i) Ni-Mo/HY-Zeolite by double impregnation (Imp). (ii) Ni/HY-Zeolite by ion exchange (Ionex). (iii) Ni-Mo/HY-Zeolite with Ni by ion exchange and Mo by impregnation (Ionex and Imp). The performance of these catalysts is given in Table 1. The product distribution varies with the method of catalyst preparation as shown in Figure 1. The difference of the products distribution may be due to the difference in the acidity and strength between the ion exchanged and double impregnation catalyst samples. This difference suggests that the introduction of Ni by ion exchange yielded catalysts with a higher acidity and a fraction of a very strong acid sites as compared to those prepared by Ni impregnation. The number of acid sites on the sulphided zeolite support increases linearly with increasing degree of cationic exchange and is same as observed by N. Davidoa [4]. After impregnation with ammonium molybdate, calcination and sulphidation, the acidity of the catalysts did not change. Ion-exchange method resulted in high dispersion loading of an active metal species on zeolite supports. All the catalysts containing both Ni and Mo were better for product selectivity than those containing only Ni as shown in Table 1. Ni/HY-Zeolite catalyst has lower hydrocracking of residual oil to products having less than 250~ boiling point, with only 56.5 vol.% as compared to both Ni-Mo/HY-Zeolite with more then 60.0 vol.%. This lower hydrogenation selectivity of the Ni/HY-Zeolite catalyst was due to absence of Mo since hydrogenation activity is generally related to Mo dispersion and the molybdemma sulphide is more active than nickel sulphide for hydrogenation. The H/C atomic ratio of the hydrogenated oil using Ni-Mo/HY-Zeolite catalyst (Ionex and Imp) is much higher compared to Ni-Mo/HY-Zeolite (Imp) or Ni/HY-Zeolite (Ionex) as
433
presented in Figure 1. This showed that Ni-Mo/HY-Zeolite catalyst which prepared by using ion exchange and followed by impregnation method, promoted better hydrodearomatization activity compared to the catalyst which was prepared by impregnation alone. The increment of H/C atomic ratio for Ni-Mo/HY-Zeolite (Ionex and Imp) is about 23.8%. Table 1 shows the viscosities of the hydrogenated oil by using catalysts with different methods of preparations. The viscosity was 61.9 cP of hydrogenated oil using Ni-Mo/HY-Zeolite (Ionex and Imp) catalyst is much lower compared to 68.1 cP using Ni-Mo/HY-Zeolite (Double Impregnation) or 70.0 cP using Ni/HY-Zeolite (Ionex) catalyst. This showed that Ni-Mo/HYZeolite (Ionex and Imp) catalyst promoted better visbreaking compared to Ni-Mo/HY-Zeolite (Imp) and Ni/HY-Zeolite (Ionex). As shown in Figure 1, the activity for HDS and HDN of residue-oil was significantly affected by the different methods of preparations of the catalysts. The Ni-Mo/HY-Zeolite (Ionex and Imp) catalyst exhibited activity superior to that of Ni-Mo/HY-Zeolite (Imp) or Ni/HY-Zeolite (Ionex). While Ni-Mo/HY-Zeolite (Imp) showed high HDS activity, its HDN activity was poor with only 23.8% HDN compared to Ni-Mo/HY-Zeolite (Ionex and Imp) with 28.4% HDN and Ni/HY-Zeolite (Ionex) with 25.1% HDN.
Table 1Performance of different Ni-Mo/HY-Zeolite catalysts Catalyst
Residue Oil Without catalyst
NiNiNi/HY NiNiNiM o / H Y Mo/HY Zeolite Mo/HY M o / H Y Mo/HY Zeolite -Zeolite (Ionex) -Zeolite -Zeolite -Zeolite (Imp) (Ionex (Ionex (Ionex (Ionex & Imp) & Imp) & Imp) & Imp) 4 4 4 5 5 3 340 340 340 300 340 340
Run Time(hrs) 4 Run Temperature 340 (~ Sulphur Content in 1.32 1.32 1.32 1.32 1.32 1.32 Residue (wt%S) Product distribution (vol. %) 1st Drop (~ 198 82 88 92 92 86 <170~ (gasoline) 33.0 31.0 25.0 26.0 28.5 <190~ (Kerosene) 0.3 8.0 10.0 8.5 10.0 11.5 <220~ (Jet Fuel) 0.7 11.0 14.0 11.0 12.5 14.5 <250~ (Diesel) 2.0 10.0 12.0 12.0 10.0 15.0 >250~ (Gasoil) 97.0 38.0 33.0 43.5 41.5 30.5 Viscosity (cP) at 50~ 60 rpm 140.0 68.1 61.9 70.0 68.1 52.0 Elemental Composition (wt%) C% 84.91 82.24 79.98 8 0 . 1 6 80.75 78.65 H% 11.33 12.68 13.20 12.36 13.05 13.50 N% 2.39 1.82 1.71 1.79 1.66 1.58 S% 1.32 0.55 0.38 0.56 0.41 0.34 H/C ratio 1.60 1.85 1.98 1.85 1.94 2.06 HDN and HDS activity (%) HDN % 23.8 28.4 25.1 30.5 33.9 HDS % 58.3 71.2 57.6 68.9 74.2 H2 charge Pressure = 2.9 MPa, Residue-oil Loading = 80g, Catalyst loading
Residue NiOil Mo/HY Without -Zeolite catalyst (Ionex & Imp) 4 4 340 340
1.32
2.00
2.00
90 20.0 12.0 16.0 12.0 40.0
178 10.0 2.0 2.5 3.0 82.5
76 28.0 9.0 12.0 11.5 39.5
60.5
138.0
66.0
80.11 12.76 1.83 0.45 1.91
84.39 11.63 1.98 2.00 1.65
78.79 12.67 1.43 0.99 1.93
23.4 65.9 = 4g
27.5 50.5
434
Figure 1. Comparison of Product's selectivity and hydrogenation activity over Ni-Mo/HYZeolite catalysts. 4. CONCLUSIONS 1. Zeolite HY-supported nickel-molybdenum (Ni-Mo/HY-Zeolite) sulphided catalysts showed large differences in catalytic behavior and is dependent on the preparation method (Impregnation vs. Ion exchange). The catalysts prepared by ion exchange showed higher hydrogenation selectivity, HDS and HDN activities than those prepared by impregnation. The binary Ni-Mo catalysts were more active than the individual Ni catalysts. 2. Higher sulphur content from 1.32 wt% S to 2.00 wt% S showed lower HDS, HDN and hydrocracking of residue-oil to fractions with boiling point less than 250~ 5. REFERENCES 1. M. Absi-Halabi, A. Stanislaus and K. A1-Dolama, Fuel, Vol. 77, No 7, (1998), 787-790. 2. P.T. Koh, A.R. Mohamed and S. Bhatia, Fuel, Vol. 77, No 11, (1998), 1221-1227. 3. C.E. Ramos-Galvan, G. Sandoval-Robles, A. Castillo-Mares, J.M. Dominguez, Applied Catalysis A, 150, (1997), 37-52. 4. N. Davidoa, P. Kovacheva, and D. Shopov, Studies in Surface Science and Catalysis, Vol 24, Elsevier Science B.V., (1985), 659.
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377 235, 343,347 351 343, 347 69, 427 401 153 361 219 365 235, 347 431 361 413 85 327 227 413 289 421 369 77, 97 243 243 401 327 351 401 351 77, 85 385 397 401 137 361 51 357 21 203, 365 85, 219 51 261 219 361 351 37 385
436 Geantet C. Giraldo S.A. Goldfarb Y.Y. Goldwasser J. Gonzalez-Jimenez F. Grange P. Griboval A. Grimblot J. Grudoski D. Grzechowiak J.R. Guirardello R. Ha B.-H. Hafner J. Hagan A.P. Halachev T. Harl~ V. Heng J.K. Herskowitz M. Ho T.C. Honma T. Honna K. Hubaut R. H utsch ka F. Iwamoto R. Iwata Y. Jouguet B. Kasztelan S. Kitaev L.Ye. Klemt A. Klimova T. Koch H. Koltai T. Komulainen S. Koyama H. Krause A.O.I. Kresse G. Kubasov A.A. Landau M.V. Le Bihan L. Leclercq C. Lee K.-H. Leglise J. Letourneur D. LiD. Li M. L6pez Cordero R. L6pez-Agudo A. Luis M.A. Maggi R. Maity S.K. Maloletnev A.S. Marroquin-S~nchez G. Matsubayashi N.
227 97 377 365 203, 365 85 361 105, 169, 357 187 337 289 243 309, 327 3 401 153,261 431 393 179 161 69, 427 203, 365 327 169 69 153 261 377 405 373 405 137 145 195 145 309 377 393 105 153 243 51 153 409 409 369 369, 3 73 203 77, 85 397 381 235, 343 69, 427
437 Mauchauss~ C. Maug6 F. Merino L. Miki Y. Minderhoud J.K. Moch ida I. Mohamed A.R. Moreno T.A. Moulijn J.A. Mujic E. Nagai M. Narasimhan C.S.L. Nie H. Ohta Y. Omi S. Paul J.F. Payen E. Peeters E. P6rez A.A. Petranovskii V. Pinz6n M.H. Pizzio L. Porta J.-A. Prasada Rao T.S.R. Prins R. Quiroz-Sosa G. Ramfrez de Agudelo M.M. Ramirez J. Rana M.S. Raybaud P. Reinhoudt H.R. Reschetilowski W. Riva R. Rives A. Roe C.L. Rota F. Rynkowski J. Sakamoto S. Sakanishi K. Salazar J.A. Salazar-Sotelo D. Sato K. Satoh S. Sau M. Schulz K.H. Scott C.E. Segawa A. Segawa K. Selvam T. Shi Y. Shibata Y. Shido T. Shimada H.
105 269 97 69, 427 3 187 431 235 211,251 279 195 297 409 161 195 327 105, 327, 361 227 235 351 97 413 385 397 113, 319, 421 347 279 369, 373 397 309 211,251 405 219 203, 365 121 319 337 195 187 279 343, 347 427 129 297 121 203, 365 389 129 145 409 389 421 69, 427
438 Shimizu T. Shin S. Shinn J. Sie S.T. Slivinsky Ye.V. Solis D. Srinivas B.N. Sun M. Suvorov U.P. Taouli A. T~t6nyi P. Toulhoat H. Travert A. Troost R. van Gestel J.N.M. van Gorsel M. van Langeveld A.D. van Veen J.A.R. V~zquez P. Verma R.P. Viljava T.-R. Vradman L. Vrinat M. Watanabe K. Wereszczako-Ziel inska I. Yamada M. Yitzhaki D. Yoneda T. Zanibelli L. Zotin J.L.
161 187 187 211,251 377 373 397 113 381 4O5 137 3O9 269 251 51 211 211,251 3,211,251 413 297 145 393 153,227, 369 389 337 161 393 389 219 227
439 S T U D I E S IN SURFACE SCIENCE A N D CATALYSIS Advisory Editors: B. Delmon, Universite Catholique de Louvain, Louvain-la-Neuve, Belgium J.T. Yates, University of Pittsburgh, Pittsburgh, PA, U.S.A. Volume
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Preparation of Catalysts I.Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the First International Symposium, Brussels, October 14-17,1975 edited by B. Delmon, P.A. Jacobs and G. Poncelet The Control of the Reactivity of Solids. A Critical Survey of the Factors that Influence the Reactivity of Solids, with Special Emphasis on the Control of the Chemical Processes in Relation to Practical Applications by V.V. Boldyrev, M. Bulens and B. Delmon Preparation of Catalysts I1. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings ofthe Second International Symposium, Louvain-la-Neuve, September 4-7, 1978 edited by B. Delmon, P. Grange, P.Jacobs and G. Poncelet Growth and Properties of Metal Clusters. Applications to Catalysis and the Photographic Process. Proceedings ofthe 32nd International Meeting ofthe Societe de Chimie Physique, Villeurbanne, September 24-28, 1979 edited by J. Bourdon Catalysis by Zeolites. Proceedings of an International Symposium, Ecully (Lyon), September 9-11, 1980 edited by B. Imelik, C. Naccache, Y. Ben Taarit, J.C. Vedrine, G. Coudurier and H. Praliaud Catalyst Deactivation. Proceedings of an International Symposium, Antwerp, October 13- 15,1980 edited by B. Delmon and G.E Froment New Horizons in Catalysis. Proceedings of the 7th International Congress on Catalysis, Tokyo, June 30-July4, 1980. Parts A and B edited by T. Seiyama and K. Tanabe Catalysis by Supported Complexes by Yu.l. Yermakov, B.N. Kuznetsov and V.A. Zakharov Physics of Solid Surfaces. Proceedings of a Symposium, Bechyhe, September 29-October 3,1980 edited by M. Lazni~ka Adsorption at the Gas-Solid and Liquid-Solid Interface. Proceedings of an International Symposium, Aix-en-Provence, September 21-23, 1981 edited by J. Rouquerol and K.S.W. Sing Metal-Support and Metal-Additive Effects in Catalysis. Proceedings of an International Symposium, Ecully (Lyon), September 14-16, 1982 edited by B. Imelik, C. Naccache, G. Coudurier, H. Praliaud, P. Meriaudeau, P. Gallezot, G.A. Martin and J.C. Vedrine Metal Microstructures in Zeolites. Preparation - Properties- Applications. Proceedings of a Workshop, Bremen, September 22-24, 1982 edited by P.A. Jacobs, N.I. Jaeger, P.Jin3 and G. Schulz-Ekloff Adsorption on Metal Surfaces. An Integrated Approach edited by J. Benard Vibrations at Surfaces. Proceedings of the Third International Conference, Asilomar, CA, September 1-4, 1982 edited by C.R. Brundle and H. Morawitz Heterogeneous Catalytic Reactions Involving Molecular Oxygen by G.I. Golodets
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Methane Conversion. Proceedings of a Symposium on the Production of Fuels and Chemicals from Natural Gas, Auckland, April 27-30, 1987 edited by D.M. Bibby, C.D. Chang, R.F. Howe and S. Yurchak Innovation in Zeolite Materials Science. Proceedings of an International Symposium, Nieuwpoort, September 13-17, 1987 edited by P.J. Grobet, W.J. Mortier, E.F. Vansant and G. Schulz-Ekloff Catalysis 1987. Proceedings ofthe 10th North American Meeting ofthe Catalysis Society, San Diego, CA, May 17-22, 1987 edited by J.W. Ward Characterization of Porous Solids. Proceedings of the IUPAC Symposium (COPS I), Bad Soden a. Ts., April 26-29,1987 edited by K.K. Unger, J. Rouquerol, K.S.W. Sing and H. Kral Physics of Solid Surfaces 1987. Proceedings of the Fourth Symposium on Surface Physics, Bechyne Castle, September 7-11, 1987 edited by J. Koukal Heterogeneous Catalysis and Fine Chemicals. Proceedings of an International Symposium, Poitiers, March 15-17, 1988 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, C. Montassier and G. Perot Laboratory Studies of Heterogeneous Catalytic Processes by E.G. Christoffel, revised and edited by Z. Paal Catalytic Processes under Unsteady-State Conditions by Yu. Sh. Matros Successful Design of Catalysts. Future Requirements and Development. Proceedings ofthe Worldwide Catalysis Seminars, July, 1988, on the Occasion of the 30th Anniversary of the Catalysis Society of Japan edited by T. Inui Transition Metal Oxides. Surface Chemistry and Catalysis by H.H. Kung Zeolites as Catalysts, Sorbents and Detergent Builders. Applications and Innovations. Proceedings of an International Symposium, Wfirzburg, September 4-8,1988 edited by H.G. Karge and J. Weitkamp Photochemistry on Solid Surfaces edited by M. Anpo and T. Matsuura Structure and Reactivity of Surfaces. Proceedings of a European Conference, Trieste, September 13-16, 1988 edited by C. Morterra, A. Zecchina and G. Costa Zeolites: Facts, Figures, Future. Proceedings of the 8th International Zeolite C°nference'Amsterdam'July 10-14, 1989. Parts A and B edited by P.A. Jacobs and R.A. van Santen Hydrotreating Catalysts. Preparation, Characterization and Performance. Proceedings of the Annual International AIChE Meeting, Washington, DC, November 27-December 2, 1988 edited by M.L. Occelli and R.G. Anthony New Solid Acids and Bases. Their Catalytic Properties by K. Tanabe, M. Misono, Y. Ono and H. Hattori Recent Advances in Zeolite Science. Proceedings of the 1989 Meeting of the British Zeolite Association, Cambridge, April 17-19, 1989 edited by J. Klinowsky and P.J. Barrie Catalyst in Petroleum Refining 1989. Proceedings of the First International Conference on Catalysts in Petroleum Refining, Kuwait, March 5-8, 1989 edited by D.L. Trimm, S. Akashah, M. Absi-Halabi and A. Bishara Future Opportunities in Catalytic and Separation Technology edited by M. Misono, Y. Moro-oka and S. Kimura
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New Developments in Selective Oxidation. Proceedings of an International Symposium, Rimini, Italy, September 18-22, 1989 edited by G. Centi and F. Trifiro Volume 56 Olefin Polymerization Catalysts. Proceedings of the International Symposium on Recent Developments in Olefin Polymerization Catalysts, Tokyo, October 23-25, 1989 edited by T. Keii and K. Soga Volume 57A Spectroscopic Analysis of Heterogeneous Catalysts. Part A: Methods of Surface Analysis edited by J.L.G. Fierro Volume 57B Spectroscopic Analysis of Heterogeneous Catalysts. Part B: Chemisorption of Probe Molecules edited by J.L.G. Fierro Volume 58 Introduction to Zeolite Science and Practice edited by H. van Bekkum, E.M. Flanigen and J.C. Jansen Volume 59 Heterogeneous Catalysis and Fine Chemicals I1. Proceedings of the 2nd International Symposium, Poitiers, October 2-6, 1990 edited by M. Guisnet, J. Barrault, C. Bouchoule, D. Duprez, G. Perot, R. Maurel and C. Montassier Volume 60 Chemistry of Microporous Crystals. Proceedings of the International Symposium on Chemistry of Microporous Crystals, Tokyo, June 26-29, 1990 edited by T. Inui, S. Namba and T. Tatsumi Volume 61 Natural Gas Conversion. Proceedings of the Symposium on Natural Gas Conversion, Oslo, August 12-17, 1990 edited by A. Holmen, K.-J. Jens and S. Kolboe Volume 62 Characterization of Porous Solids I1. Proceedings of the IUPAC Symposium (COPS II), Alicante, May 6-9, 1990 edited by F. Rodriguez-Reinoso, J. Rouquerol, K.S.W. Sing and K.K. Unger Volume 63 Preparation of Catalysts V. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Fifth International Symposium, Louvain-la-Neuve, September 3-6, 1990 edited by G. Poncelet, P.A. Jacobs, P.Grange and B. Delmon New Trends in CO Activation Volume 64 edited by L. Guczi Catalysis and Adsorption by Zeolites. Proceedings of ZEOCAT 90, Leipzig, Volume 65 August 20-23, 1990 edited by G. Ohlmann, H. Pfeifer and R. Fricke Dioxygen Activation and Homogeneous Catalytic Oxidation. Proceedings of the Volume 66 Fourth International Symposium on Dioxygen Activation and Homogeneous Catalytic Oxidation, Balatonf~red, September 10-14, 1990 edited by L.I. Simandi Structure-Activity and Selectivity Relationships in Heterogeneous Catalysis. Volume 67 Proceedings of the ACS Symposium on Structure-Activity Relationships in Heterogeneous Catalysis, Boston, MA, April 22-27, 1990 edited by R.K. Grasselli and A.W. Sleight Catalyst Deactivation 1991. Proceedings of the Fifth International Symposium, Volume 68 Evanston, IL, June 24-26, 1991 edited by C.H. Bartholomew and J.B. Butt Zeolite Chemistry and Catalysis. Proceedings of an International Symposium, Volume 69 Prague, Czechoslovakia, September 8-13, 1991 edited by P.A. Jacobs, N.I. Jaeger, L. Kubelkova and B. Wichterlova Poisoning and Promotion in Catalysis based on Surface Science Concepts and Volume 70 Experiments by M. Kiskinova
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Catalyst Deactivation 1994. Proceedings of the 6th International Symposium, Ostend, Belgium, October 3-5, 1994 edited by B. Delmon and G.F. Froment Catalyst Design for Tailor-made Polyolefins. Proceedings of the International Symposium on Catalyst Design for Tailor-made Polyolefins, Kanazawa, Japan, March 10-12, 1994 edited by K. Soga and M. Terano Acid-Base Catalysis I1. Proceedings of the International Symposium on Acid-Base Catalysis II, Sapporo, Japan, December 2-4, 1993 edited by H. Hattori, M. Misono and Y. Ono Preparation of Catalysts VI. Scientific Bases for the Preparation of Heterogeneous Catalysts. Proceedings of the Sixth International Symposium, Louvain-La-Neuve, September 5-8, 1994 edited by G. Poncelet, J. Martens, B. Delmon, P.A. Jacobs and P. Grange Science and Technology in Catalysis 1994. Proceedings of the Second Tokyo Conference on Advanced Catalytic Science and Technology, Tokyo, August 21-26, 1994 edited by Y. Izumi, H. Arai and M. Iwamoto Characterization and Chemical Modification of the Silica Surface by E.F.Vansant, P.Van Der Voort and K.C. Vrancken Catalysis by Microporous Materials. Proceedings of ZEOCAT'95, Szombathely, Hungary, July 9-13, 1995 edited by H.K. Beyer, H.G.Karge, I. Kiricsi and J.B. Nagy Catalysis by Metals and Alloys by V. Ponec and G.C. Bond Catalysis and Automotiye Pollution Control II1. Proceedings of the Third International Symposium (CAPoC3), Brussels, Belgium, April 20-22, 1994 edited by A. Frennet and J.-M. Bastin Zeolites: A Refined Tool for Designing Catalytic Sites. Proceedings of the International Symposium, Qu6bec, Canada, October 15-20, 1995 edited by L. Bonneviot and S. Kaliaguine Zeolite Science 1994: Recent Progress and Discussions. Supplementary Materials to the 10th International Zeolite Conference, Garmisch-Partenkirchen, Germany, July 17-22, 1994 edited by H.G. Karge and J. Weitkamp Adsorption on New and Modified Inorganic Sorbents edited by A. Da,browski and V.A. Tertykh Catalysts in Petroleum Refining and Petrochemical Industries 1995. Proceedings of the 2nd International Conference on Catalysts in Petroleum Refining and Petrochemical Industries, Kuwait, April 22-26, 1995 edited by M. Absi-Halabi, J. Beshara, H. Qabazard and A. Stanislaus 1lth International Congress on Catalysis - 40th Anniversary. Proceedings ofthe 1lth ICC, Baltimore, MD, USA, June 30-July 5, 1996 edited by J. W. Hightower, W.N. Delgass, E. Iglesia and A.T. Bell Recent Advances and New Horizons in Zeolite Science and Technology edited by H. Chon, S.I. Woo and S. -E. Park Semiconductor Nanoclusters - Physical, Chemical, and Catalytic Aspects edited by P.V. Kamat and D. Meisel Equilibria and Dynamics of Gas Adsorption on Heterogeneous Solid Surfaces edited by W. Rudzi~ski, W.A. Steele and G. Zgrablich Progress in Zeolite and Microporous Materials Proceedings ofthe 1lth International Zeolite Conference, Seoul, Korea, August 12-17, 1996 edited by H. Chon, S.-K. Ihm and Y.S. Uh
445 Hydrotreatment and Hydrocracking of Oil Fractions Proceedings ofthe 1st International Symposium / 6th European Workshop, Oostende, Belgium, February 17-19, 1997 edited by G.F. Froment, B. Delmon and P. Grange Volume 107 Natural Gas Conversion IV Proceedings of the 4th International Natural Gas Conversion Symposium, Kruger Park, South Africa, November 19-23, 1995 edited by M. de Pontes, R.L. Espinoza, C.P. Nicolaides, J.H. Scholtz and M.S. Scurrell Volume 108 Heterogeneous Catalysis and Fine Chemicals IV Proceedings of the 4th International Symposium on Heterogeneous Catalysis and Fine Chemicals, Basel, Switzerland, September 8-12, 1996 edited by H.U. Blaser, A. Baiker and R. Prins Volume 109 Dynamics of Surfaces and Reaction Kinetics in Heterogeneous Catalysis. Proceedings ofthe International Symposium, Antwerp, Belgium, September 15-17, 1997 edited by G.F. Froment and K.C. Waugh Volume 110 Third World Congress on Oxidation Catalysis. Proceedings ofthe Third World Congress on Oxidation Catalysis, San Diego, CA, U.S.A., 21-26 September 1997 edited by R.K. Grasselli, S.T. Oyama, A.M. Gaffney and J.E. Lyons Volume 111 Catalyst Deactivation 1997. Proceedings ofthe 7th International Symposium, Cancun, Mexico, October 5-8, 1997 edited by C.H. Bartholomew and G.A. Fuentes Volume 112 Spillover and Migration of Surface Species on Catalysts. Proceedings of the 4th International Conference on Spillover, Dalian, China, September 15-18, 1997 edited by Can Li and Qin Xin Volume 113 Recent Advances in Basic and Applied Aspects of Industrial Catalysis. Proceedings ofthe 13th National Symposium and Silver Jubilee Symposium of Catalysis of India, Dehradun, India, April 2-4, 1997 edited by T.S.R. Prasada Rao and G. Murali Dhar Volume 114 Advances in Chemical Conversions for Mitigating Carbon Dioxide. Proceedings of the 4th International Conference on Carbon Dioxide Utilization, Kyoto, Japan, September 7-11, 1997 edited by T. Inui, M. Anpo, K. Izui, S. Yanagida and T. Yamaguchi Volume 115 Methods for Monitoring and Diagnosing the Efficiency of Catalytic Converters. A patent-oriented survey by M. Sideris Volume 116 Catalysis and Automotive Pollution Control IV. Proceedings ofthe 4th International Symposium (CAPoC4), Brussels, Belgium, April 9-11, 1997 edited by N. Kruse, A. Frennet and J.-M. Bastin Volume 117 Mesoporous Molecular Sieves 1998 Proceedings of the 1st International Symposium, Baltimore, MD, U.S.A., July 10-12, 1998 edited by L.Bonneviot, F. Bdland, C. Danurnah, S. Giasson and S. Kaliaguine Volume 118 Preparation of Catalysts VII Proceedings of the 7th International Symposium on Scientific Bases for the Preparation of Heterogeneous Catalysts, Louvain-la-Neuve, Belgium, September 1-4, 1998 edited by B. Delmon, P.A. Jacobs, R. Maggi, J.A. Martens, P. Grange and G. Poncelet Volume 119 Natural Gas Conversion V Proceedings ofthe 5th International Gas Conversion Symposium, Giardini-Naxos, Taormina, Italy, September 20-25, 1998 edited by A. Parmaliana, D. Sanfilippo, F. Frusteri, A. Vaccari and F. Arena Volume 120A Adsorption and its Applications in Industry and Environmental Protection. Vol I: Applications in Industry edited by A. Dabrowski
Volume 106
446 Volume 120B Adsorption and its Applications in Industry and Environmental Protection. Vol I1:Applications in Environmental Protection edited by A. Dabrowski Volume 121 Science and Technology in Catalysis 1998 Proceedingsof the Third Tokyo Conference in Advanced Catalytic Science and Technology, Tokyo, July 19-24, 1998 edited by H. Hattori and K. Otsuka Volume 122 Reaction Kinetics and the Development of Catalytic Processes Proceedings ofthe International Symposium, Brugge, Belgium, April 19-21, 1999 edited by G.E Fremont and K.C. Waugh Volume 123 Catalysis: An Integrated Approach Second, Revised and Enlarged Edition edited by R.A. van Santen, P.W.N.M. van Leeuwen, J.A. Moulijn and B.A. Averill Volume 124 Experiments in Catalytic Reaction Engineering by J.M. Berty Volume 125 Porous Materials in Environmentally Friendly Processes Proceedings ofthe 1st International FEZA Conference, Eger, Hungary, September 1-4, 1999 edited by I. Kiricsi, G. PaI-Borbely, J.B. Nagy and H.G. Karge Volume 126 Catalyst Deactivation 1999 Proceedings of the 8th International Symposium, Brugge, Belgium, October 10-13, 1999 edited by B. Delmon and G.F. Froment Volume 127 Hydrotreatment and Hydrocracking of Oil Fractions Proceedings of the 2nd International Symposium/7th European Workshop, Antwerpen, Belgium, November 14-17, 1999 edited by B. Delmon, G.F. Froment and P. Grange
