Catalysis: A Review of Chemical Literature: Vol 15 (Specialist Periodical Reports)

Catalysis Volume 15

A Specialist Periodical Report

Catalysis Volume 15 A Review of Recent Literature Senior Reporter James J. Spivey, Research Triangle Institute, Research Triangle Park, North Carolina, USA

Reporters F? Aghalayam, University of Massachusetts, Amherst, MA, USA T. Baba, Tokyo Institute of Technology, Japan S. Bordawekar, University of Virginia, Charlottesville, VA, USA B.H. Davis, University of Kentucky, Lexington, KY, USA R.J. Davis, University of Virginia, Charlottesville, VA, USA E.J. Doskocil, University of Virginia, Charlottesville, VA, USA M. Machida, Miyazaki University, Miyazaki, Japan Y. Ono, National lnstitution for Academic Degrees, Yokohama, Japan Y.K. Park, University of Massachusetts, Amherst, MA, USA A. Ueno, Shizuoka University, Shizuoka, Japan D.G. Vlachos, University of Massachusetts, Amherst, MA, USA Y. Zhang, University of Kentucky, Lexington, KY, USA

ROYAL SOCIETY OF CHEMISTRY

ISSN 0 140-0568 ISBN 0-85404-219-9 @] The Royal Society of Chemistry 2000

All rights reserved. Apart from any fair dealing for the purpose of research or private study, or criticism or review us permitted under the terms of the UK Copyright, Designs and Patents Act, 1988, this publication may not be reproduced, stored or trunsmitted, in any form or by m y means, without the prior permission in writing of The Royal Society of Chemistry, or in the cuse of reprogruphic reproduction only in accordance with the terms of the licences issued by the Copyright Licensing Agency in the UK, or in uccordance with the terms of the licences issued by the appropriute Reproduction Rights Organizution outside the UK, Enquiries concerning reproduction outside the terms stuted here should be sent to The Royul Society of Chemistry ut the address printed on this page. Published by The Royal Society of Chemistry, Thomas Graham House, Science Park, Milton Road, Cambridge CB4 OWF, UK For further information see our web site at www.rsc.org Typeset by Computape (Pickering) Ltd, Pickering, North Yorkshire, UK Printed and bound by Athenaeum Press Ltd, Gateshead, Tyne &Wear

Preface

The challenges that the chemical industry and research institutions face are increasing. The cost-efficient and environmentally sound conversion of natural resources into fuels, chemicals, and energy require innovative solutions. Catalysts are essential to these processes, and the subjects of the chapters presented here reflect this. In two separate chapters, Robert Davis, Eric Doskocil, and Shailendra Bordawekar (University of Virginia), and Yoshio Ono (National Institution for Academic Degrees, Japan) and Toshihide Baba (Tokyo Institute of Technology) review catalysis by solid bases. These reviews reflect the interest of catalyst researchers in this area. Although acid catalysis is much more widely studied and applied industrially, there are key reactions that require base catalysts, such as aldol condensations. In some cases, the synthesis of heterogeneous base catalysts to replace aqueous base catalysts can have positive environmental effects by eliminating wastewater streams that are otherwise produced. The challenge is to maintain industrially practical selectivity and activity without significant deactivation. Masato Machida (Miyazaki University) has provided a review of solid sorbents for No, removal. These materials can be used in conjunction with catalytic reduction, especially when the NO, concentrations are low. Carbons, zeolites, and modified aluminas can be used - both in stationary and mobile sources. He shows how the NO, separation and catalytic reduction can be combined into a cyclical process for NO, control. P. Aghalayam, Y.K. Park, and D.G. Vlachos (University of Massachusetts) review progress in catalysts for the partial oxidation of light alkanes. Most of the interest in this area, and the focus of this review, is on the conversion of methane. However, the authors have expanded their review to oxidative dehydrogenation of ethane and other alkanes. In an closely related chapter, Akifumni Ueon (Shizuoka University) describes recent work in the partial oxidation of methane, with a special focus on silocomolybdic acid catalysts. These solid acids are one of several classes of catalysts being investigated for this process, and they are more selective for oxygenates such as formaldehyde and methanol - both industrially important intermediates now produced by other processes. Finally, Burt Davis and Yongquing Zhang (University of Kentucky) have reviewed the status of indirect liquefaction, and the critical role that catalyst development must play in bringing these processes to wider commercial practice. This review clarifies the relationships among the wide range of V

vi

Preface

processes that have been developed and offers insight into the challenges that remain. Volume 16 is underway and I look forward to bringing it to you. As always, comments are welcome. James J. Spivey Research Triangle Institute Research Triangle Park, NC USA Email:[email protected]

Contents

Chapter 1 Strong Solid Bases for Organic Reactions By Yoshio On0 and Toshihide Baba 1 Introduction

1

1

2 Role of Solid Base and Basic Sites as a Catalyst 2.1 Abstraction of Protons 2.2 Activation of Reactants without Proton Abstraction 2.3 Cooperative Action of Acidic and Basic Sites

2 2

2 2

3 Base Strength of Basic Sites 3.1 H - Acidity Function 3.2 Indicator Method 3.3 Other Methods for Determining Basic Strength

4 Base Strength and Catalytic Reactions

5

5 Solid Base Materials 5.1 Alkaline Earth Oxides 5.2 Alkali Metals Supported on Metal Oxides 5.3 Alkali Metals on A1203Treated with Alkali Metal Hydroxides 5.4 KNH2/A1203 5.5 Alkali Metal Compound Supported on Alumina 5.5.1 KNOJA1203, KHCOJA1203, K2C03/A1203 5.5.2 LiOH/A1203, NaOH/A1203,KOH/A1203, RbOH/A1203, CsOH/A1203 5.6 Zeolites 5.7 Modified Zeolites 5.8 E~,Yb/A1203 5.9 KF/A1203 5.10 Mixed Oxides

8 8

14 14 15 16 16 18

6 Reactions Catalysed by Solid Bases

20

~ _ _ _

Catalysis, Volume 15 0The Royal Society of Chemistry, 2000

vii

9 10 10 14 14

...

Contents

Vlll

6.1 Isomerizations 6.1.1 Isomerization of Alkenes 6.1.2 Isomerization of Alkenyl Amines 6.1.3 Isomerization of Ally1 Ethers 6.1.4 Isomerization of Alk- 1-ynes 6.2 Nucleophilic Addition Reactions of Phenylacetylene 6.2.1 Dimerization of Phenylacetylene 6.2.2 Reactions of Alk-1-yne with Ketone or Aldehyde 6.3 Aldol-type Condensations 6.3.1 Aldol Condensations 6.3.2 Claisen-Schmidt Reactions 6.3.3 Knoevenagel Condensations 6.3.4 Michael Additions 6.3.5 Dehydrocondensation with Methanol 6.4 Nucleophilic Substitution at Silicon - Si-C Bond Formation 6.4.1 Metathesis of Trimethylsilylacetylene 6.4.2 Cross-metathesis of Alkynes 6.4.3 Reaction of Alk-1-yne with Silane 6.4.4 Reaction of Alk-1-ene with Diethylsilane 6.4.5 Reaction of Toluene with Diethylsilane 6.5 Miscellaneous Reactions 6.5.1 Tishchenko Reaction 6.5.2 Reactions of Silanes with Ketones or Aldehydes 6.5.3 Reaction of Aniline with Trimethylsilylacetylene 6.5.4 Ring Opening of Epoxides with Me3SiCN 6.5.5 Hydrocyanation 6.5.6 Alkylations 6.5.7 Formation and Ring Transformations of Heterocycles

20 20 20 21 21 21 21

7

Concluding Remarks - Unsolved Problems

35

References

36

Chapter 2 Catalysis by Solid Bases By Eric J. Doskocil, Shailendra Bordawekar and Robert J. Davis

22 25 25 25 27 28 28 28 29 30 30 31 31 32 32 33 33 34 34 34 35

40

1 Introduction

40

2 Types of Basic Catalysts 2.1 Supported Alkali Metal Oxides

41 41

Contents

ix

2.2 2.3 2.4 2.5

Hydrotalcites Zeolites and Mesoporous Oxides Alkali Metals Supported on Nanoporous Carbon Basic Phosphate Oxynitride Catalysts

3 Characterization of Solid Base Catalysts 3.1 Theoretical Ranking of Solid Basicity 3. I. 1 Sanderson Intermediate Electronegativity 3.1.2 Molecular Orbital Calculations 3.2 Experimental Ranking of Solid Basicity 3.2.1 Infrared Spectroscopy of Adsorbed Probe Molecules 3.2.2 Visible Absorption Spectroscopy of Adsorbed Iodine 3.2.3 Microcalorimetry of Adsorbed Probe Molecules 3.2.4 Temperature Programmed Desorption (TPD) of Probe Molecules 3.2.5 X-Ray Absorption Spectroscopy 3.2.6 X-Ray Photoelectron Spectroscopy

41 42 44 45 45 46 46 46 47 47 50 52 54 54 56

4 Probe Reactions over Solid Base Catalysts 57 4.1 Double-bond Isomerization 58 59 4.2 Alcohol Dehydrogenation 4.3 Hydrogenation Reactions 60 4.4 Condensation Reactions 61 4.4.1 Aldol Condensation 61 4.4.2 Knoevenagel Condensation 62 4.5 Alkylation Reactions 63 4.6 Side-chain Alkenylation of o-Xylene with 1,3-Butadiene 64 4.7 Miscellaneous Reactions 65

5 Conclusions

67

Acknowledgements

67

References

67

Chapter 3 Solid Sorbents for Catalytic NO, Removal By Masato Machida

73

1 Introduction

73

2 Materials for NO, Sorption 2.1 Carbonaceous Solids 2.2 Zeolites

74 74

75

Conten ts

X

2.3 2.4 2.5 2.6 2.7 2.8

Alumina Alkaline Solids Transition Metal Oxides containing Alkaline Earths NO, Intercalation Compounds Non-alkaline Solid Oxides Heteropoly Compounds

76 77 78 82 83 87

3 Regeneration of NO, Sorbent 3.1 Pressure Swing Process 3.2 Thermal (Temperature) Swing Process 3.3 Reduction-Oxidation Cycles

88 89 90 90

4 Practical Applications of NO, Sorption 4.1 Flue Gas Treatment for Stationary Sources 4.2 Automobile NO, Storage Catalysts

90 90 92

5 Conclusion

93

References Chapter 4 Partial Oxidation of Light Alkanes in Short Contact Time Microreactors By P. Aghalayam, Y.K. Park and D. G. Vlachos

1 Introduction

93

98

98

2 CH4 Partial Oxidation on Platinum and Rhodium Catalysts 2.1 Unsupported Catalysts 2.2 Foam and Extruded Monoliths 2.3 Fluidized Beds 2.4 The Effect of the Support and Pore Diameter

99 101 103 103 104

3 Influence of Operating Conditions 3.1 Temperature 3.2 Flow Velocity and Contact Time 3.3 Pressure 3.4 Dilution 3.5 Inlet Fuel Composition 3.6 Influence of Different Surface Coverages

104 104 104 107 108 108 109

4 Analysis of the Effect of Operating Conditions 4.1 Model 4.2 Flow Velocity 4.3 Temperature

110 110 112 113

xi

Contents

5 Bifurcation of Methane-Oxygen Mixtures Near Catalysts 5.1 Influence of Fuel Composition 5.2 Influence of Pressure and Flow Velocity 5.3 The Inhibiting Role of the Catalyst in Gas-phase Ignition 5.4 Importance of Gas-phase Reactions

118 119

6 Direct versus Indirect Path to Syngas Formation

121

114 114 118

7 A Quantitative Reaction Mechanism for Oxidation of Methane 7.1 Currently Proposed Mechanisms 7.1.1 Selectivity Mechanisms 7.1.2 Catalytic Ingition Mechanisms 7.1.3 Other Proposed Reaction Pathways 7.2 Limitations of the Existing Detailed Surface Reaction Mechanisms 7.2.1 Adsorption of CH4 7.2.2 The Role of Adsorbed Oxygen 7.2.3 Adsorbate-Adsorbate Interactions 7.2.4 Other Limitations

126 127 128 129 129

8 Partial Oxidation of Ethane and Higher Alkanes 8.1 Production of Olefins from Alkanes using Pt and Rh 8.2 The Mechanism of Alkane Dehydrogenation

130 131 131

9 Conclusions and Outlook

133

125 125 125 126 126

Acknowledgement

134

References

134

Chapter 5 Indirect Liquefaction - Where Do We Stand? By Yongqing Zhang and Burtron H. Davis

138

1 Introduction

138

2 Today’s Commercial Operations 2.1 South Africa 2.2 Sasol 2.3 Mossgas 2.4 Shell

139 140 140 145 145

3 Large Pilot/Demonstration Plant Operations 3.1 Rheinpresussen-Koppers 3.2 British Fuel Research Station

148 148 153

Contents

xii

3.3 US Bureau of Mines 3.4 Stanolind/Carthage Hydrocol

153 155

4 Standard Oil Co. (New Jersey) - Exxon

156

5 Pilot Scale Operations 5.1 Syntroleum 5.2 Gulf Oil 5.3 Rentech 5.4 Chinese Studies 5.5 Mobil Oil

161 161 170 171 172 172

6 Process and Economic Evaluations

174

7 Potential Commercial Operations

175

8 Summary of Current Status

180

Acknowledgement

182

References

182

Chapter 6 Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts By Akifumi Ueno

185

1 Introduction

185

2 Partial Oxidation of Methane on Mo031Si02 and Alkali Metal-doped Mo03/Si02

186

3 Active Sites on Mo03/Si02and Reaction Mechanism of Selective Oxidation of Methane

192

4 Characterization of Surface Species Generated on Mo03/Si02 by Heat Treatments

197

5 Catalytic Activity of SMA Generated On or Impregnated on Si02

204

6 Other Topics Concerning Heteropoly Acid Catalysts or Partial Oxidation of Methane

208

7 Summary of the Selective Oxidation of Methane on Mo03/Si02 Catalysts

21 1

References

212

1 Strong Solid Bases for Organic Reactions BY YOSHIO O N 0 AND TOSHlHlDE BABA

1

Introduction

Carbanions are important intermediates in many organic reactions such as isomerizations, additions, alkylations, and cyclizations. They are formed by abstraction of a proton from a C-H bond of an organic molecule by a base. These organic reactions often require a stoichiometric amount of liquid base to generate carbanions and produce a stoichiometric amount of metal salts as a by-product. For example, the methylation of phenylacetonitrile with methyl iodide proceeds in the presence of base under a phase-transfer condition. PhCH2CN

+ CH31 + NaOH

-

PhCH(CH3)CN

+

Nal

+

H20

(1.1)

In this case, more than a stoichiometric amount of sodium hydroxide is required to neutralize the hydrogen iodide produced and to keep the system basic. Furthermore, a stoichiometric amount of sodium iodide is inevitably formed and has to be disposed of in an appropriate manner. Organometallic compounds such as Grignard reagents and alkyl lithium serve as donors of carbanion-like species. Here, again, a stoichiometric use of these reagents is required. Therefore, there is a need to develop solid bases to avoid these problems. Solid base catalysts have many advantages over liquid bases. They are noncorrosive and environmentally benign, presenting fewer disposal problems, while allowing easier separation and recovery of the products, catalysts and the solvent. Thus, solid base catalysis is one of the economically and ecologically important fields in catalysis and the replacement of liquid bases with heterogeneous catalysts is becoming more and more important in the chemical industry. * Furthermore, high activities and selectivities are often obtained only by solid base catalysts for various kinds of reaction. Since the ability of bases to abstract a proton from a C-H bond is directly connected to the base strength, stronger bases are in general more effective in forming carbanions. Alkaline earth oxides such as magnesium oxide are strongly basic when properly pretreated. Extensive works by Tanabe, Hattori and their co-workers have been carried out using these material^.^.^ Recently, other strong base catalysts have been reported. Potassium amide supported on alumina (KNH2/A1203) is effective for a number of baseCatalysis, Volume 15 0The Royal Society of Chemistry, 2000 1

2

Catalysis

catalysed Even toluene is activated to react with silanes at 329 K,8 and the isomerization of 2,3-dimethylbut- 1-ene proceeds even at 20 1 K.495 Potassium fluoride supported on alumina (KF/A1203) has been used by organic chemists for a long time,10-12but Tsuji and Hattori revealed that this catalyst becomes much more active when pretreated at 573-673 K under vacuum. l 3 Yamaguchi et al. reported that catalysts prepared by loading alkalimetal salts such as KN03, followed by heating at 773-873 K, were very strongly basic and active for the isomerization of cis-but-2-ene at 273 K.14 Fu et al. used alkali-metal compounds supported on alumina for the reaction of catechol with dimethyl carbonate and found that the rate and selectivity depended strongly on the catalyst used.I5- Furthermore, modified zeolites and calcined hydrotalcites are often reported as strong bases. In this review, we will describe the preparation and characterization of these strong bases. Then, application of these catalysts to a variety of catalytic reactions is described. The reactions include the isomerization of alkenes and alkynes, the dimerization of alkynes, aldol reactions, and the formation of Si-C, Si-N and Si-0 bonds. 2

Role of Solid Base and Basic Sites as a Catalyst

2.1 Abstraction of Protons - On the surface of solid bases, there are specific sites or centers, which function as a base. Basic sites (centers) abstract protons from the reactant molecules (AH) to form carbanions (A-),

Here, the basic site B- on the solid surface acts as a Bronsted base. Stronger bases can abstract a proton with molecules with higher pK, values.

2.2 Activation of Reactants without Proton Abstraction - Reactants such as ketones and aldehydes are often activated by bases without proton transfer, as expressed by the following equation.

I

R2

B

Here, the basic sites B - act as a Lewis base. It should be noted that a same surface site can serve as a Bronsted base as well as a Lewis base, depending on the nature of the adsorbate.

2.3 Cooperative Action of Acidic and Basic Sites - Magnesium oxide is active for the hydrogenation of 1,3-butadiene. It is assumed that hydrogen heterolytically dissociates in the presence of a pair of a coordinatively unsaturated Mg2+ and an oxide Hydrogen adsorption is schematically expressed as:

1: Strong Solid Bases for Organic Reactions

3

(1.4)

3

Base Strength of Basic Sites

3.1 H- Acidity Function - The H - acidity function is defined as a measure of the ability of the basic solution to abstract a proton from an acidic neutral solute.

To determine the H - value of a solution, the concentrations of AH and A have to be measured accurately. When half of a solute AH is deprotonated in the solution, i.e., [A-] = [AH], the H - value of the solution is equal to the pK, value of AH. The basic strength of a solution is stronger when a neutral molecule of larger pK, value is deprotonated. Tanabe proposed transferring this concept to solid bases as a measure of their strength.2 The base strength of solid bases is expressed by means of the H - value, equated to the highest among the pK, values of the adsorbates from which the basic site is able to abstract a proton. Tanabe defined solid superbases as materials with H - values higher than 26. This value, like that of superacids ( H - < -12), is 19 units from a neutral solution of 7. In the use of this concept for solid bases, two important points should be noted: (a) In the discussion of solid bases, the H - value is treated as a parameter to describe the nature of individual basic sites. It is often assumed that there are a certain number of basic sites on solid surfaces and that each of the sites has its own basic strength. In the original definition, H - scale is used to describe basic property of the solution, not that of individual basic molecules (or ions) in the solution. (b) In principle, the idea of the H - scale is only applicable to the Bronsted base. It is not, at least directly, related to the ability of the sites to function as Lewis bases, as shown in eqn. (1 3. 3.2 Indicator Method - The H - values of basic solutions are determined by using indicator molecules.18If the pK, value of the indicator AH is known, the H - value can be calculated by determining the ratio of [AH]/[A-I. To cover a wide range of the H - scale, a series of indicators with different pK, values have been selected to obtain the accurate value of [AH]/[A-1. In the case of solid bases, the color change of indicator molecules upon adsorption is taken as a measure of basic strength.* If the color change of the indicator is observed, the H - value of the basic sites on the solid is higher than the pK, value of the indicator. Similarly, if the indicator does not change color

4

Catalysis

upon adsorption, the H- value of the sites is judged to be less than the pK, value of the indicator. By using indicators of different pK, values, the H value of the basic sites can be determined. It is important to know that the color change is due to proton abstraction by basic sites and not due to other types of interactions such as charge-transfer between the adsorbate and the surface.

3.3 Other Methods for Determining Basic Strength - Temperature programmed desorption of carbon dioxide is often used. When the interaction between basic sites and carbon dioxide is stronger, the molecule desorbs at higher t e r n p e r a t ~ r e . ~One . ' ~ disadvantage of using carbon dioxide is that this molecule adsorbs on solid surfaces in several different forms. For example, as revealed by infrared spectroscopy, carbon dioxide is adsorbed on alkaline earth oxides to form a unidentate complex as well as a bidentate complex.20 Since the interaction of carbon dioxide with the surface does not involve a proton-transfer process, the result may not be directly related to the Bronsted basicity of the sites. The XPS binding energy value of elements depends on the charge carried by the atom. The binding energy of 0 1 , is then expected to decrease with increasing negative charge on the oxygen. The 0 1 , binding energy of X- and Y-type zeolites decreases with increasing Si/Al ratio and decreasing electronegativity of the counter cation.21'22Since the XPS binding energy is measured as the average of those for all of the 0 atoms in the material, the method is not applicable for the materials where only a fraction of the oxygen ions are active as basic sites, as in most alkaline earth oxides. Infrared spectroscopy of adsorbed molecules is often used for characterizing surface bacisity. Pyrrole is an amphoteric molecule. The basic strength may be estimated from the shift of NH vibration upon its interaction with basic sites through hydrogen bonding. For example, when pyrrole interacts with framework oxygen ions in zeolites, the NH vibration is shifted to low wavenumbers from 3430 cm- in the pure liquid to around 3200 cm- .23 The extent of the shift increases with the negative charge on oxygen, which is calculated from Sanderson's electronegativities. Though pyrrole adsorption has become a popular technique, the spectra are rather complex and the molecule is not always stable on the surface; it polymerizes or dissociates on some oxides.24 The shift of C-D stretching mode of adsorbed CDC13 is also a measure of the basic strength. Berteau et al. examined the base properties of modified aluminas by IR spectroscopy of probe molecules and C 0 2 TPD.25 Acetylene and substituted acetylenes have also been used as probe molecules for surface basicity.26 Bosacek proposed the use of 13C NMR of adsorbed methyl iodide for the basicity of zeolites.27Methyl iodide heterolytically dissociates and the methyl group attaches to the lattice oxygen. The chemical shift of the carbon, therefore, reflects the actual electronegativity of the oxygen. With this method, Hunger et al. confirmed that strongly basic sites were created by incorporating alkali metal oxides into the zeolite pores.28

1: Strong Solid Bases for Organic Reactions

4

5

Base Strength and Catalytic Reactions

Catalytic reactions provide an accurate measure of basic strength, especially when the reaction starts by formation of carbanions by abstraction of a proton from the reactant, since the ease of carbanion formation depends on the PKa value of reactant. In Table 1.1, PKa values of various compounds are listed. Isomerization of alkenes such as but-1 -ene proceeds through the formation of allylic anions, which are formed by abstraction of a proton from but-1-ene, as shown in reaction ( 1.6).29

Since the PKa value of alkenes is high (PKa of C3H6=35), strong bases are required to activate the alkene molecule. Thus, alkene isomerization is an appropriate test reaction for strong solid bases. Moreover, the reaction is mechanistically simple. This makes the interpretation of experimental results straight forward. Table 1.2 shows the catalytic activities of various solid bases for the isomerization of 2,3-dimethylbut- 1-ene (DB-1) to 2,3-dimethylbut-2-ene after 20 h.30The activities vary significantly from one catalyst to the other, reflecting a wide variety of the base strength and the number of the basic sites. KY (K+exchanged Y zeolite) has no activity, indicating that no strong basic sites exist on KY. On the other hand, there are groups of catalysts that have very high activities: alkali amides on A1203,alkali compounds on A1203,CaO and MgO. Since the conversion over these catalysts is close to the equilibrium value at 3 13 K, it is hard to know the relative ranking of activities for these materials from Table 1.2. Table 1.3 shows the results when the isomerization was carried out at much lower temperature, i.e. 201 K, and a shorter reaction time on these materials, RbNH2/A1203, a series of alkali hydroxide supported on A1203,and K loaded on A1203 prepared by the deposition of K vapor (K/A1203), being added to the list. The reaction is very fast over RbNH2/A1203, KNH2/A1203, CsOH/A1203 and CaO even at 201 K. The order of the activities for the most active class of solid base catalysts is as follows: RbNH2/A1203> KNH2/A1203 > CaO > MgO > CsOH/A1203 > KN03/A1203> RbOH/A1203> KOH/A1203> K/A1203 The PKa value of propene is 35 and we can assume that the pK, value of DB-1 is not far from this value. Therefore, all of these catalysts can be classified as solid-superbases. CaO was reported to have basic sites stronger than H - = 26 by an indicator method.2 As mentioned above, KY does not show the catalytic activity for the isomerization of DB-1, whereas it catalyses a Knoevenagel reaction of benzaldehyde with ethyl ~yanoacetate.~' O

C

H

-

O + NCCH2COOEt

O

CN C

H

=

(1.7)

& COOEt

6

Catalysis

Table 1.1 Approximate acidities of various common reagents PKl

4A

9

PK,

ref:

NH3

35

a

H2C=CH2

36.5

b

37

b

37

a

CH4

40

b

H3C-CH3

42

b

H3C-CH2-CH3

44

b

45

b

ref: a

H H

OH

9

11

a

13

a

15

C

15.75

a

16.7

C

y

3

H H

CH30H CH3CH20H

O

C

E

C

H

17 18

a

18.5

b

a

0 II

CrCH3

19

a

25 25

a,b

CH3CN CH3C02C2H5

25

a

25

a

HCGCH

a

H.O. House, Modern Synthetic Reuction, W.A. Benjamin, INC., Menlo Park, CA, USA, 1972, p. 494. D.J. Cram, Fundamentuls of Curbanion Chemistry, Academic Press, New York and London, 1965.

a

This indicates that basic sites of KY are able to abstract a proton from the latter, which has a pK, value of 8.6. Therefore, when this reaction proceeds over a solid catalyst, it is judged that the solid has basic sites stronger than H- =8.6. Though the H - values of a base catalyst can be used to decide whether a

1: Strong Solid Basesfor Organic Reactions

7

Table 1.2 The catalytic activities of solid base catalysts for the isomerization of 2,3-dimethylbut-l-ene at 313 K after 20 h 303a

Catalyst

Metals supported from liq. NH3b*C KNH2/A1203(2.6) K(N H 3)/AI2O3(2.O) Na(NH3)/A1203(3.5) E~(NH3)/A1203(0.5) Y b(N H3)/A1203(0.5) E~(NH3)/A1203(0.5) Y b( N H3)/A1203(0.5) Alkali metals/A1203 KN03/A1203(5.O) KOH/A1203(5.0) KF/A1203(5.0) Metal oxides CaO MgO SrO BaO Sm203 EU203 Yb203 A1203 Mixed oxidesd 4 CaO, A1203 4 MgO, A1203 Zeolites KY

Pretreatment TemperaturelK Timelh

573 423 423 523 473 423 473 873 673 623

Conversion (%)

91.3 89.4 89.3 83.4 29.4 2.3 1 .o 3 3 3

90.7 86.0 25.6 89.8 84.9 1.6 0.4 1.2 1.o 0.9 3.9

1073 873 1023 1073 773 923 923 773 1073 773

1 4

11.9 1.3

773

3


Conditions: catalyst = 0.25 g, 2,3-dimethylbut-l-ene= 3 ml (24 mmol). Number in parentheses: supported amount in mmol g-'. KNH2/A1203 prepared by loading KNH2 on A1203 from anmmoniacal solution of KNH2, M(NH3) (M = K, Na, Eu and Yb) prepared by impregnation from the metal dissolved in liquid ammonia. Prepared from hydrotalcite structure.

reactant can be activated to form carbanions by the solid base or not, it is not absolute. As stated above for alkali-exchanged faujasite, although the H values of alkali-exchanged zeolites are estimated to be 10- 13 from the rates of Knoevenagel condensations at 363-443 K,31 they can catalyse the reaction of phenylacetonitrile with dimethyl carbonate at 533 K.32

Rb- and Cs-exchanged X-zeolites can catalyse the side chain alkylation of toluene with methanol at 700 K.33

8

Catalysis

Table 1.3 The catalytic activities of solid base catalysts for the isomerization of 2,3-dimethylbut-l-ene at 201 K307a

Pretreatment temperaturelK

Catalyst

React ion timelmin

Conversion

90.2 87.3 93.8 49.4 4.2 0.1

(“w

R bNHz/A1203(2.6)b’c KHN2/A1203(2.6)b”

573 573

KN03/A1203(2.6) K$03lA1203(2.6) NaNH2/A1203(2.6)

873 873 523

10 10 30 30 30 10

CsOHIA1203 RbOH/A1203 KOH/A1203 NaOH/A1203 LiOHlA1203

873 873 873 873 873

30 30 30 30 30

67.8 36.4 34.1 0 0

CaO

998

10

63.0

K/AI,O~(12.0)

573

10

7.0

~

a

~~

Conditions: catalyst = 0.25 g, 2,3-dimetylbut-l-ene = 3 ml (24 mmol), pretreatment time = 1 h. Number in parentheses: supported amount in mmol g-I. 0.063 g.

These facts clearly show that these catalysts can activate phenylacetonitrile (pKa=21.9) and toluene (pKa=37) at 533 and 700 K, respectively. Base strength increases with temperature. Weak bases (as measured at room temperature) can be catalysts for a variety of reactions at higher temperatures, as long as they are stable.

5

Solid Base Materials

5.1 Alkaline Earth Oxides - Alkaline earth oxides (MgO, CaO, SrO and BaO) are active for a various types of reactions including isomerization of alkenes. To obtain a high activity it is essential to remove adsorbed molecules such as carbon dioxide and water. The catalytic activities of these oxides depend on the pretreatment temperature. The dependence of the catalytic activities of MgO on outgassing temperature for various reactions is shown in Figure 1.1. 3 When the pretreatment temperature is low (below 700 K), MgO shows no activity for the isomerization of but-1-ene. The catalytic activity develops at a pretreatment temperature of 800 K and declines at higher pretreatment temperatures. The maximum activity for H-D exchange between CH4 and Dz appears at 923 K, while the activities for hydrogenations develop at a higher pretreatment temperature of 1200- 1300 K. The dependence of the catalytic activities on pretreatment temperature indicates that there are at least three types of basic site on the surfkce of MgO. A model of MgO surface shows that there are several types of oxygen anions with different coordination number on the surface.34 It is plausible that each type of oxygen anion

I : Strong Solid Bases for Organic Reactions

9

Pretreatment temperature/K

Figure 1.1 Effect of pretreatment temperature on the catalytic activities of MgO for different types of r e a ~ t i o n0 . ~ But-I-ene isomerization at 303 K, A CH4-D2 exchange at 673 K, A amination of 1,3-butadiene with dimethylamine at 273 K, 0 hydrogenation of 1,3-butadiene at 273 K, 4 hydrogenation of ethene at 523 K.

manifests its own basic strength, and changes in amount, with pretreatment conditions. Oxygen anions at low coordination numbers exist at corners, edges and high Miller index surfaces. As pretreatment temperature increases, desorption of adsorbed molecules such as carbon dioxide occurs and oxygen anions become available for the reactants. Desorption of adsorbed molecules starts from weaker basic sites and more severe pretreatment is required for generating stronger basic sites. At the same time, pretreatment at higher temperature causes the rearrangement of the surface structure. These two factors induce a complex dependence of the catalytic activities with pretreatment temperature. The base strength of alkaline earth oxides was determined by an indicator method.35 The acid strength of the strongest sites on MgO and CaO exceeds H - = 26, showing that these oxides are superbases. Higuchi et al. estimated the number of active centers on CaO as 2.7 mmol g- by examining the inhibition effect of trichloroacetic acid on the cyanosilylation of benzophenone on this catalyst.36

5.2 Alkali Metals Supported on Metal Oxides - Alkali metals loaded on supports by deposition of the metal vapor have been reported as highly active catalysts for the isomerization of alkenes and related corn pound^.^^^^* For example, sodium metal deposited on alumina (Na/A1203) isomerizes but- 1 -ene and pent-1-ene at room t e r n p e r a t ~ r eMalinowski .~~ and Kijenski also reported that sodium metal deposited on MgO (NdMgO) showed a high catalytic

10

Catalysis

activity for the isomerization of alkenes at 293 K, and the base sites of Na/ MgO are stronger than H- = 35.37 They proposed the generation of strongly basic sites by the reaction of the alkali metal with the holes trapped on oxygen anions near the cationic ~ a c a n c y . ~ ~ - ~ ' 5.3 Alkali Metals on A1203 Treated with Alkali Metal Hydroxides - Suzukamo and co-workers prepared strongly basic solid catalysts by the reaction of NaOH and Na.42343The catalyst will be denoted as Na/NaOH/A1203. A typical preparation procedure is as follows. Alumina is added to sodium hydroxide by portions at 583-593 K with stirring and the water generated is removed by a flow of nitrogen. The stirring is continued for 3 h at the same temperature to give white solid (A). The sodium is then added, and the reaction mixture is stirred for 1 h at the same temperature to give a pale blue solid. The intermediate solid (A) is disordered sodium aluminate which does not show Xray diffraction lines due to sodium metal. The 01,photoelectron peak shifts to lower binding energy in the order of Na/NaOH/A1203, solid (A), and yalumina, and this suggests that 0 on Na/NaOH/A1203 atoms have increased negative charge and a higher electron pair donating ability that is strong enough to exhibit high b a ~ i c i t yThey . ~ ~ suggest ~~ that ionization of metallic sodium on the disordered aluminate plays a key role in generating extremely strong basic sites on Na/NaOH/A1203. Na/NaOH/A1203 is an effective catalyst for the isomerization of 5-vinylbicycl0[2.2.1]hept-2-ene to 5-ethylidenebicyclo[2.2.l]hept-2-enewhich is an important component of ethylene-propylene synthetic rubber. The reaction proceeds completely, even at 243 K.42-43

(1.10)

Ethylation of cumene proceeds at 313 K. A very high selectivity (99.6%) is obtained at a cumene conversion of 99.9% over K/KOH/A1203, prepared in a similar manner to Na/NaOH/A1203. 5.4 KNH2/A1203- Potassium supported on alumina prepared from the ammoniacal solution is very active for the isomerization of pent-1-ene, even at 201 K. The conversion reaches 98% in 6 min.4 This activity is far greater than that of Na/A1203 which is prepared by the vapor deposition of Na on alumina. The detailed study of K supported on alumina prepared from liquid ammonia revealed that the active species in this catalyst is potassium amide.4 Further studies show that the amide species is more effectively formed when Fez03 is added as a catalyst for converting K into KNH2 in liquid a m m ~ n i a . The ~ . ~ very high activity of KNH2/A1203is clear from Tables 1.2 and 1.3. KNH2/A1203 is prepared as f01lows:~~~ alumina and a small amount of

I : Strong Solid Basesfor Organic Reactions

11

300400500600700

Evacuation Temperature / K Figure 1.2 Effect of evacuation temperature on the catalytic activity of KNH21A1203for the isomerization of 2,3-dimethylbut-l-ene.4*5Reaction conditions: 201 K; catalyst weight=0.063 g; supported amount as K metal=2 mmol g alumina- l ; 2,3-dimethylbut-l-ene = 24 mmol; reaction time = 10 min.

Fe203 are heated in the reactor under vacuum at 773 K for 3 h. A piece of potassium metal is then put into the reactor under nitrogen. After evacuation, ammonia is liquefied into the reactor to dissolve the metal. The blue color due to solvated electrons disappears within about 10 min, indicating the formation of potassium amide. After 1 h, the reactor is warmed to room temperature in order to remove most of the ammonia and is then heated under vacuum at a higher temperature for 1 h. Figure 1.2 shows the dependence of the catalytic activity on evacuation temperature of KNH2/A1203for the isomerization of 2,3-dimethylbut- 1-ene at 201 K.475 The catalytic activity strongly depends on the heating temperature and reaches the maximum at around 573 K. At 673 K, the activity sharply declines due to the decomposition of KNH2, as evidenced by the evolution of hydrogen and nitrogen. This result shows that high temperature treatment (573 K) is essential to obtain a strongly basic solid catalyst, although KNHz supported on alumina dried at 338 K was proposed as a base catalyst by H ~ r b e r t . ~ ~ - ~ ~ Figure 1.3 shows the dependence of the catalytic activity for the isomerization of 2,3-dimethylbut-1-ene on the loaded amount of p o t a ~ s i u m Alumina .~ has no activity under the reaction conditions. The isomerization does not proceed at all when the loaded amount is below 3 wt% as K. This potassium content is almost equal to the amount of OH groups on alumina surface, indicating that the OH groups react with potassium, which may be converted

Catalysis

12

80

1

I

I

I

I

1

60 -

-

40

-

20

-

0

0

2

4

6

8 1 0 1 2 1 4

Amount of K(Meta1) / wt% Figure 1.3 The dependence of loaded amount of potassium on the catalytic activity of KNH2IA1203 for the isomerization of 2,3-dimethylbut-1-ene (24 mmol) at 201 K. Reaction time = 10 min; catalyst = 0.063 g.5

Table 1.4 Influence of supports on the catalytic activity of KNH2" Support

Reaction timelmin

Conversion (YO)

A1203 A1203-Mg0 (Mg/Al = 2 ) CaO Si02

10 30

70 19

30 30 30 10

19 0 0 63

Ti05 CaO (pure) a

Reaction conditions: 201 K, catalyst: 63 g, reactant: 24 mmol. Without KNH2.

into OK groups, having no activity for the isomerization. The activity arises when the loaded amount exceeds 3%wt. Alumina is a unique support for KNH2. Table 1.4 shows the effect of supports on the catalytic activity of KNH25 for the isomerization of 2,3dimethylbut- 1-ene. Though KNH2/A1203 has a very high activity, KNH2/Si02 and KNH2/Ti02 are totally inactive. A1203-Mg0 (AYMg = 1/2) obtained by the decomposition of hydrotalcite has some activity. This material is known to be a solid base, as described later. KNH2/CaO also have some activity, but the activity is much lower than CaO or KNH2IA1203. Thus, the property of the support is important to the development of the high catalytic activity in KNH2/support systems. The nature of the strong interaction of KNH2 and alumina is still an important problem to be solved in future.

1: Strong Solid Bases for Organic Reactions

13

Wavenumbers / cm-' 3-mahyl-1 -but-

Figure 1.4 The infrared spectrum of KNHJA1203 and its reaction with 0 2 and 3methylbut-1 -ene. (a) KNH21A1203 as prepared; (b) sample then exposed to Dz at 373 K for 30 min; (c) sample treated with 3-methylbut-1-ene at room temperature for IO min.5

The infrared spectrum of KNH2/A1203 is shown in Figure 1.4(a).5 Two bands at 3238 and 3190 cm-' and a single band at 1541 cm-I are ascribed to N-H stretching and NH2 bending bands, respectively, indicating the presence of KNH2 on the surface. The band at 1939 cm-' may be assigned to an N = N band. When the sample was exposed to deuterium at 373 K for 30 min the three bands due to KNH2 disappeared, and three new bands appeared at 241 1, 2343 and 1132 cm- [Figure 1,4(b)]. These three bands are attributed to ND2 groups, indicating that the exchange reaction between KNH2 and D2 occurred to form KD2 on the alumina surface. When the sample which shows the bands due to the KNH2 groups was exposed to 3-methylbut-1-ene at room temperature for 10 min, the bands due to ND2 groups disappeared and bands due to NH2 groups reappeared, besides the bands due to CH stretching and bending [Figure 1.4(c)]. This shows that the NH2 groups are involved in the isomerization of alkenes. In fact, it was confirmed that D atoms were contained in the reaction products in the gas phase. The NH2 groups in KNH2/A1203 react even with methane (pK, = 40).48 Thus, the H-D exchange reaction proceeds between KND2/A1203 and CH4 at room temperature. The H-D exchange reaction was also observed between KND2/A1203 and C2H6.48The rate of the exchange of KND2 with CH4 was faster than that with C2H6, consistent with the difference in their acidities.

14

Catalysis

5.5 Alkali Metal Compound Supported on Alumina 5.5.I KN031A1203, KHC03/A1203, K2C03/A1203 - Potassium compounds such as KN03, KHC03 and K2CO3 supported on alumina and heated at high temperature are highly active catalysts for the isomerization of cis-but-2-ene at 273 K.35 Tables 1.2 and 1.3 show that these materials are active for the isomerization of 2,3-dimethylbut- 1-ene.30 The supported compounds are decomposed at least partly during heat-treatment and the origin of the basic sites is not clearly identified. These catalysts, however, have an advantage in the ease of the preparation and strong basicity. 5.5.2 LiOHIA1203, NaOHIA1203, KOHIA1203, RbOHIAl203, CsOHlA120~-

The loading of NaOH and KOH was often used as a method for suppressing the acidity of alumina. Fu and co-workers found that alkali metal hydroxides supported on alumina offer great opportunities as solid bases.15-17 In the vapor-phase reaction of catechol with dimethyl carbonate, LiOHlA1203 gave a high selectivity for anisole, whereas CsOH/A1203 showed a very high activity for veratrole ( 1,2-dimethoxybenzene). KOHIAl2O3, RbOHlA1203 and CsOHI A1203 exhibit very high catalytic activity for the isomerization of 2,3dimethylbut-1-ene, even at 201 K when treated at high temperatures as shown in Table 1.3, indicating that these are very strong solid bases.30 LiOHlA1203 and NaOH/A1203 have no activity for the isomerization under the same reaction condition^.^^ Zeolites - As mentioned above, alkali ion-exchanged zeolites, especially faujasites such as KY, are weak bases, They can be used to activate even toluene at high temperature^.^^ One of the advantages of weak bases such as zeolites is easy handling. They can be handled in ambient atmosphere since the adsorption of carbon dioxide or water is not too strong and can be removed by high temperature treatment. They can be used even in cases where these molecules are involved as reactants or products, as in reactions (1.8) and (1.9). The basic strength of alkali ion-exchanged zeolites has been rationalized by the negative charge of the oxygen atoms calculated using the Sanderson's electronegativity equalization principle.49 The base strength of X-zeolites is higher than the corresponding Y-zeolites. The base strength depends on the exchangeable cations and increases in the order, Li < Na < K c Rb c Cs.49950 This tendency was also confirmed experimentally by the XPS study of the zeolites5' as well the infrared study of adsorbed pyrr01e.~~ The basic strength of alkali ion-exchanged zeolites can be estimated by using the Knoevenagel condensation reactions, reaction ( 1.7), as described above.31 The activity of the zeolites is Li < Na < K < Cs and Y < X zeolites. The order of activity is in conformity with the order of the average negative charge on the oxygen atom calculated by the electronegativity equalization principle. They concluded that most of the basic sites on alkaline X and Y zeolites have H- < 10.3 and sites with H - < 13 were present only in CsX. This catalyst is more active than pyridine (pKb=8.8), and less active than piperidine (pKb = 11.12). 5.6

1: Strong Solid Bases for Organic Reactions

15

5.7 Modified Zeolites - As mentioned above, alkali ion-exchanged zeolites are weak bases. Various efforts have been made to increase the base strength of alkali metal-ion-exchanged zeolites. Martens et al., formed metallic sodium particles in zeolites by the decomposition of occluded sodium Zeolites thus prepared exhibit catalytic activity for the isomerization of butenes at 300 K and the side a chain alkylation of toluene with ethylene at 523 K. Strongly basic catalysts can be prepared by loading CsNaX and CsNaY with cesium acetate followed by thermal decomposition of the acetate into the The catalytic activity for the dehydrogenation of isopropyl alcohol to acetone increased an order of magnitude by loading Cs onto CsY. The isomerization of but-1-ene proceeds over these catalysts at 273 K. Higher basic strength of the materials compared with the parent zeolite was confirmed by IR of adsorbed CHC13 and 13C MAS NMR of adsorbed methyl iodide.28 Various evidence including 133CsNMR indicates that the active species is nanophase cesium oxide occluded in the supercage of the zeolites. Recently, Zhu et al., reported that KN03/KL formed superbasic sites upon heating at 673 K with H- =27 determined by using Hammet indicator^.^^ They also reported that MgO dispersed on KL or NaY by microwave irradiation was strongly basic and that these materials catalysed the isomerization of cis-but-2-ene at 273 K.58 The catalysts were prepared by grinding a mixture of MgO and the zeolite in a rotor, followed by irradiation in a microwave oven (2450 MHz) for 20 min. Another method for obtaining strongly basic catalysts is loading of ytterbium or europium on Y-zeolites from the ammoniacal s o l ~ t i o n . ~These ~-~~ zeolites thus loaded with Yb or Eu have a high catalytic activity for the isomerization of but-1-ene at 273 K when they are heated under vacuum at 500 K. The materials also catalyse the Michael reaction.59 Thus, reaction of cyclopent-2-enone with dimethyl malonate over Eu-loaded KY gave a coupling product in an 81% yield at 303 K for 20 h. The catalytic activity is greatly influenced by the alkali-metal cations present in the Y-zeolite as shown in Figure 1.5.59 The high cisltrans ratio of but-2-ene produced indicates that the isomerization proceeds via an anionic intermediate and that higher basicity is obtained in the case of RbY or CsY. The nature of the catalyst depends very much on the evacuation temperature after loading these metals from the ammoniacal solution. The effect of evacuation temperature on the catalytic activities of ytterbium introduced into KY for but- 1-ene isomerization, Michael addition of cyclopenten-2-one with dimethyl malonate, and ethylene hydrogenation is shown in Figure 1.6.60The high catalytic activities for base-catalysed reactions, namely, but- 1-ene isomerization and the Michael addition, develop at an evacuation temperature of around 500 K. At 900 K, hydrogenation activity arises. The catalytically active species for base-catalysed reactions are amide or imide of low valent rare-earth metals based on photoluminescence, temperature programmed desorption, IR and XANES spectroscopy. These catalytically active species seem to be formed during heating the catalyst by the

Catalysis

16

Initial Rate / lo4 m d g& min"

5 Li Na K Rb

cs

3.0 4.3 10.3 10.5 10.3

Eu

I

i i 1 I

I

I.s 4.4 9.1 9.6 13.0

i

W

1 r

.

1

r

.

i

=

I

Yb 1 I

1 I .

H-Y EwY

20

3

I

LS Na K Rb CS

15

10

l

-

l

-

l

.

0.90

Figure 1.5 Catalytic activities of Eu and Yb supported on alkali-cation-exchanged Y zeolites. 57 Reaction conditions: 273 K , catalyst weight = 0.50 g ; but-I-ene = 20.5 kPa; degree of ion-exchange: Li 67%, Nu I OOYO,K 99%, Rb 47% and CS 390/0.

reaction of the metals and adsorbed ammonia, and, at higher temperature, transformed into their nitrides, which may be active species for hydrogenation.

5.8 Eu,Yb/AI2O3 - Eu and Yb loaded on alumina from the ammoniacal solution and evacuated at an appropriate temperature shows the basic properties as those on zeolite described above.62Again, the catalytically active species are plausibly amide or imide of these rare-earth metals. Their catalytic activities are much higher when loaded on alumina than on zeolites. As shown in Table 1.2, these catalysts have activities for the isomerization of 2,3dimethylbut- 1-ene even at 3 13 K. 5.9 KF/A1203 - Potassium fluoride supported on alumina (KF/A1203) is a unique base catalyst, and many applications of this material have been developed by organic chemists. KF/A1203 is a convenient catalyst to handle and is now a commercially available common reagent. Examples of the reactions catalysed by KF/A1203 are given in reactions (1.1 1)-( 1.14).10-12,63 It is well established that reaction (1.15) occurs during the preparation of the catalyst. Weinstock et al., concluded that the active species was potassium hydroxide or the aluminate formed on alumina.64 On the other hand, Ando and coworkers concluded that remaining fluoride ions are the source of the strong

I : Strong Solid Bases for Organic Reactions

17

3004005006007008009001~ evacuation temperature/K

Figure 1.6 The catalytic activities

of YblNa-Y for .) the isomerization of but-1-ene, the hydrogenation of ethene, and 0 the Michael reaction of cyclopent en-2-one with dimet hyl malonate plotted vs. evacuation temperature. Fraction of Yb2' 0 and Ybj' also plotted vs. evacuation temperature.60 Isomerization of but-1-ene (20.5 kPa) carried out at 273 K over Yb(6.6 wt%)/ Na Y. Hydrogenation of ethene carried out at 273 K over Yb(6.6 wt%)lK Y; ethene: 13.3 kPa, hydrogen: 26.6 kPa. Michael reaction of cyclopenten-2-one (1 5 mmol) with dimethyl malonate ( 1 5 mmol) carried out at 323 K for 20 h. (a

(-.-)

PhCHO + CH2(CN)2 PhCHO + PhCH*=CN PhCHO + PhCHClCN 12KF + A1203 + 3H20

___c

-

PhCH=C(CN)2

(1.12)

PhCH=C(CN)Ph

(1.13)

phk%:N H O

(1.14)

2KfiIFe

+

6KOH

(1.15)

There is another complicating factor. Since preparation and basicity. drying (usually below 473 K) are done in an ambient atmosphere, adsorption of water and carbon dioxide is unavoidable. The effects of these adsorbed molecules have never been studied in detail. The effect of evacuation temperature on the catalytic activity of KF/A1203 for the isomerization of pent- 1-ene and the Michael addition of nitromethane to buten-2-one is shown in Figure 1.7.13The activity of KF/A1203 for pent-lene isomerization showed a sharp dependence on the pretreatment temperature. The sharp maximum of the activity was observed at around 623 K. The dependence of the activity for the Michael addition is not so sharp, though the maximum is also observed at 623 K. The sharp dependence of the catalytic activity on evacuation temperature is also found for the Tishchenko reaction of benzaldehyde66 and the disproportionation of trimethylsilylacetylene.8 In

Catalysis

18

100

80

60

40 20

0

Pretreatment Temperature / K Figure 1.7 Effect of pretreatment temperature on the catalytic activities of KFIA1203for the isomerization of pent-]-ene and Michael reaction between nitromethane and buten-2-0ne.I~ Pent-1-ene isomerization at 273 K; 0 Michael addition at 273 K.

the case of the reaction of benzaldehyde, no activity was observed at an evacuation temperature of 473 K, whereas it was about 50% upon evacuation at 623 K. These phenomena indicate that the active species developed upon evacuation at high temperatures may be entirely different from those of KF/ A1203used in usual practice. While KF/A1203 is a useful catalyst, further study is definitely required before the catalytically active species in ‘KF/A1203’ is fully understood. 5.10 Mixed Oxides - Various types of mixed oxides show basic properties

especially when they contain Mg or Ca as a component. In a number of combinations containing MgO, the number of basic sites is higher than that of neat MgO, probably because of increased surface area, but the basic strength of mixed oxides is usually much lower than alkaline earth oxides. In the case of MgO doped with Cs (as 3 mol% as Cs), the strength of the basic sites are higher than those of pure Mg0.67 Decomposition of hydrotalcite and its analogs at about 673 K is a typical way of preparing Mg-A1 mixed o ~ i d e s . ~ ~Corma - ~ O and co-workers studied the catalytic activity of a calcined Mg-A1 hydrotalcite for the Knoevenagel condensation of benzaldehyde with various activated methylenic compounds with different pK, values, and concluded that the calcined hydrotalcite has basic sites with H - up to 16.5, most of them being in the range of 10.7 < H - <

I : Strong Solid Bases for Organic Reactions

19

13.3.71The value indicates that calcined hydrotalcite should be classified as a mild solid base. The base-catalysed reactions over calcined hydrotalcite include the polymerization of P-propi~lactone,~~ aldol condensation of acetone,68 cross aldol condensation of acetone and formaldehyde to methyl vinyl k e t ~ n e ,and ~ ~the ,~~ isomerization of pent- 1-ene.69Recently, Kaneda and co-workers reported that Mg-A1 mixed oxides obtained by the calcination of hydrotalcites at 673 K catalysed the addition reaction of C02 with various kinds of epoxides to form cyclic carbonates under an atmospheric C02 pressure.75 0

(1.16) R'

For example, styrene oxide was converted into styrene carbonate at 373 K after 5 h with 94% yield. It is important to note that these mixed oxides usually have acid sites as well as basic sites. The role of acid sites should not be underestimated. Actually, the cooperative actions of acid sites with strong basic sites are proposed for the mechanism of reaction.75 In most of the catalytic applications of hydrotalcites, they have been used in their calcined form as mixed oxides. There are, however, several reports on the catalysis by hydrotalcites with the as-synthesized s t r ~ c t u r e . ~ ~ - ~ ' Organic halide exchange reactions producing alkyl bromides and iodides have been catalysed by hydrotalcites intercalated by Cl-, Br- or I - .76-78 Hydrotalcites also catalyse the disproportionation of HSi(OCH3)3 to form tetramethoxysilane and ~ i l a n e . ~ ~ 4HSi(OCH3)3

-

3Si(OCH3)4 + SiH4

(1.17)

Constantino and Pinnavaia carried out the reaction of 2-methyl-3-butyn-201 over a hydrotalcite,80.8' [Mg2.34A1(OH)6.68](Co~)0.~2.6H*o, treated at various temperatures. Acetone and acetylene were the only products over the temperature range of 353-423 K, demonstrating that all of the hydrotalcitederived catalysts have highly basic selectivities.82The thermal activation below the structural decomposition point ( C 523 K) afforded the catalysts with activities approximately an order of magnitude higher than the amorphous metal oxides formed by its thermal activation at 723 K.80,81

Catalysis

20

6

Reactions Catalysed by Solid Bases

6.1 Isomerizations 6.I . 1 Isomerization of Alkenes - Isomerization of alkenes proceeds through anionic intermediates by abstraction of an allylic proton from alkene molecules by solid bases. In the case of but-1-ene isomerization, a high cisltrans ratio of but-2-ene is a characteristic of the base-catalysed isomerization. On the other hand, the isomerization over solid acids proceeds through carbenium ions to afford a cisltrans ratio of near unity. Since pK, values of alkenes are high, strongly basic catalysts are required for the isomerization. In Tables 1.2 and 1.3 are listed solid base catalysts active for the isomerization of 2,3dimethylbut- 1-ene. Among them, alkali amide supported on A1203 and alkaline earth metal oxides show very high activity for the isomerization. Reaction (1.6) is catalysed by Na/NaOH/A1203,42.43KNH2/A1203,4Mg0,83 and Ca0.84 Isomerization of alkenes over KF/A1203 was also reported.85 The following are examples of alkene isomerizations of some more complex molecules. Catalysts used are alkaline earth o ~ i d e s . ~ ~ , ~ ~

8

(1.18)

(1.19)

(1.20)

(1.21)

6.1.2 Isomerization of Alkenyl Amines - Allylamines are also isomerized to enamines 1-N-pyrrolidino-2-propene to 1-N-pyrroridino- 1-propene over alkaline earth oxides at 3 13 K.88 N,N-Diethyl-3,7-dimethyloct-2-enylamine 1 gives the corresponding enamine 2 with a 100% E configuration in the presence of KNH2/A1203at 353 Kq6 The isomerization of N,N-diethyl-3,7-dimethylocta-2(Z),6-dienylamine3 proceeds at 313 K to afford exclusively N,N-diethyl-3,7-dimethylocta1,3-

1: Strong Solid Bases for Organic Reactions

21

dienylamine 4. The ratio of EIZ of the double bond at the 3-position in the product is almost 1: 1. The double bond at the 1-position is 100% E. 6.1.3 Isomerization of Ally1 Ethers - Ally1 ethers are also effectively isomerized over alkaline earth oxides.89 C=C-C-OC~HS

c=c-c-0

*C-C=C-OC2HS

-

c-c=c-0

( 1.24)

(1.25) (1.26) (1.27)

6.1.4 Isomerization of AZk-I-ynes - KNH2/A1203 is effective for the isomerization of alkynes. Thus, hex-1-yne is isomerized exclusively to hex-2-yne in a 92% yield in dioxane at 333 K in 20 h. CaO, prepared by evacuating CaC03 at 998 K also catalysed the isomerization of hex-1-yne. The yields of hex-2-yne and hex-3-yne were 79 and 13%, respectively at 313 K in 20 h. However, KNOJ A1203 and CsOHIA1203 showed no catalytic activity under the same reaction conditions.

6.2 Nucleophilic Addition Reactions of Phenylacetylene 6.2.I Dimerization of Phenylacetylene - Addition of phenylacetylene to double or triple bonds of various molecules occurs plausibly via formation of the alkynide ions, PhC = C - . The base-catalysed dimerization of phenylacetylene was first reported by Malkhasyan et aZ.89They found that phenylacetylene reacts in the presence of metallic sodium in an aprotic polar solvent to give 1,3-diphenylbut-1-en-3-yne and diphenylbut- 1,3-diyne. Later, Trofimov et al., further demonstrated that stirring of phenylacetylene with KOH-DMSO suspension in a ball-mill at room temperature for 1 h gave a mixture of E and 2 isomers of 1,4diphenylbut-1-en-3-yne in the ratio of 6: 1.90 It has been also reported that phenylacetylene polymerizes in the presence of sodium amide and phasetransfer catalyst^.^ * Phenylacetylene dimerizes in the presence of a KNH2IA1203 catalyst to afford (2)and (E)-1,4-diphenylbut-l-en-3-yne 5, the ratio of 2 :E being 96 :4, the yield being 97% in 20 h at 363 Ke7 The dimerization proceeds through the formation of an alkynide-type

Catalysis

22 2PhC=CH

-

Ph\ /C=CPh C=C\ H / H

Ph\

+

I H

iC="\C ECPh

(1.28)

E

Z 5

intermediate and its addition to the triple bond of another molecule of phenylacetylene (Scheme 1).

The predominance of head-to-head over head-to-tail dimerization is probably caused by the severe steric repulsion between the alkynide (i) and phenylacetylene in comparison with the case of head-to-tail dimerization. The high stereoselectivity is also caused by the stability of the Z-type intermediate (ii) over the corresponding E-type intermediate because of the steric repulsion between the bulky substituent and solid surface in the latter case. 6.2.2 Reactions of Alk-I-yne with Ketone or Aldehyde - As mentioned above, PhC = C- anions are generated on the basic sites of KNH2/A1203.This result indicates that alk-1-ynes such as PhCECH will also react with ketones and aldehydes through the formation of alkynyl anions over solid bases as shown in Scheme 2, where B- stands for a basic site on a solid surface.

Scheme 2

The reaction of phenylacetylene with cyclohexanone in dioxane in the presence of KNH2/A1203 afforded 1-phenylethynylcyclohexanolin a 67% yield in 20 h at 363 K.5792 (1 -29)

PhCrCH + 0 PhCEC

I : Strong Solid Bases for Organic Reactions

23

Thus, the dimerization of phenylacetylene did not occur at all, though it yielded lP-diphenylbut- 1-en-3-yne, 5 in the absence of cyclohexanone. Recently, Tzalis and Knochel have reported that most terminal alkynes undergo addition to secondary or tertiary aliphatic aldehydes or aliphatic ketones to give the corresponding propargyl alcohols in the presence of a catalytic amount of CsOH.H20, in yield between 60 and 9 6 Y 0 . ~ ~ The reaction of phenylacetylene with benzaldehyde in the presence of CsOH/A1203 in DMF gives chalcone, 6, in a 75% yield in 20 h at 343 K, together with 3% of 7.5992

PhCECH +

O

H

C

-

G

(1.30)

Chalcones are usually prepared by the cross-afdol condensation of benzaldehyde and phenyl methyl ketone in the presence of a base such as KOH.94 The reaction (1.30) offers a new synthetic route for chalcones. Phenylacetylene also reacts with 2-furylaldehyde to yield an a$-unsaturated ketone as a main product in the presence of Cs2C03/A1203, a-alkynyl alcohol being a side product.92

8

9

The yields of 1-furyl-3-phenylprop-2-en1-one, 8 and 1-furyl-3-phenylprop2-yn- 1-01, 9 were 49% and 5%, respectively, when 5 mmol of PhC = CH was reacted with an equimolar amount of 2-furylaldehyde over 0.125 g of Cs2CO3/ A1203at 363 K for 20 h. The reaction mechanism for the formation of 6 was investigated by using deuterated benzaldehyde (PhCD0).92 Thus, the reaction of phenylacetylene with deuterated benzaldehyde was carried out at 343 K for 20 h using CsOHl A1203 as a catalyst. Deuterated 6 (6-D) was a sole deuterated product, no other products containing D atoms being found, though the ratio of 6-D to 6 was 82/18 which was determined by * HNMR. This indicates the scrambling of H and D atoms may occur on the surface of the catalyst. On the basis of this observation, we propose the reaction scheme shown in

24 PhCECH

+

Catalysis

*‘.a /

D

6

6-D

Scheme 3. A unique feature of this mechanism is 1,3-shift of D-(H-) in the intermediate i to form ii. This type of hydride transfer has never been reported. 1,3-Diphenylprop-2-yn-l-o1 7 is plausibly formed by protonation of i, as expected from Scheme 2. The fact that the main product is 6 indicates that the transformation of the anionic intermediate (i) to (ii) is much faster than the protonation of (i). Obviously, alcohols are the products in the case of the reactions of alk-1-yne with ketones since the rearrangement of the anionic intermediates is not possible.

7-D

6-D

Scheme 3

I : Strong Solid Basesfor Organic Reactions

25

The reaction Scheme 3 suggests that there is a possibility that 6 is prepared by the isomerization of 7,if the protonation of (i) and the deprotonation of 7 are both reversible. H\

/p (1.33) w

7

6

When the isomerization of 7 (1 mmol) was carried out over Cs2C03/A1203in 1,4-dioxane (3 ml) at 303 K, 6 was obtained selectively. Thus, the yields of 6 were 34,92 and 97% after 1 h, 3 h and 20 h, re~pectively.~~ The isomerization of alkyn- 1-yl alcohol to an a,P-unsaturated ketone has not been reported. This novel rearrangement was observed also in the case of 9.92The yield of 8 was 98% in the reaction for 20 h at 363 K in the presence of CS~CO~/A~~O~. H

o-c=c-&; / \

O

\ I

-

(1.34)

/

9

8

6.3 Aldol-type Condensations - It is well known that +unsaturated ketones can be synthesized by aldol-type condensation reactions which proceed through formation of anions by proton abstraction by basic sites from active methylene compounds. Since the pK, values of active methylene groups are not so high, solid bases with mild basic strength are often effective for this type of condensation reaction. 6.3.I Aldol Condensations - Alkaline earth oxides show high catalytic activities for aldol condensation of acetone to diacetone alcohol at 273 K.94 The order of the catalytic activities is BaO > SrO > CaO > MgO. BaOH calcined at 473 K was also reported to be highly selective for this reaction.95 In the reactions of acetone with various substituted benzaldehydes, the expected aldol condensation products undergo dehydration to give, selectively, the corresponding a,P-unsaturated ketones in the presence of calcined hydroxide (Mg-Al-oxide) at 333 K as shown in Table 1.5.96 R

R\ RCHO

+

H IH

c=c

H3C’ C ‘CH3

H’

)-CH3 0

+

H20

( I .35)

6.3.2 Claisen-Schmidt Reactions - Among a,P-unsaturated ketones, stylyl ketones and distylyl ketones are obtained in high yield by the Claisen-Schmidt reaction in the presence of Ba(OH)2 treated at 483 K.97

26

Catalysis

Table 1.5 Aldol condensation between acetone and substituted benzaldehydes catalysed by Mg-A1 hydrotalcite at 333 Kg6va Ra

Reaction timelh

Me-

M

e O

e

Yieldb(%)

1

9F

1

96

1.5

95

0.5

98

1.5

94

1

96d

14

8c

MeO'

No reaction

15 OH

All reactions were performed on 2 mmol substrate in 10.3 ml (140 mmol) acetone using 0.2 g of Mg-A1 hydrotalcite. 'H NMR yields based on aldehyde. Aldol to dehydrated product ( I : I). Isolated yield.

a

(1.37) R4

eR2

I : Strong Solid Bases for Organic Reactions

27

Claisen-Schmidt condensation reactions of 2-hydroxyacetophenones and benzaldehyde derivatives proceed at 423 K in air using calcined hydrotalcites (mixed Mg-A1 oxide) as the catalysts.98 0

(1.38)

R’ I

Calcined hydrotalcite

*

RV

R

1+

H20

0 10

The primary product, 10, is further isomerized to flavone, 11, under the reaction conditions. 10

-

R*Rl

(1.39) 0 11

Reaction ( 1.40) proceeded to yield vesidryl (2’,4‘,4-trimethoxychalcone) in an 85% yield at 443 K.

(1.40) OCH3

6.3.3 Knoevenagel Condensations - Knoevenagel condensations are the reactions between ketone and active methylene compounds such as malononitrile and ethyl cyanoacetate. Calcined hydrotalcite is very effective for Knoevenagel reactions.99KnoeveR: ,C=O

R2

/CN

+

H2C, Y

-

R:

,CN

,C=C, R2 y

+

H20

(1.41)

28

Catalysis

nagel condensations also proceed over a variety of mildly basic catalysts: alkali-ion-exchanged zeolites;41alkali-ion-exchanged sepiolite;loo t-BuOK supported on xonotolite.lO1 6.3.4 Michael Additions - The reaction involves the nucleophilic addition of carbanions to a$-unsaturated carbonyl compounds. KFIA1203 catalyses Michael additions. -13 Meyer et al. prepared four strongly basic solid bases, namely, NdNaOH/A1203, Cs,O/A1203, Cs,O/CsX, and Na/NaX and carried out Michael addition of ethyl acrylate and acetone to form 5-oxohexanoic acid ethyl ester at 363 K.lo2

Na/NaX showed the highest activity (84% conversion in 24 h) with a selectivity of 57%, whereas higher selectivities of 65-70% were obtained with Cs,O/A1203 and Na/NaOH/A1203, but with lower conversion (44-5 1%). Cs,O/CsX showed a lower activity. 6.3.5 Dehydrocondensation with Methanol - The methyl and methylene groups at the a-position of saturated ketones, esters and nitriles are converted into vinyl groups by their vapor-phase reactions with methanol. 103*104 The catalysts are MgO modified with transition metal ions. For example, acetonitrile reacts with methanol over Mn-modified MgO to yield acrylonitrile at 648 K.

The reactions plausibly proceed through aldol-type condensations between an active methylene compound and a formaldehyde-like intermediate, formed by dehydrogenation of methanol on the surface.

6.4 Nucleophilic Substitution at Silicon - Si-C Bond Formation - The reactions of carbanions with silanes are one of the most important methods for achieving Si-C bond forming synthetic sequences.lo3-lo9 The most common sources of ‘carbanions’ are alkyl lithium and Grignard reagents. These reactions can be expressed in general as follows: R1 RLi + R2-ki-X A3

-

R1 I

R2-Si-R I R3

+ LiX

(1.44)

The detailed mechanism of the nucleophilic substitution has been extensively discussed. In many cases, the leaving group X is a halide ion such as C1-. Carbanions can be generated from various sources by the action of bases, but this method has rarely been used for Si-C bond formation. The general scheme of the reaction over solid base catalysts is expressed in Scheme 4.

I : Strong Solid Bases for Organic Reactions RH + B-

-

29

R- + BH

R’ I R- + R2-Si-X

R2-Si-R

A3

k3

X- + BH

R’ 1

+

x-

B- + HX

Scheme 4

Here, B- and X- denote the basic sites of catalysts and the leaving group from the Si atom, respectively. As shown in the following sections, the leaving groups in this case are hydride or an alkynide ion. This offers the possibility that novel routes for organosilicon compounds may be developed. 6.4.I Metathesis of Trimethylsilylacetyleene - Silylalkynes are commonly prepared from alkynes through reactions of the alkynide anion or its equivalent with a suitable silyl chloride. For example, PhCECSiMe3 is prepared as follows. PhCECH

n-BuLi

PhC=CLi

-

Me3S iCI

PhCrCSiMe3

(1.45)

The reaction of trimethylsilylacetylene over various solid base catalysts leads to the metathesis of the alkyne, the products being bis(trimethylsily1)acetylene and acetylene, as shown in Table 1.6. l o 2Me3SiCECH

-

Me3SiC=CSiMe3

+

HCECH

(1.46)

The basic sites abstract a proton from Me3SiC= CH to form the alkynide ion, Me3SiC=C-, which attacks at the silicon atom of another molecule of the reactant . The reaction is very selective, no other products being observed. The yield over KNH2/A1203 and KF/A1203 is limited by the equilibrium. The partial removal of acetylene from the system gave a higher yield of Me3SiC= CSi Me3 (87%) over KF/A1203 at 273 K. Furthermore, the reaction of Me3SiC = CSiMe3 and HC = CH gave Me3SiC= CH over KNH2/A1203. Metathesis of Et3SiC = CH also proceeded over KF/A1203 or KNH2/A1203 to selectively afford Et3SiC = CSiEt3. The yields of Et3SiC = CSiEt3 in benzene were 84 and 47% over KF/A1203 and KNH2/A1203, respectively, at 333 K in 2 h.”’ 2Et3SiC=CH 2Me2(EtO)SiC=CH

-

-

EtsSiC=CSiEt3

+

HC=CH

Me2(Et0)SiC-CSi(OEt)Me2

+

(1.47) H C ECH

MeZ(Et0)SiC = CH also undergoes metathesis over these catalysts. The main product was Me2(Et0)SiC = CSi(EtO)Me2. The Mez(Et0)SiC =CH gave

30

Catalysis

Table 1.6 Catalytic activities of various solid base catalysts for the metathesis of MejSiC= CHI Catalyst

Pretreatment

Yield of Me3SiCr CSi1kfe3~

Alkaline compound supported on A1203 /mmol g-alumina-

’

KF/A1203 KNH21A1203 MgO C SOH/A1203 KOH/A1203 KzC03/A1203 CaO

(Yo) 77 76 74 74 64 11 2

5

673 K, 3 h 573 K, 1 h 773 K, 3 h 673 K, 3 h 673 K, 3 h 673 K, 3 h 998 K, 3 h

2.6

5 5 5

Reaction conditions: 293 K, 30 min, catalyst weight: 0.25 g, Me3SiC=CH: 13.5 mmol. yields were calculated on the basis of Me3SiC= CH.

a

The

Table 1.7 Cross-metathesis of alk-1-ynes with MejSiC= CH.llO*a R - C = CH + Me3SiC = CH -+ R-C = SiMe3 + HC =CH Reactant

Conc /mmol

Catalyst

Reaction temp./K

Yield of R-C= SiMejb (”/)

Ph-C = CH Ph-C = CH Ph-C = CH Ph-C = CHc t-Bu-C CH n-Bu-C = CH

9.0 9.0 9.o 4.5 8.0 8.7

KF/A1203 K2C03/A1203 KNH2/Al203 KF/A1203 KF/A1203 KF/A1203

318 3 18 3 18 318 303 318

93 91 75 81 91 87

Conditions: catalyst = 0.125 g, R-C =CH/Me3SiC= CH = 2 : 1, reaction time = 2 h. The yield of R-C =CSiMe3 was calculated on the basis of Me3SiC= CH. R-C = CH/Me3SiC= CH = 1 : 1 .

a

77 and 61% yields over KNH2/A1203 and KF/A1203, respectively, in 1 h at 313 K. 6.4.2 Cross-metathesis of Alkynes - Cross-metathesis proceeds when alk- 1-ynes are reacted with silyl acetylene as shown in Table 1.7.l lo R-CZCH

+

HCECSiMe3

-

R-CECSiMe3

+

HCrCH

(1.48)

R = Ph, f-Bu and n-Bu

KFIA1203 and K2C03/A1203 gave R-C -= CSiMe3 with high yields, whereas KNH2/A1203 showed lower catalytic activity than KF/A1203 and K2C03/A1203. 6.4.3 Reaction of Alk-1-yne with Silane - The reaction of t-butylacetylene (tBuCrCSi(H)Et,) with diethylsilane in the presence of KF/A1203 at 303 K affords the silylacetylene and hydrogen. l o

I : Strong Solid Bases for Organic Reactions +

t-BuC=CH

H2SiEt2

-

31 t-BuC=CSi(H)Et2

+

(1.49)

H2

The yield of t-BuC =CSi(H)Et2 was 83% in 2 h. Dehydrogenative coupling reactions between monosubstituted silanes and alk- 1-ynes are catalysed by solid bases. l 2 For example, phenylsilane reacts with alkynes to produce monoalkynylphenylsilane, dialkynylphenylsilane and trialkynylphenylsilane in the presence of MgO. PhSiH3 + HCGCR

-

Ph(RCEC)SiH2 + Ph(RC32)2SiH + Ph(RC=C)3Si

(1 S O )

R = Ph or butyl

The reaction of phenylsilane and m-diethynylbenzene leads to formation of a polymer.

These polymers are soluble in benzene and THF, and showed very little weight loss in thermal treatment under argon. The high heat-resistance property of the polymer was attributed to a crosslinking reaction involving the Si-H bonds and C-C bonds. 6.4.4 Reaction of Alk-I-ene with Diethylsilane - KNH2/A1203 catalyses the reaction of hex- 1-ene with diethylsilane. C ~ H T C H ~ C H = C H+~ Et2SiH2

-

C3H7CH=CHCH2SiEt2H

+

H2

(1.52)

When 0.2 g of the catalyst was used, KNH2/A1203 gave C3H7CH = CHCH2SiEt2H in 22% yield in 20 h at 329 K. The amounts of hex-1-ene and diethylsilane were 24 mmol and 1.5 mmol, respectively. The ratio of 2 :E in the product alkene was 77 :23. The isomerization of hex-1-ene to hex-2-ene and hex-3-ene also proceeded, the yields of hex-2-ene and hex-3-ene being 6.6% (2:E = 62 :38) and O.6%, respectively, as hex- 1-ene base. 6.4.5 Reaction of Toluene with Diethylsilane - Because of the strong basicity of

KNH2/A1203the catalyst can activate toluene, whose pK, value is 35. The reaction of diethylsilane with excess toluene over KNH2/A1203 at 329 K affords diethylbenzylsilane in 74 and 85% yields after 20 h and 40 h,

8CH3

Et2SiH2

+

Et Et-Si-CH2

I

H

(1.53)

32

Catalysis

Table 1.8 Reactivates of alkylbenzene with EtzSiH2 over KNH2IAI203 for the formation of the corresponding benzylsilanes9*" Reactant

Toluene Ethylbenzene Propylbenzene Isopropylbenzene

PKO

Amount of reactantlmmol

35

28

-

25

-

21 29

37

Yield ("h)

74 23 7.0 2.0

Conditions: catalyst weight = 0.20 g, reaction temperature = 329 K, reaction time = 20 h, EtZSiH, = I .5 mmol, amount of K metal = 2.6 mmol per 1 .OO g of A1203.

a

respe~tively.~ RbNH2/A1203and K2C03/A1203also afford diethylbenzylsilane in 73 and 30% yields, respectively, at 303 K after 20 h. The reaction of Et2SiH2 with other alkylbenzenes, such as ethylbenzene, also proceeded in the presence of KNH2/A1203 to selectively afford the correEthylbenzenegave a 23% yield of sponding benzylsilanes, as shown in Table 1 diethyl-1-phenylethylsilane. Propylbenzene and isopropylbenzene reacted with Et2SiH2 to give diethyl(1-phenylpropyl)silane and diethyl(1-phenyl-1-methylethy1)silane in 7 and 2% yields, respectively. As shown in Table 1.8, alkylbenzenes with more acidic protons showed higher reactivity towards Et2SiH2, indicating that the abstraction of a proton from alkylbenzene by basic sites on the solid base is the rate-determining step. The reactions proceed through carbanion intermediates, which causes nucleophilic substitution at silicon. Benzene has a pK, value of 37, comparable to propylbenzenes and therefore is expected to undergo a similar reaction. When Et2SiH2 (3.1 mmol, 0.4 cm3) was stirred with 1.0 g of KNH2/A1203 in benzene (68 mmol, 6 cm3), diethylphenylsilanewas obtained in a 7.5% yield in 20 h at 329 K9

(1.54)

6.5 Miscellaneous Reactions 6.5.1 Tishchenko Reaction - The Tishchenko reaction is a dimerization of aldehydes to form esters. Thus, benzaldehyde is converted into benzylbenzoate. Alkaline earth oxides are very effective catalysts for this reaction. Among them, BaO is most active. KF/A1203, when pretreated at 623 K, is also a very effective catalyst for the Tishchenko reaction of benzaldehyde, and gives benzyl benzoate in 94 and 990/0yields in 3 and 20 h, respectively, at 323 K.66 The reaction presumably proceeds via following mechanism (Scheme 5):

I : Strong Solid Bases for Organic Reactions

33

Scheme 5

6.5.2 Reactions of Silanes with Ketones or Aldehydes - Cyanotrimethylsilane reacts with ketones or aldehydes to form addition compounds in the presence of solid bases. The effective catalysts include MgO, CaO, CaF2, CaSi03, hydroxyapatite and K3P04.36 @ 3 i O

+

Me3SiCN

-

(1.55) CN

Hydrosilylation occurs by the reaction of triethoxysilane with a variety of ketones or aldehydes in the presence of CaO or hydroxyapatite at 363 K. l4 Benzaldehyde reacts with Et3SiH and PhMe2SiH to afford the corresponding silyl benzyl ether in the presence of KF/A1203 in 93 and 99% yields, respectively, at 303 K in 1 h. lo O

C

H

O

+

R’ R2-2i-H

-

R’

I

@H20-;i-R2

R3

A3

(1.56)

R’ = R2 = R3 = Et R’ = Ph, R2 = R 3 = Me

However, MgO and CaO did not show the catalytic activity for the reaction of benzaldehyde with Et3SiH or PhMeZSiH at all. Acetophenone also reacts with PhMe2SiH over KF/A1203 to afford ( 1-phenyl-1-ethoxy)dimethylphenylsilane, whose yields were 43 and 54%after 20 h at 303 and 353 K, respectively. l o 0 II

C Ph’ ‘CH3

+

H-Si

-

CH3

I

H

CH3

6.5.3 Reaction of Aniline with Trimethylsilylacetylene - When aniline was reacted with trimethylsilylacetylene at 318 K for 20 h, N-trimethylsilylaniline was formed in a 90% yield in the presence of Mg0.5 0

N

H

2 + Me3SiCFCH

-

e N H - S i M e 3 + HCFCH

(1.58)

Catalysis

34

K2C03/A1203, KNH2/A1203, KF/A1203 also showed catalytic activities, the yields of N-trimethylsilylaniline being 80, 76 and 67%, respectively, under the same reaction conditions. On the other hand, KNH2/A1203 showed a low catalytic activity, the yield of N-trimethylsilylaniline being 47% at 323 K in 20 h. 6.5.4 Ring Opening of Epoxides with M e $ X N - CaO and MgO are effective for nucleophilic ring opening of epoxides with Me3SiCN to afford P-trimethylsiloxynitriles in a high regioselective manner. Epoxides quantitatively produce P-trimethylsiloxynitriles by highly regioselective (>94%) attack of cyanide on the less substituted carbon of epoxides without any formation of isocyanides.

-

R'

CN

R218iXoi~R3 Me3SiO H

(1.59)

Ring opening of 2,3-epoxyhexan-1-01s with Me3SiCN also takes place on CaO in high yields and regioselectivities. * l5 6.5.5 Hydrocyanation - Kabashima and Hattori examined the catalytic activities of various solid bases for cyanomethylation of methanol with acrylonitrile to form 3-methoxypropanenitrile at 323 K.'16 The alkaline earth oxides and KOH/A1203 showed the highest activity. Alkaline earth hydroxides also show the high activity. The catalytic activity of MgO for this reaction was not deteriorated by exposure to air. Kumbhar showed that Mg-A1 hydrotalcite modified by thermal decomposition followed by rehydration is highly active for cyanoethylation of alcohols with acrylonitrile. l 7 The catalyst is also air stable. Therefore, active sites for the hydrocyanation are certainly different from those for alkene isomerization and could be basic hydroxide groups rather than oxide ions. 6.5.6 Alkylations (a) Side Chain Alkylation of Alkylaromatics. Alkylation of alkylbenzenes with alkenes or alcohols over base catalysts yields the products alkylated at the side chain, whereas ring alkylation proceeds over acidic catalysts. To abstract a proton from the alkyl groups, strongly basic catalysts and/or severe reaction conditions are required, as expected from their pK, values. In the reaction of toluene with methanol, alkali-ion-exchanged zeolites, especially, RbX and CsX, give ethylbenzene and styrene as products, whereas acidic zeolites afford ~ y l e n e sThe . ~ ~reaction requires a temperature of 750 K. Strong solid base catalysts such as N a N 3 / ~ e o l i t eand ~ ~ Na/NaOH/A120343 catalyse the side-chain alkylation of alkylaromatics with alkenes under mild conditions. Alkylation of o-xylene with butadiene over Na/K2C03 yields 0tolylpent-2-ene, which is a precursor for 2,6-dimethylnaphthalene.I (b) Alkylation of Phenols. Alkylation of phenol with mehanol over solid base

I : Strong Solid Bases for Organic Reactions

35

catalysts occurs at ortho-positions to afford o-methylphenol and 2,6-dimethylphenol, the latter being a monomer for a heat-resisting resin. The favorable catalysts are MgO and Mg-containing oxides.'Y2 On the other hand, acidic catalysts give all three isomers of cresols. The reaction of catechol with dimethyl carbonate proceeds over a series of alkali metal salts supported on alumina at 583 K. The product distribution depends very much on the catalyst used. KN03/A1203 showed the highest activity to gave veratrole, 14, in a high yield (98%),17 and CsOH/A1203 gave guaiacol and catechol carbonate, 13, in almost 1:1 ratio.I6 A weak base, LiOH/ A1203,gave guaiacol, 12, in a high selectivity (84%).15 H3C0,

C

,0CH3

b

(1.60)

6.5.7 Formation and Ring Transformations of Heterocycles - 4-Methylthiazole, a fungicide, can be synthesized by a vapor-phase reaction. l7 Thus, Cs-loaded ZSM-5 zeolite gives excellent performance, e.g. activity, selectivity, and lifetime. The reaction is run at 700 K in the presence of water vapor. (1.61) H3C'

The ring transformation of y-butyrolactone into y-butyrothiolactone proceeds over alkali-ion-exchanged faujasite. A 100% conversion was obtained over CsY at 603 K. + H2S

7

-

Qo

+

H20

(1.62)

Concluding remarks - Unsolved Problems

Use of solid base catalysts in organic synthesis has been increasing in recent years because of strong demand for environmentally benign processes. The methods of preparing solid base catalysts and the characterization of the surfaces have also been developed. Strongly basic solid catalysts will further extend the area of the reactions in various fields of chemistry. However, the nature of basic sites is not always clear. For example, the discussion about the active sites of KF/A1203 remains controversial. Though

Catalysis

36

KN03/A1203or K2C03/A1203are strong bases and very useful catalysts, the origin of the basic property is not fully understood. In many cases, use of A1203as a support is essential to generate high catalytic activities for various reactions. Furthermore, their catalytic activities increase enormously upon heat treatment under vacuum. It is not clear whether the heat treatment is effective simply because of elimination of adsorbed molecules such as H20 or CO2, or whether it leads to development of new active sites by the chemical interaction of alkaline metal salts and the support. At the moment, it is essential to describe the pretreatment as well as preparation conditions to be clearly identified. Base-catalysed reactions are often carried out in the liquid phase. Because of slow diffusion of reactants and products, catalysts having large pores as well as high surface areas are very desirable. There have been several attempts to create basic sites in mesoporous materials and their activities for Knoevenagel and Michael reactions were tested.' Similar catalysts which are also useful for less acidic reactants (higher pK, values) such as alkenes are certainly needed. In general, H20 and/or CO2 easily poison solid base catalysts. It is an intriguing problem to prepare a water- or carbon dioxide-resistant solid base. For example, creating basic sites in hydrophobic micro- or mesoporous environments may be of interest. More work is needed for developing shape-selective and stereoselective solid base catalysts.

References 1 2

7 8 9 10 11 12 13 14 15

K. Tanabe and W.F. Holderich, Appl Catal., A, 1999,161,399. K. Tanabe, M. Misono, Y. Ono and H. Hattori, New Solid Acids and Bases (Stud. SurJ Sci. Catal., 51), Kodansya-Elsevier, 1989. H. Hattori, Chem. Rev., 1995,95,527. T. Baba, H. Handa and Y. Ono, J. Chem. Soc., Faraday Trans., 1994,90, 187. Y. Ono and T. Baba, Catal. Today, 1997,38,321. H . Handa, T. Baba, H. Yamada, T. Takahashi and Y. Ono, Catal. Lett., 1997, 44, I 19. T. Baba, A. Kato, H. Handa and Y. Ono, Catal. Lett., 1997,47, 77. T. Baba, A. Kato, H. Takahashi, F. Toriyama, H. Handa, Y. Ono and H. Sugisawa, J. Catal., 1998,176,488, T. Baba, H. Yuasa, H. Handa and Y. Ono, Catal. Lett., 1998,50,83. T. Ando, Stud Surl: Catal., 1994,90,9. J.H. Clark, D.E. Cork and M.S. Robertson, Chem. Lett., 1983, 1145. D.E. Bergbreiter and J.J. Lalonde, J. Org. Chem., 1987,52, 1601. H. Tsuji, H . Kabashima, H. Kita and H. Hattori, React. Kinet. Catal. Lett., 1995, 56, 363. T. Yamaguchi, J.-H. Zhu, Y. Wang, M. Komatsu and M. Ookawa, Chem. Lett., 1997,989. Y. Fu, T. Baba and Y. Ono, Appl. Catal., A, 1998,166,425.

I : Strong Solid Basesfor Organic Reactions 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34

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61 62 63

64 65 66 67 68 69 70 71 72 73 74 75 76 77 78 79 80 81 82 83 84 85 86 87 88 89 90 91 92

I : Strrong Solid Basesfor Organic Reactions 93 94 95 96 97 98 99 100 101 102 103 104 105 106 107 108 109 110 111

112 113 114 115 116 117 I18

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2 Catalysis by Solid Bases BY ERIC J. DOSKOCIL, SHAILENDRA BORDAWEKAR AND ROBERT J. DAVIS

1

Introduction

Solid acids have received significant attention in the literature due to their industrial importance in replacing liquid acids in petroleum processing. Hence, a great deal of research has involved determining the nature and strength of the active sites on these materials. In contrast, relatively few studies involve solid bases, despite their potential as important industrial catalysts. Tanabe and Holderich performed a statistical survey of industrial processes using solid acids, solid bases and acid-base bifunctional catalysts, and have counted 103, 10 and 14 industrial processes for these types of catalysts, respectively.' Clearly, solid acids have received most of the attention in industrial applications. However, as more novel base materials become known and new basecatalysed reactions are found to be commercially relevant, additional studies of solid bases will be necessary for these catalysts to achieve the industrial success of solid acids. Solid base catalysts exhibit high activities and selectivities for many kinds of reactions, including some condensations, alkylations, cyclizations and isomerizations which are carried out using liquid bases as catalysts in industrial applications.2 Many of these applications require stoichiometric amounts of the liquid base for conversion to the desired product. Replacement of these liquid bases with solid base catalysts would allow for easier separation from the product as well as possible regeneration and reuse. These basic solids have the added advantages of being non-corrosive and environmentally friendly, which allows for easier disposal. Significant interest in strong solid base catalysts began shortly after the early investigations of Pines that involved sodium metal doped onto a l ~ m i n a . ~ Strong solid bases are now recognized as effective catalysts for a wide range of organic transformations such as double-bond isomerization of olefins, sidechain alkylation of aromatics, dehydrogenation of alcohols and Knoevenagel condensation of A molecular level understanding of solid basicity is required before structure-function properties of new materials can be effectively predicted. A significant fraction of the research on solid. bases therefore involves the correlation of base strength to catalyst composition. However, strong bases are also poisoned by carbon dioxide and water, which are common side products Catalysis, Volume 15 0The Royal Society of Chemistry, 2000 40

2: Catalysis by Solid Bases

41

in catalytic reactions. New base catalysts that are more resistant to deactivation by these molecules need to be developed. In addition, the search for novel solid bases that catalyse transformations with high product selectivity, high reaction rate and low deactivation rate is an ongoing process. This chapter discusses various solid base catalysts currently of interest to researchers. In addition, commonly used characterization methods and catalytic probe reactions are also presented.

2

Types of Basic Catalysts

2.1 Supported Alkali Metal Oxides - Alkali metals and metal oxides are among the strongest bases known. Indeed both rubidium and cesium oxide have been reported to be 'superbases', possessing a Hammett basicity function exceeding +26 (see Section 3).9 These oxides have a very low surface area that limits their ability to be effective catalysts. Therefore, supporting these oxides on higher surface area carriers is expected to give materials exhibiting this strong basicity as well. Because cesium has been shown to create stronger basic centers than other alkali metals in Group IA, significant attention in the literature has focused on supported cesium catalysts. In addition to the alkali metal oxides mentioned above, alkali metals supported on basic alkaline earth oxides like MgO and CaO have been found to be effective solid base catalysts. These materials can be prepared by heating the carrier in the presence of a vapor-phase alkali metal.lO*llDoping of various alkali metal salts on MgO also produces both basic and superbasic catalysts used for the oxidative coupling of the aldol condensation of acetone,21 the oxidative methylation of acetonitrile with methane,22 and the decomposition of but-2-an01 and 2-methyl-3-buten-2-01.~~ Alkali metal salts doped onto other oxide supports are also known to be solid bases. The active materials are prepared by heating the sample to high temperatures to decompose the supported alkali precursor. Shen et al. have determined the optimal decomposition temperatures for various alkali metal nitrates on y-A1203in order to form a layer of surface alkali metal oxide with minimal migration of the alkali metal into the Alkali metal doped onto alumina has also been used for COS h y d r o l y ~ i s ,isomerization ~~ of cis-but-2ene,26-27 and the synthesis of catechol carbonate.28Supported rubidium acetate has been decomposed on a variety of common oxide carriers of various acidity (magnesia, titania, alumina and silica) to study the effect of the support on the resulting alkali metal species.29 Not only did rubidium reduce the support acidity, but it also increased the density of base sites. Additional studies on strontium-modified supports, prepared via the decomposition of impregnated strontium acetate, showed that samples containing Sr had weaker base sites than materials containing similar loadings of Rb.30 2.2 Hydrotalcites - There has been an increased interest in the use of mixed oxides both as basic catalysts and supports. Hydrotalcites, or Mg-A1

42

Catalysis

Table2.1 Turnover rates over MgO, Mg: A1 mixed oxides and A1203 (adapted from reJ:31) Catalysts

Acetone formationb (10-3 s-1)

Turnover rates Propeneformationb (10-4 s-1)

But-I -ene Isomerization'

(s-9

~~

MgO Mg:Al 5 : 1 Mg:A1 3: 1 Mg:Al 2: 1 A1203 a

2.6 3.3 2.3 2.7 1.9

Calcined for 4 h at 823 K in flowing He. Reaction temperature was 340 K.

0.82 2.1 4.4 11.5 750

0.22 0.18 0.10 0.093 0.0071

Propan-2-01 decomposition temperature was 593 K.

hydroxycarbonates, are layered double hydroxides that form Mg-A1 mixed oxides upon calcination. The acidity and basicity of these mixed oxides have been characterized by catalytic probe reactions like the decomposition of propan-2-01 (see Section 4.2) and isomerization of b~t-l-ene.~' The effects of Mg:Al atomic ratio on turnover rates for these two reactions are given in Table 2.1. Since turnover frequencies are based on the adsorption capacity of carbon dioxide (a common probe for basic sites), the rate of a structure insensitive base-catalysed reaction should be relatively constant at a given temperature and pressure. The dehydrogenation of propan-2-01 to acetone requires the presence of a base site. Indeed, the invariance of the turnover frequency for acetone formation from propan-2-01 over MgO, Mg :A1 mixed oxides and A1203 supports the idea of structure insensitivity. In contrast to the dehydrogenation results, the turnover rate for propan-2-01 dehydration increased by three orders of magnitude as the catalyst was changed from magnesia to alumina. Apparently, the surface sites counted by C02 adsorption do not function equally for alcohol dehydration. This is expected since dehydration is usually considered an acid-catalysed reaction. In addition, the mechanism for secondary alcohol dehydration over basic oxides differs from that over alumina. It is of interest to note that the turnover frequency of base-catalysed isomerization of but- 1-ene (Table 2.1) was also a strong function of composition, with magnesia being the most active. Adsorption of C02 most likely does not provide an accurate measure of active isomerization sites on the mixed oxides. Apparently, the ensemble of surface atoms required for butene isomerization is larger than that associated with CO2 adsorption.

2.3 Zeolites and Mesoporous Oxides - Other mixed oxides containing Al, Si, Mg and Ti have been prepared with varying degrees of acidity and basicity. Some of these mixed oxide materials can be formed into highly porous materials with complex three-dimensional structures. These include zeolites and mesoporous mixed oxides like MCM-41. The aluminosilicate form of MCM-41 is an amorphous mixed oxide that has been templated with large

.

2: Catalysis by Solid Bases

43

Figure 2.1 Schematic of cage structures of (a) X zeolite and (b) occluded molecules in an X zeolite supercage

surfactant molecules during synthesis to yield a solid with highly oriented unidimensional mesopores. Zeolites are high surface area, microporous materials consisting of interconnected Si04 and A104 tetrahedral units linked through one oxygen. The linked tetrahedral units form a three-dimensional aluminosilicate that contains interconnected pore structures. These structures are typically formed from the linking of cages and prisms composed of the tetrahedral units. A threedimensional representation of zeolite X can be seen in Figure 2.l(a) where the vertices represent the location of Si and A1 atoms. Oxygen atoms are omitted for clarity. For each A104 tetrahedron present in the zeolite framework, a chargecompensating cation, which is typically a proton or metal cation, is necessary to balance the charge since A13+ has a lower valency than Si4+.The existence of both positive and negative charges accounts for the acid-base character of the zeolite structure. It is easy to see that increasing the fraction of A104 tetrahedral units present in a zeolite increases the net negative charge associated with the framework, which results in a zeolite containing more basic sites. Some of the counter-cations in the zeolite pores can be readily exchanged with other cations, which may alter the pore-size of these materials as well as change the acid-base character of the zeolite framework. Exchanging zeolites with a less electronegative charge-balancing cation such as cesium creates a more basic zeolite. Ion-exchanged zeolites have been used to catalyse many reactions known to occur on a base site, such as toluene alkylation with methanol. In addition, cation-exchanged mesoporous Al-Si mixed oxides like MCM-41 have been used for the base-catalysed Knoevenagel condensation of benzaldehyde and ethyl ~ y a n o a c e t a t e . ~ ~ Occlusion of alkali metal oxide clusters in zeolite cages via decomposition of impregnated alkali metal salts results in a further increase in the basicity of these r n a t e r i a l ~ . ~ The . ~ * supported ~ ~ - ~ ~ species are typically introduced through wet impregnation of a solution containing the solvated precursor into the zeolite pores. The resulting zeolite contains the occluded alkali metal precursor in the pores as seen in Figure 2.l(b).

44

Catalysis

A significant contribution to the understanding of base catalysts involving supported alkali metal species has been achieved with these materials. However, the nature of the occluded alkali species has been elusive, and the actual form of the occluded oxide is still in question. Recently, NMR spectroscopy has been used to study cesium species loaded into zeolites in excess of ion-exchange c a p a ~ i t y . In ~ ~addition, .~~ an NMR study of a variety of alkali metal oxides, superoxides and peroxides suggests that the peroxide and superoxide phases should be considered for their role in catalytic reactions as both a base and an oxidant.40 Kloestra et al. have successfully created a re-usable catalyst consisting of binary cesium-lanthanum oxide supported on MCM-41 to carry out the Knoevenagel addition of enolates to benzaldehyde in aqueous media.41 Zeolites containing alkali metals in the cages have been investigated as high strength solid base catalysts. In addition, a novel preparation technique has been introduced in the literature in which alkali metal clusters are prepared via the decomposition of alkali metal azides in the presence of a z e ~ l i t e . ~ ~ - ~ O Unlike their alkali metal counterparts, the azides themselves are stable at room conditions. The catalytic properties of azide-supported zeolites containing ionic and/or metal clusters of alkalis have been investigated using the isomerization of b ~ t - l - e n e . ~ ~Decomposition *~**~~ of the alkali a i d e in the pores of the zeolite forms either ionic or neutral alkali metal clusters, depending primarily on the heating rate used to decompose the supported azide; faster heating rates (ca. 25 K min-I) result in ionic clusters whereas slower heating rates (ca. 1 K min-') give primarily neutral metal clusters.49Since both of the alkali clusters can be formed upon introduction of an alkali metal to zeolite, it is important to determine the chemical state of the alkali metal when studying its role in catalysis. Electron spin resonance (ESR) spectroscopy has been used to determine whether alkali metal or ionic clusters are present in the cages of X and Y ~ e o 1 i t e ~ . ~ ~ - ~ ~ ~ ~ * ~ ~ ~ ~ ~ 2.4 Alkali Metals Supported on Nanoporous Carbon - Recently, Stevens and co-workers developed strong basic catalysts obtained from entrapping welldispersed cesium in nanoporous carbon (NPC).53-57 By incorporating Cs metal through vapor deposition, a catalyst containing atomically dispersed Cs in the nanopores is created (Cs/NPC). Macropores provide for easy transport to and from the active alkali metal species. These materials were shown to be thermally stable, not pyrophoric, and active for a variety of reactions including butene isomerizati~n,~~ benzene coupling,5s aldol condensations6 and sidechain alkylation of toluene with p r ~ p e n e Magnetic .~~ susceptibility and EPR measurements indicated the presence of unpaired electrons in these materials and led to the speculation that electrons from cesium are donated to the carbon s ~ p p o r t . Based ~ ~ * ~on~ these findings, the authors conclude that the chemistry over these catalysts proceeds through surface anions rather than radicals and the catalyst functions more like a Lewis base.ss*s7 Since cesium is atomically dispersed within the nanopores of NPC, Stevens et al. investigated these materials after oxidation of the supported alkali

2: Catalysis by Solid Bases

45

The catalytic activity and X-ray diffraction patterns of these materials were compared to NPC catalysts prepared via impregnation of cesium hydroxide. The X-ray diffraction lines did not correspond to the X-ray patterns for the hydroxides, carbonates or oxidess6 In addition, the catalysts prepared with aqueous impregnation of CsOH were less active and selective for the basecatalysed condensation of acetone to isophorone than the oxidized Cs/NPC catalyst. The controlled oxidation of a supported alkali metal, like that for Cs/ NPC, provides a new synthesis technique for the preparation of supported alkali metal oxides than the decomposition of supported alkali salts mentioned previously.

2.5 Basic Phosphate Oxynitride Catalysts - A relatively new solid base has been prepared by the activation of aluminophosphate oxide in high temperature ammonia to form aluminum phosphate oxynitride (AlPON).s8-60 The amount of nitrogen incorporated into the oxynitride depends upon the length of time and concentration of ammonia flow as well as the nitridation temperature? The same synthesis technique has been used to prepare zirconophosphate oxynitrides (ZrPONs), aluminovanadate oxynitrides (VAlPONs) and galloaluminophosphate oxynitrides (AlGaPONs). Two types of nitrogen-containing species exist on the surface and in the bulk of these catalysts. The first type is surface NH, species (1 5 x 5 4) which are formed after ammonia interacts with surface hydroxyls. The second type is bulk nitride anions, N3-, which exchange with the oxygen atoms in the parent oxide. The acid-base nature of these materials is dependent upon the nitrogen :oxygen ratio; both the number of base sites and the base strength increase with an increase in the amount of nitrogen substituted into the oxynitride structure.mv62 Fripiat et al. suggest that nitridation follows a three-stage mechanism in which flowing ammonia molecules initially react with the surface. This is followed by substitution of ammonia into terminal hydroxyls to form metal amino groups and, ultimately, diffusion of nitrogen into the bulk structure.62Samples having surface areas in excess of 200 m2 g-' can be prepared with this method. Since the basicity of the surface of these phosphate oxynitrides is controlled by the nitrogen content, the basic properties of the solid can easily be adjusted. Understanding the role of nitrogen in the creation of base sites is important to synthesize effective solid base catalysts.

3

Characterization of Solid Base Catalysts

From the classical Lewis definition, the base strength of a solid catalyst is determined by its ability to donate an electron pair to an adsorbed molecule. Typical measurements of basicity have been obtained previously by using titration of adsorbed indicators having a wide range of pK, values. For a reaction of an acid indicator BH with a solid base B*, BH + B* g B-

+ B*H+

(2.1)

Catalysis

46

the Hammett basicity function, H-,is defined by

where [BH] is the concentration of the indicator and [B-] is the concentration of its conjugate form. One problem with using adsorbed indicators to evaluate basicity is the interference of indicator reactions that are not due to acid-base chemistry. In addition, evidence of reaction is often provided by a color change, which requires the use of colorless catalysts. Clearly, there is a need for other methods to probe surface basic sites. 3.1 Theoretical Ranking of Solid Basicity 3.1.1 Sanderson Intermediate Electronegativity - Another way to arrive at a scale for solid basicity is to calculate the partial negative charge on the basic center. For metal oxides, the basic site is likely to be a surface oxygen anion. The challenge is to calculate the electronic charge associated with the basic surface oxygen. One very simple method that is often used to compare structurally similar compounds is based on the intermediate electronegativity principle of Sanderson. The Sanderson electronegativity (SE,int) of a solid is calculated from the mean of the individual electronegativities of the elements present, where SE,zis the individual electronegativity of element 2 and x is the atomic ratio of element Z present. The basic idea is that an equalization of the electronegativities in a compound results from electron transfer that occurs in the formation of the compound.63The calculation of the intermediate electronegativity takes only the composition of the compound into account, ignoring the effect of structure and surface composition. From the Sanderson intermediate electronegativity, the partial negative charge on oxygen can be calculated by using the following equations: ASE.0

= 2.08.(SE,o)1n

(2.4)

High partial negative charge on oxygen indicates strong basicity. Since the oxygen is the basic site on oxides and mixed oxides, many basic materials have been ranked using the partial negative charge on ~ x y g e n .For ~ example, Barthomeuf has shown that partial negative charge on oxygen increased with increasing aluminum content of zeolites.64In addition, Mortier found that the partial negative charge on a zeolite framework oxygen increased with increasing electropositivity of the alkali metal counter-cati~n.~~ 3.1.2 Molecular Orbital Calculations - The use of molecular orbital calculations to understand solid catalysts has been receiving greater attention in the

2: Catalysis by Solid Bases

47

literature. However, many of the model systems that have been studied are aimed at understanding solid acidity instead of solid basicity. Surfaces are typically represented by small three-dimensional clusters of atoms or periodic arrays of atoms. Various probe or reactant molecules are then brought into close proximity to simulate adsorption. The binding energy of the adsorbate molecule is determined by subtracting the energies associated with the initial bare cluster and free molecule from that of the final complex. High binding energies for acid probe molecules are typically interpreted as an indication of strongly basic adsorption sites. From these calculations, a fundamental understanding of heterogeneous base catalysis can be investigated at a molecular level. This field is still in its infancy and will surely grow in the near future. Many of the ab initio calculations involving solid bases deal with the alkaline earth oxides MgO and Ca0.66v67Using proton adsorption to probe MgO clusters, Kawakami and Yoshida found that the fewer Mg atoms coordinated to a central basic oxygen and more 0 atoms coordinated to the adjacent Mg atoms created stronger base sites.66 Pacchioni et al. used (0Mg5)8+ and (0Ca5)8' clusters to investigate COZ and SO2 adsorption on MgO and Ca0.67 They found that both of these probe molecules, which are acidic, bound more weakly to MgO and more strongly to CaO. The differences in basicity were attributed to the electrostatic stabilization of the 02-ion at the surface by the Madelung potential of the ionic crystal.67For alkali metal compounds, which are known to be strong bases, Burk and Koppel performed ab initio calculations to predict the geometries and proton affinities of lithium, sodium and potassium oxides and hydroxides, confirming that the alkali metal oxides were indeed more basic than the corresponding hydroxides.68

3.2 Experimental Ranking of Solid Basicity 3.2.I Infrared Spectroscopy of Adsorbed Probe Molecules - As mentioned above, a measure of basicity is often obtained by using titration of adsorbed indicators over a wide range of pK, values. Since these measurements are carried out in the liquid phase, it can be more advantageous to investigate the adsorption of gaseous acidic probe molecules to determine strengths of adsorption sites relevant to catalysis. Infrared spectrometric studies of various probe molecules adsorbed on metal oxides and zeolites have been reviewed by L a ~ a l l e y The . ~ ~ basic sites in alkali metal-modified zeolites are commonly characterized by IR spectroscopy coupled with temperature programmed desorption of adsorbed probe molecules like carbon d i o ~ i d e , ~ * ~ O - ~ ~ and c h l ~ r o f o r r n . ~ ~ - ~ ~ The criteria for selecting appropriate probe molecules for investigation of surface base sites using IR spectroscopy has been reviewed p r e v i o ~ s l y . ~ ~ ~ ~ ~ * * One of the major problems with this technique is that many of the commonly used probe molecules decompose or react upon interaction with the basic site and, therefore, do not effectively interrogate the catalyst surface. For example, pyrrole has been found to dissociatively chemisorb on highly basic metal oxides like Tho2 and CeOz, forming the C4H4N- pyrrolate anion.69Thus, a

48

Catalysis

judicious choice of probe molecule is important in characterizing surface basicity. A probe molecule should undergo specific chemical interactions with the base sites without dramatically altering the catalyst surface. Thus, the search for a unique, widely applicable adsorbate molecule to probe the surface base sites of heterogeneous catalysts is a daunting process. No single probe is universally adept at characterizing the active sites on all basic solids. Carbon monoxide is an effective probe for the hydroxy groups of metal oxides. The OH groups of metal oxides like MgO and CeOz are highly basic and are not affected by CO adsorption at 78 K.69 These OH groups are bonded to only one metal atom. However, the OH groups bridging between metal atoms undergo hydrogen bonding with adsorbed CO. The Bronsted acidity of metal oxides also increases with increasing coordination number of the OH groups. Thus, low temperature IR spectroscopy of adsorbed CO can be used to characterize the acid-base properties of metal oxides. Carbon monoxide adsorption has also been used to interrogate the basic 02-sites on metal oxides. Pyrrole adsorption has been found to be useful for probing the basicity of zeolites. An increase in solid base strength has been correlated to a shift in the NH vibration frequency to lower wavenumbers in the IR spectrum for -84 The NH vibration frequency is numerous alkali-exchanged generally found between 3400 and 3000 wavenumbers and can be complicated by the OH vibration frequencies of surface hydroxy species. The overlap of these spectral features is particularly troublesome for metal oxides where pyrrole forms a hydrogen-bonded species with either surface 02- or OHions. Shifts in the NH vibration frequency to lower wavenumbers have been correlated to higher strength adsorption sites for various metal o ~ i d e s . ~ ~ ~ When the 02-species is highly basic, the surface OH species are unperturbed and the H atom of the pyrrole molecule is localized near the basic oxygen, undergoing dissociative chemisorption as seen in Figure 2.2(a). When the 02species are less basic, the surface oxygen forms an NH- -0bridge with pyrrole [Figure 2.2(b)].69 Interaction with surface hydroxy groups leads to hydrogenbridged species depicted in Figure 2 . 2 ( ~ ) Complexities .~~ in the IR spectrum result from the interaction between surface hydroxys and pyrrole since hydroxy species act as both a basic surface species as well as a product formed from pyrrole dissociati~n.~~ Downward shifts to lower wavenumbers for two vR(ring) modes of pyrrole adsorbed on metal oxides has been used to rank base strength.85However, changes in the vR(ring) modes of pyrrole did not show any recognizable trend for ranking the base strength of zeolites of different basicity.74From the work using pyrrole as a probe of basic surfaces, it is evident that great care must be taken when interpreting the IR spectrum of the adsorbed species. Carbon dioxide is probably the most widely used probe for surface basicity. Since carbon dioxide is a weakly acidic molecule, it selectively adsorbs on base sites. Carbon dioxide adsorption on metal oxides and zeolites results in many different surface species, all of which are easily seen by IR spectroscopy. Figure 2.3 shows the various modes of C02 adsorption on metal oxides.87 These modes are summarized below:

2: Catalysis by Solid Bases

49

(c)

IM".

H

/ ....,

H

k

-=-.... /

Osurfacc

Figure 2.2 Adsorption of pyrrole on different strength base sites

0

ti

0-

P\ I (-M-O-M-0-M-) 0

0

t

o-c4 I 1

(-M-O-M-)

(-M-O-M-)

0

II

C ' 0 '0 I 1 (-M-O-M-)

\, 7../o

0 I (-O-Me-)

(4

t /* F (-M-O-M-) O \

(f)

Figure 2.3 Adsorption of CO, on different types of sites in metal oxidesg7

50

Catalysis

(a) Adsorption on a surface hydroxy group resulting in the formation of a bicarbonate. (b) Adsorption on a metal cation and dissociation of the resultant species. (c) Adsorption on a metal cation and the adjacent oxygen ion to form a bidentate carbonate. (d) Adsorption on an oxygen vacancy to form a surface carbonyl group. (e) and (f) Adsorption on the metal cations with participation of oxygen in excess leading to the formation of a unidentate carbonate. One or more of these surface adsorption modes may be preferred, depending on the physicochemical properties of the metal oxide, the synthesis method, starting materials and pretreatment conditions. On basic oxides, carbon dioxide chemisorbs onto a surface oxygen to form a carbonate species which exhibits vibrational stretching frequencies in the 1200- 1800 cm- range of the IR spectrum. The common forms of carbonate found on basic solids are bicarbonate (a), bidentate (c), and unidentate (e, f) carbonates. A unidentate carbonate is indicative bf a high strength base site and is formed when the carbon from COZ interacts with a basic oxygen. This unidentate species gives characteristic 0-C-0 stretching frequencies in the 15 10- 1560 cm - and 1360- 1400 cm- ranges for asymmetric and symmetric stretches, respectively. A bidentate carbonate is indicative of a medium strength base site and is formed when a carbon and an oxygen from C02 interact with a basic oxygen and a surface metal cation, respectively. This bidentate carbonate gives 0-C-O stretching frequencies in the 1610-1630 cm-I and 1320-1340 cm-' ranges for asymmetric and symmetric stretches, respectively. A bicarbonate is indicative of a low strength base site and is formed when carbon dioxide interacts with a surface hydroxy group. A bicarbonate species shows asymmetric and symmetric stretches at about 1650 and 1480 cm-', respectively, as well as a C-OH bending mode near 1220 cm- Therefore, by investigating the location of stretching frequencies for carbon dioxide adsorbed on a basic catalyst, the strength of the base sites present on the surface can be qualitatively determined. 3.2.2 Visible Absorption Spectroscopy of Adsorbed Iodine - The visible absorption spectrum of iodine has been used recently to rank donor strength of the basic oxygens in various zeolites.88Choi et al. have shown that a blue-shift occurs in the visible spectrum of adsorbed iodine with increasing electropositivity of the alkali counter-ion from Li to K, as well as with increasing aluminum content in the zeolite framework.88These observations are consistent with the ranking of zeolite basicity determined by infrared studies69and alcohol decomposition reactions.89This correlation has also been made for the blue-shift in the visible absorption band of an iodine donor-acceptor complex to its heat of formation for solvated molecules.90-92 Figure 2.4 shows a sample spectrum for zeolite KY containing adsorbed iodine. The high energy band located at short wavelength is attributed to charge-transfer upon complexation whereas the low energy band at longer

2: Catalysis by Solid Bases

51

-Charge

Transfer Band

1.o

Visible Absorption Band

n

0.5

0.0

240

320

280

360

400

440

480

520

Wavelength (nm) Figure 2.4 UV- Vis diffuse reflectance spectra for iodine adsorbed on K Y at room temperatureg5

0

5

10

15

20

25

-AW,, (kcal mol-1) Figure 2.5 Partial negative charge on oxygen as a function of energy calculated from the blue-shvt of adsorbed iodine in relation to gaseous iodineg5

wavelength is due to the visible absorption of the complexed i ~ d i n e .For ~~,~~ both X and Y zeolites, a blue-shift was observed in the low energy bands when sodium was exchanged with p o t a s ~ i u mChoi . ~ ~ et al. found similar results for analogous zeolites.88The cesium-exchanged X and Y zeolites have absorption features at almost the same wavelength as their corresponding. potassiumexchanged zeolites. Figure 2.5 shows the correlation of the partial negative charge on the framework oxygen atoms with the energy shift of the iodine absorption maximum with respect to 12 in the gas phase. The partial negative

52

Catalysis

charge on oxygen was calculated from the Sanderson intermediate electronegativity principle, which is based solely on the overall sample c o m p ~ s i t i o nA .~~ linear relationship between energy shift and partial negative charge on oxygen was observed for the alkali metal-exchanged zeolites. The incomplete levels of Cs exchange account for the similar absorption features for CsX and KX as well as CsY and KY. The use of iodine as a visible probe for donor strength has been extended to other metal oxides typically used as catalyst supports. Pure silica MCM-41 containing adsorbed iodine had a visible band at 508 nm, which is close to that observed for gaseous I2 (520 nm), indicating that the interaction with amorphous silica is relatively weak.95 Similar values have been seen as well for zeolitic materials with high WA1 ratios.88Alumina exhibited a significant blueshift of the low energy absorption peak to 410 nm.95Magnesia, a classic basic oxide, gave adsorbed I2 an even more significant blue-shift to 366 nm, confirming the high donor strength of its surface.95 The adsorption of iodine as a probe for surface basicity has been shown to be effective in determining the relative basicity for alkali metal-exchanged zeolites. Blue-shifts in the visible absorption spectra of iodine correlated well with basic strength of alkali-exchanged zeolites. This technique has been extended to characterize alkali species supported in zeolite pores and on metal oxides. However, the iodine appeared to react irreversibly on these strongly basic samples, possibly forming an adsorbed triiodide ion.95The reactivity of iodine on strong bases limits the usefulness of this method to the characterization of weak-to-moderate solid bases. 3.2.3 Microcalorimetry of Ahorbed Probe Molecules - In addition to spectroscopic techniques, heat-flow microcalorimetry has been receiving greater attention as a useful method for measuring the heats of adsorption of various acidic probe molecules on basic materials. The heat released during adsorption is a measure of the strength of the bond formed between the adsorbed species and the catalyst surface, thereby giving a direct measurement of chemical reactivity of the surface. This technique can be used to measure the heats of adsorption as a function of the surface coverage. Thus, the distribution in the strengths of the surface sites can be obtained. Carbon d i o ~ i d e , * ~ ~ - dioxidelo*-lo3 and hexafluoroisopropano1104 loo, ~sulfur have been used as adsorbates in microcalorimetry studies. In addition, Huang et al. have carried out adsorption microcalorimetry of pyrrole to rank the basicity of alkali-exchanged zeolites. lo5 Carbon dioxide adsorption microcalorimetry has also been used to characterize the basicity of zeolites containing occluded cesium oxide formed by decomposition of occluded cesium acetate.'O Even though C02 adsorbs in a variety of configurations depending on the nature of the surface, higher values for the heats of adsorption are generally expected when carbon dioxide adsorbs on highly basic surface sites. Complementary vibrational spectra are needed to identify the mode of adsorption and correctly interpret results from rnicr~calorimetry.~~ Figure 2.6 compares the differential heat of C 0 2 adsorption for cesium-

2: Catalysis by Solid Bases

53

+I

+CSX 0

50

100

150

200

250

C02Uptake ( p o l g-1) Figure 2.6 Differential heats of adsorption as a function of CO, uptake for CsX zeolites (adaptedfrom ref: 100)

exchanged X zeolite, with and without occluded cesium oxide. For the sample with occluded alkali metal oxide, there were a few sites with high heats of C02 adsorption (> 100 kJ mol- '). The adsorption energies decreased rapidly to a plateau, indicating that a majority of the sites had uniform strength. Incorporation of occluded CsO, in CsX increased the total number of base sites as well as their strength. Even though the number of base sites increased with increasing amounts of occluded CsO,, the strengths of a majority of the sites in CsO,/CsX seemed to be unaffected by the CsO, loading. Most of the adsorption sites in each of the CsO,/CsX samples had energies close to 85 kJ mol-'. It is believed that decomposition of cesium acetate occluded in cesiumexchanged zeolites leads to the formation of cesium oxide ~ 1 ~ ~ t e These occluded moieties are apparently more basic than the framework oxygen atoms of the zeolite. Lasperas et al. studied a series of CsX and CsY zeolites with different loadings of occluded Cs by TPD of adsorbed C02.36J08They observed that the amount of C02 desorbed (up to 773 K) was proportional to the amount of excess cesium, up to 16 Cs atoms per unit cell of CsX. Since the stoichiometry was measured to be about one CO2 molecule desorbed for every two added Cs atoms, Lasperas et al. concluded that the composition of the occluded species was Cs20. Table 2.2 shows the number of C02 molecules adsorbed per unit cell as a function of the number of excess Cs atoms in CsX zeolites prepared by Bordawekar and Davis, loo The C 0 2 adsorption capacity increased linearly with the amount of occluded Cs, which is consistent with the observation of Lasperas et al. However, Bordawekar and Davis found that about one C02 molecule was adsorbed for every four occluded Cs atoms, which does not support the ideal formula of Cs20 for the occluded cesium oxide. loo 369108

r

~

.

Catalysis

54

Table 2.2 Carbon dioxide adsorption capacity of Cs0,-containing zeolitesIm ~~

~

Excess Cs atoms per unit cell of CsX 0.0 1.3 4.0 6.7 10.7 16.0 21.3

C 0 2 adsorbed (molecules per unit cell) 0.47 0.61 1.09 1.61 2.42 3.97 5.83

3.2.4 Temperature Programmed Desorption ( T P D ) of Probe Molecules - This method is used to measure the number and base strengths of sites found on solid base catalysts. Since strongly bound probe molecules have high binding energies, increased temperatures are necessary to desorb these adsorbates. Experiments are typically performed under identical experimental conditions (heating rates and sample size) so that a qualitative comparison can be made between samples, By using thermal-gravimetric analysis, the number of adsorbed probe molecules can be measured directly and, assuming a ratio of adsorbate molecules to active sites, the number of sites can be counted. During a TPD experiment, the amount of desorbed molecules is often monitored by mass spectrometry and the surface interactions are explored with infrared lo spectroscopy. Numerous texts describe in detail the TPD TPD of adsorbed carbon dioxide has been widely used to probe basic materials. For example, rubidium-modified supports have been investigated using stepwise TPD of C02.29 The addition of Rb species to supports like MgO, A1203, Ti02 and Si02, via the decomposition of supported rubidium acetate, increased the surface density of adsorbed COZ over that of the pure support. The high desorption temperatures required to liberate C 0 2 from RbOWMgO indicated the formation of very strong base sites. Carbon dioxide temperature programmed desorption has also been used to measure the base strengths of various alkali metal-containing (exchanged and occluded) zeolite^.^^^^ In addition, the basicity of alkaline earth oxides determined from CO2 TPD is MgO
3.2.5 X-Ray Absorption Spectroscopy - The lack of long range order associated with supported phases often prevents structural characterization by techniques like X-ray diffraction, which require samples with substantial crystallinity.

2: Catalysis by Solid Bases

55

X-Ray absorption spectroscopy (XAS) is a technique that is able to investigate the local structure near a particular element. When using a standard of known composition, the oxidation state, coordination number, nearest-neighbor interatomic distance(s) and Debye-Waller factors can be determined. Excitation of a core level electron results in a step change in the absorption coefficient associated with a particular element, and this phenomenon is known as the X-ray absorption edge. The X-ray absorption near edge structure (XANES) technique gives valuable chemical information in the vicinity of the absorbing atom. A large absorption peak that sometimes occurs immediately after the absorption edge in the spectrum is known as the ‘white line’, and is a measure of the density of unoccupied states near the absorbing atom. Thus, in addition to the oxidation state of an element, the relative number of accessible electronic states can be derived. X-Ray absorption spectroscopy has been used to study many catalytic materials. However, systematic investigations of supported alkali metals and their effect on activity in base-catalysed reactions have not been performed, despite the valuable local structural information this technique can give. Nevertheless, Clausen et al. have investigated the local structure around Rb promotors for ammonia synthesis catalysts using X-ray absorption spectroscopy and found that the calcined catalyst contained a disordered rubidium oxide phase, which is assumed to enhance the surface basicity of the catalyst. The R b - 0 phase changed to a lower coordinated environment and shorter interatomic distances during reduction of the catalysts up to 673 K. In addition, Tsuji et al. studied MgO occluded in NaY and MgY zeolites.6 Analysis of the extended X-ray absorption fine structure above the magnesium K-edge suggested that MgO crystallites formed in the zeolite supercages retain the rock salt crystal structure associated with bulk magnesia. However, the near edge spectrum associated with the supported sample was modified from that of bulk MgO, which was attributed to less charge separation in the former. Thus, Tsuji et al. propose that the basic oxygen atoms associated with the intrazeolitic magnesia clusters are less basic than those associated with bulk magnesia. X-Ray absorption spectroscopy has been used as a tool to study alkali metal-support interactions for rubidium deposited on pure oxides at the Rb Kedge.29As seen in Table 2.3, the Rb-O interatomic distances determined from EXAFS obtained at 77 K were found generally to decrease with increasing Sanderson electronegativity of the support.29 In fact, the reduced intensity of the Rb-0 peak in the radial structure function for rubidium-modified Si02 indicated the formation of a highly disordered rubidium silicate phase due to strong interaction between silica and the supported alkali species. The decrease in the interatomic distances over a range of supports of various electronegativity coincided with a decrease in the strengths of the base sites present on each support. This result indicates that strongly basic alkali-containing catalysts should utilize basic carriers of low electronegativity to minimize alkalisupport interactions that reduce the base strength of the supported species. Although rubidium oxide is classified as a superbase, it does not have as

56

Catalysis

Table 2.3 Interatomic distances of Rb-0 for various Rb-modijied supports (adaptedfrom ref 29) Catalyst

Rb/MgO Rb/TiOz Rb/A 1 2 0 3 Rb/SiOz

Interatomic distancelA

2.94 2.86 2.91 2.80

Sanderson int ermedideate electronegativity of the support

2.80 3.36 3.70 4.26

high a base strength as that of cesium oxide. * Cesium itself has been studied very little in the literature using X-ray absorption spectroscopy. Only recently has X-ray absorption spectroscopy been used to study cesium in catalytic materials for the determination of the Cs-0 distance in CsZSM-5 zeolite.'16 Doskocil and Davis recently published a study of the Cs LIII edge of various Cs-containing compounds and catalyst^.^^ The LIII edge corresponds to the excitation of a 2~312electron into unoccupied states of s- and d-type symmetry. The higher the intensity of this 'white line' transition, the lower the ground state electron occupation of the low lying vacant states associated with cesium. For Cs, the edge structure is likely to be dominated by the unoccupied 6s states. However, a density functional theory study of a Cs-O dimer revealed that the highest occupied molecular orbital is formed by the overlap of an 0 2p orbital with that of Cs 5d orbital. Thus, Cs atoms that are highly coordinated to oxygen may also have a smaller LIIIedge feature due to fewer available 5d states around Cs. Nevertheless, as long as the Cs-0 interatomic distance is identical, the white line intensity should correlate with the ionicity of the Cs-O bond, which decreases the s-type orbital population near Cs. Figure 2.7 compares the white line intensity at the LIIIedge for CsX and Cs2CO3. The higher intensity of the white line associated with Cs-exchanged X zeolite results from the ionicity of the Cs-zeolite interaction. Apparently, the bonding in cesium carbonate is more covalent than in the zeolite, which results in a greater density of filled s-type states for the bulk compound. Both impregnated and ion-exchanged cesium catalysts showed a more intense white line than any Cs compound investigated, indicating the ionic character of cesium atoms present in the zeolite catalysts compared to those in bulk phase compounds. 3.2.6 X-Ray Photoelectron Spectroscopy - Since the binding energy of electrons in a basic oxide is related to basic strength, XPS can be used to directly measure the core level binding energy of electrons associated with surface oxygen atoms present in the sample. Shifts to lower energy in the 0 1 , binding energy indicate stronger electron pair donation and stronger basicity. l 7 The 01, binding energies were obtained for numerous simple oxides by Dimitrov et al. * Okamoto et al. used XPS to investigate various alkali metal-exchanged zeolites and observed that the 01, binding energy decreased with decreasing electronegativity of the counter-ion, i.e. from Na to Cs, as well as showing a decrease with decreasing Si :A1 ratio. I l 9

2: Catalysis by Solid Bases

-10

57

-5

0

5

10

15

20

Energy (eV) Figure 2.7 Comparison of white line intensity at the CS LllI edge of CsX and C S , C O ~ . ~ ~

In addition to investigating the basic surface, XPS has also been used to measure the binding energies of electrons in probe molecules that are chemisorbed on a base site. Chloroform adsorbed in ion-exchanged zeolites120and on KN03 or K2CO3 loaded onto A1203 have been investigated by measuring the C12p binding energies.121Xie et al. showed that chloroform was adsorbed only on the strongly basic oxygen anions directly adjacent to the alkali metal cations. 120 Furthermore, the negative partial charge on chlorine increased with decreasing C12p binding energy. The XPS of adsorbed chloroform is a useful technique for evaluating zeolite basicity.120 Borade et al. found that shifts in the N1, binding energy of pyrrole, an acidic probe molecule, also provided a large enough scale to effectively characterize base strength. 122

4

Probe Reactions over Solid Base Catalysts

Base-catalysed reactions occur via carbanion intermediates as opposed to acid-catalysed reactions that involve carbocation intermediates. Hence, the product selectivities in base-catalysed reactions are different from acidcatalysed pathways. Moreover, coking of the catalyst is seldom observed in base-catalysed reactions, since coking is due to reactions which proceed via a carbenium ion. Numerous catalytic reactions have been investigated using solid bases. Some of these have been mentioned earlier in this chapter when introducing various

58

Catalysis

types of basic catalysts and their past uses. Reviews by Hattori,'17 Barthomeuf? and Ono and Baba2 present the wide variety of reactions that have been studied. Below, some of the heterogeneous base-catalysed reactions that are commonly used to study solid bases are summarized. This section will not discuss every reaction catalysed by solid bases, but will instead focus on reactions of interest as probes of solid basicity.

4.1 Double-bond Isomerization - The base-catalysed isomerization of olefins was first studied by Pines et a13 Most isomerization studies have been performed on simple olefins. However, a few studies describe the double-bond migration of unsaturated compounds that contain atoms like N and 0, such as ally la mine^'^^ and propen-2-yl ethers, respectively.124The isomerization reaction is initiated when an allylic proton is abstracted by a base site, forming n-allylic anions in which the cis form is typically more stable than the trans form. Since the initial studies by Pines, the double-bond isomerization of but1-ene has emerged as a standard probe reaction in the catalysis community for characterizing the basicity of solid catalysts. Base-catalysed isomerization of but- 1-ene selectively results in double-bond migration without any skeletal isomerization. The cisltrans ratio of the product but-2-enes is typically greater than four over solid bases, indicating an anionic reaction mechanism. The isomerization of but- 1-ene has been studied over many solid bases, including alkaline earth oxides, calcined hydrotalcites,3 127and alkali-exchanged zeolites as well as zeolites containing occluded alkali oxide^.^,^^ The activity of MgO and SrO for but- 1 -ene isomerization was found to be a function of the pretreatment t e m p e r a t ~ r e . ' ~ Mohri ~ . ' ~ ~ et al. claimed that but- 1-ene isomerization over SrO actually occurs on acid-base pair sites.'26 They attribute the formation of trans-but-2-ene largely to the acid sites on the surface, presumed to be Sr2+,and the formation of cis-but-2-ene to the basic 02-ions. In the case of alkali-modified zeolites, occlusion of alkali oxide clusters in the zeolite cavities resulted in an increase in both the activity for but- 1-ene isomerization and the cisltrans ratio of the product but-2-enes, compared to the ion-exchanged material^.^.^^ There is a direct correlation between activity of zeolites for but-1-ene isomerization and the loading of occluded alkali metal oxide in the pores.loO Other double-bond migration reactions have been investigated as possible probe reactions for evaluating the nature of solid base surfaces. One of these additional reactions is the isomerization of 3-carene to 2-carene. This reaction effectively ranked basicity for simple oxides like MgO and CaO, but proved to be less suitable for highly basic materials like sodium-loaded NaX zeolite due to the formation of several by-products. 128*129High rates of by-product formation during the double-bond migration of 2- and 3-carene oxides are why these reactions have not received much attention for characterizing base catalysts. 130 Recently, Handa et al. have carried out the isomerization of 2,3-dimethylbut-1-ene in the liquid phase for a variety of basic catalysts, including alkaline earth oxides, various mixed oxides and zeolites. The authors

'

7

259126

2: Catalysis by Solid Bases

59

suggest that this alkene isomerization is an effective new probe reaction for studying strong basic sites.

4.2 Alcohol Dehydrogenation - Alcohol decomposition occurs over both acids and bases. On an acidic catalyst, alcohols predominantly dehydrate to ethers or olefins, whereas, on basic catalysts they predominantly dehydrogenate to aldehydes or ketones. However, alcohol dehydration can also occur over basic catalysts by a mechanism different from acid-catalysed dehydration.' l 7 Alcohol decomposition has been studied over many metal oxides, including Ti02,132 Ce02133 and Zn0.134 Kim and Barteau observed that aliphatic alcohols were adsorbed on Ti02 in two different states, namely, molecular alcohol and alkoxide. 132 They found that the molecularly adsorbed alcohols desorbed intact upon heating, whereas the alkoxide species followed multiple pathways: recombinative desorption to form dialkyl ethers, a-hydrogen elimination to form carbonyl compounds, P-hydrogen elimination and C-0 bond cleavage to form olefins. One probe reaction that has been used extensively to investigate both acid and base sites is the decomposition of p r o p a n - 2 - 0 1 . ~ ~35~ ~137 * *Dehydra~~'~~~~ tion of propan-2-01 to produce propene occurs primarily over acid sites, and dehydrogenation to produce acetone occurs primarily over basic sites. Neither of these reactions require strong sites and, therefore, can be useful in investigating subtle changes in the acid-base character of metal oxide catalysts. The product selectivities in the decomposition of propan-2-01 have been used to characterize alkali- and alkaline earth-modified metal oxide^.^^.^^ Incorporation of alkali and alkaline earth elements on various metal oxides increased the selectivity to acetone in the decomposition r e a c t i ~ n . Hathaway ~ ~ . ~ ~ and Davis have studied the decomposition of propan-2-01 over alkali-modified zeol i t e ~ . ~ *They ' * ~ found that incorporation of occluded cesium oxide clusters in cesium-exchanged Y zeolite resulted in an order of magnitude increase in acetone production. On a surface area basis, the acetone activity of alkali oxide/zeolite catalyst was comparable to that of MgO. lo6 Results for the reaction of propan-2-01 after 2 h at 563 K on CsX zeolites containing occluded Cs oxide can be found in Figure 2.8. For CsX without any occluded cesium species located in the cages, the selectivity to acetone was 9%. The occlusion of only 1.3 CsOx per unit cell resulted in a significant increase in the selectivity due primarily to a four-fold increase in the rate of acetone production. Higher loadings of excess cesium further increased the acetone selectivity. An order of magnitude increase in the loading of cesium oxide resulted in an analogous order of magnitude increase in the rate of acetone production over the range of Cs loadings presented. The occlusion of CsOx in the zeolite pores did not significantly affect the specific rate of propene production over the range 0-10.7 excess Cs per unit cell, indicating that the supported species did not reduce support acidity. This result contrasts with what has been observed for rubidium oxide supported on classic metal oxides.29It appears that the occluded species in zeolites catalyse the reaction

Catalysis

0

2

4

6

a

10

12

Excess Cs Atoms per Unit Cell Figure 2.8 Rate of acetone and propene formation from propan-2-01 reaction over CsX zeolites containing occluded cesium oxide52

without significantly affecting the acid-base properties of the zeolite framework. Recently, Handa et al. suggested the use of two complementary probe reactions to characterize solid base catalysts.131In addition to the liquid phase isomerization of 2,3-dimethylbut-1-ene to 2,3-dimethylbut-2-ene(mentioned at the end of Section 4.1), the authors suggest using the vapor phase decomposition of 2-methyl-3-butyn-2-01 to characterize the acid-base character of moderately basic solids. Acidic catalysts produce 3-methyl-3-buten-1-yne and/ or 3-methyl-2-buten-1-al, whereas basic catalysts yield acetylene plus acetone. Amphoteric catalysts produce 3-hydroxy-3-methylbutan-2-oneand/or 3methylbut-3-en-2-one. Analysis of the product composition can effectively characterize the acid-base character of the catalyst surface.

4.3 Hydrogenation Reactions - Olefin hydrogenation is also catalysed by solid bases. Hydrogenation reactions on solid bases normally involve heterolytic cleavage of H2 into H+ and H - . The characteristics of solid bases in hydrogenation reactions are very different from those associated with conventional transition metal catalyst^."^ Base catalysts are more active for the hydrogenation of conjugated dienes than mono-olefins. Also, over solid bases, 1,4-addition of H atoms is favored in conjugated dienes, as opposed to the 1,2addition commonly seen over conventional hydrogenation catalysts. Hattori et al. have shown that the molecular identity of H2 was maintained during the hydrogenation of 1,3-butadiene over MgO, i.e. both the H atoms in a H2 molecule were incorporated in the same hydrogenated molecule.13* The hydrogenation of olefins also occurs over alkali-modified zeolite^.^ Tsuji et al. have studied the hydrogenation of 1,3-butadiene over various basic zeolites.6 They found that alkali-exchanged and Mg-exchanged zeolites, as well as zeolites containing occluded alkali oxides, were inactive for the hydrogena-

2: Catalysis by Solid Bases

61

tion reaction at 423 K. However, Mg-exchanged Y zeolite containing occluded MgO clusters was active in the reaction.6 The authors claim that heterolytic dissociation of H2 requires a highly acidic metal cation, which is offered by Mg2+ in occluded MgO, but which is absent in the alkali oxides occluded in zeolites. 4.4 Condensation Reactions - Condensation of ketones and/or aldehydes are useful reactions in organic synthesis because they form new C-C bonds between two molecules. Condensations are usually catalysed by bases, but have been shown to proceed over acids as well. Both self- and cross-condensations can occur if more than one aldehyde is present. One of the challenges in designing base catalysts for cross-condensation reactions is to influence the surface base properties to decrease the ratio between self- and cross-condensation products.

4.4.I Aldol Condensation - Aldol condensation reactions are used for organic synthesis of molecules that contain a double bond conjugated with a carbonyl group. 39 Dehydration of the condensation product gives an a,P-unsaturated molecule. Various aldehydes and ketones have been used for aldol condensation, including formaldehyde, acetaldehyde140-*42and acetone.4-56”39*143-150 The condensation of acetone has been the most widely studied of all aldol condensations for solid bases and will be the focus of this section. Figure 2.9 shows the schematic for the aldol condensation of acetone. The aldol condensation of acetone is initiated by the extraction of the acidic a-proton of acetone to form the enolate anion. Another acetone molecule is then attacked by this anion to form a tetrahedral alkoxide ion which is protonated to form diacetone alcohol. This branch of the aldol condensation is catalysed by many metal oxides, including alkaline earth oxides, La2O3, Zr02, Si02-Al203 and Nb205.143The order of catalytic activities on a unit surface area basis is reported to be BaO > SrO > CaO > MgO > La203 > Z r 0 2>> SO2A1203> Nb2O5. In the case of MgO, addition of water resulted in an increase in the activity and selectivity to diacetone alcohol. 143 Hence, the researchers propose that the basic OH- ions either retained on the surface or formed via dehydration of diacetone alcohol are the active sites for the reaction. Di Cosimo et al. carried out the aldol condensation of acetone over alkaliand alkaline earth-modified MgO catalysts. 139The authors found that mesityl oxide, isomesityl oxide and isophorone (Figure 2.9) were the major products formed in the reaction and that the activity of MgO was enhanced by addition of alkali or alkaline earth. The Li/MgO catalyst produced a larger amount of isophorone than any of the other promoted MgO samples. Since the condensation of acetone to give isophorone requires very strong base sites, the authors conclude that Li/MgO was the most basic of all the promoted MgO catalysts. The authors believe that addition of lithium to MgO results in the replacement of the Mg2+ ions in the MgO lattice with Li+. This causes a straining of the Mg-0 bonds and formation of (Li+O-) species leading to the generation of very strong base sites.

’

Catalysis

62

'

CH3

R

&-

CH3

o=cI

H,C-C-CH,

I

O=?

I

Diacetone Alcohol

CH*

CH3

1

H,C-C-CH,

Acetone

I

OH

f 9

HJ-C-CH,

o=c I

CH

II

H,C

/C\

CH,

o=c

I

CH,

I 4%

H,C

CH,

0

Isophorone

Figure 2.9 Schematic of the aldol condensation of acetone over solid base catalysts

Aldol condensation of acetone has also been reported to increase the formation of methyl isobutyl ketone via additional hydrogenation of the double bond of mesityl oxide as well as to decrease the selectivity to isobutene when more basic alkali-modified zeolites were 50 In addition, reactions have been performed using oxidized cesiumhanoporous carbon catalysts in an attempt to minimize the formation of higher molecular weight condensation products obtained during liquid phase reactions and, therefore, maximize formation of the more industrially important i ~ o p h o r o n eOver . ~ ~ the temperature range of 473 to 573 K, the selectivity to isophorone was between 54 and 6l%, which was maintained even when the catalyst deactivated. However, at temperatures greater than 523 K, the activity of the catalyst decreased because of the accumulation of higher molecular weight products remaining in the carbon pores.56 4.4.2 Knoevenagel Condensation - Knoevenagel condensation involves the formation of a C-C bond via the reaction of an aldehyde or a ketone with an active methylene compound. Usually, this reaction employs compounds containing a methylene group activated by two electron-withdrawing moieties.

2: Catalysis by Solid Bases

63

CHO 0

+

It

CH3

Benz aldehyde

-C-

0

It CH2- C

- OC&

Ethyl acetoacetate

0

CH=YI

I1

C- CH3

@ r0c2Hs + H20

Figure 2.10 Knoevenagel condensation of benzaldehyde with ethyl acetoacetate

Commonly used active methylene compounds include malonates, acetoacetates, acetonitriles, acetylacetone and malonodinitrile. Figure 2.10 shows the Knoevenagel condensation of benzaldehyde and ethyl acetoacetate. Martin-Aranda et al. have studied the condensation of benzaldehyde with malonic esters over alkali-modified carbon catalysts and found the order of activity for the modifier to be Li < Na < K c Cs.I5' The condensation reaction has also been studied over alkali-exchanged faujasites,* faujasites containing occluded alkali oxides34 and hydr~talcites.'~~ In the case of faujasites, incorporation of occluded alkali oxides resulted in an increase in the activity for the reaction compared to the parent alkali-exchanged zeolite.34 Also, for ion-exchanged faujasites, substitution of Si by Ge during the zeolite synthesis resulted in an increase in the activity for Knoevenagel condensation.*53This increase in activity has been attributed to an increase in basicity resulting from Ge sub~titution.~ As mentioned earlier, novel basic catalysts with high specific surface areas have been synthesized by activation of an aluminophosphate oxide precursor under ammonia to form aluminum phosphate oxynitride (A1PON).58 These AlPON catalysts were found to be more active than MgO for Knoevenagel condensation and their activity was found to be a function of their nitrogen content 4.5 Alkylation Reactions - Alkylation reactions yield very different product selectivities over acidic and basic catalysts. Usually, alkylation of aromatics over basic catalysts results in side-chain alkylation products, whereas over acidic catalysts ring-alkylates are formed. For instance, over acidic zeolites, toluene reacts with methanol to form primarily xylenes. However, over basic zeolites, styrene and ethylbenzene are the dominant products. The alkylation of toluene with methanol over solid base catalysts is an attractive single step route for styrene synthesis, The traditional process for styrene manufacture involves Friedel-Crafts alkylation of benzene with ethylene to make ethylben-

64

Catalysis

zene, followed by dehydrogenation of ethylbenzene. The dehydrogenation step is energy intensive and hence economically demanding. In the alkylation of toluene with methanol over solid bases, formaldehyde formed in situ from methanol dehydrogenation is the alkylating agent.4 Alkali-exchanged zeolites X and Y are active for the side-chain alkylation of toluene with methan01.~An undesirable side reaction is the decomposition of formaldehyde to carbon monoxide. Occlusion of alkali metal oxides in zeolites results in increased basicity and a greater activity for toluene alkylation at the expense of formaldehyde decomp~sition.~ Recently, Wieland et al. found that cesium-exchanged zeolites L and p were also active for toluene alkylation with methan01.~Although the L and p zeolites required higher temperatures to achieve aromatic yields similar to alkali-exchanged faujasites, very little carbon monoxide was produced over these catalyst^.^ The side-chain alkylation of aromatics with olefins is also catalysed by solid bases. However, this reaction requires greater base strength than alkylation with methanol. Pines and co-workers did pioneering work on the side-chain alkylation of alkylbenzenes with olefins. 154-156 They carried out the reaction using sodium metal in the presence of organic promoters such as anthracene, o-chlorotoluene, o-toluic acid, pyridine and organic peroxides. The sodium metal reacted in situ with the organic promoter to form an organosodium compound which then catalysed the alkylation reaction. Martens et al. have studied the side-chain alkylation of alkylaromatics with ethylene over zeolites containing occluded alkali metals. 57 They observed that the alkylation activity of Na metal in NaX zeolite was higher than that of Na metal in NaY zeolite. Also, the alkylation activity for isopropylbenzene was lower than that of ethylbenzene, possibly due to steric effects.157 Basic zeolites also catalyse the side-chain alkylation of heteroaromatics. For instance, the alkylation of aniline with methanol occurs over alkali-exchanged zeolites and the selectivity to the N-alkylate increases with increasing basicity of the ~ a t a l y s tMagnesium-aluminum .~ mixed oxides prepared by calcination of hydrotalcites have been used for the alkylation of phenol with methanol to make m e t h o x y b e n ~ e n e . ~ ~ ~ 4.6 Side-chain Alkenylation of o-Xylene with 1,3-Butadiene - Polyethylene naphthalate (PEN) is a polyester which has the potential to be an important new plastic due to its increased rigidity, strength, heat stability and barrier properties compared to that of polyethylene terephthalate (PET) which is commonly used today. 159 The increased stability over other polyesters which contain single-ring structures comes from the double-ring structure of the naphthalene found in the polymer backbone. The primary use of PEN has been for specialty film production in the photographic industry. However, widespread use of PEN will not occur until the cost of producing naphthalene dicarboxylate (NDC), the main raw material used in PEN production, is lowered. Once this occurs, PEN will be able to better compete in such industrial applications as packaging materials, heavy-duty products and shatter-resistant bottles. 59

2: Catalysis by Solid Bases

65

One possible precursor for NDC is produced from the side-chain alkenylation of o-xylene with 1,3-butadiene to form the monoalkenyl aromatic hydrocarbon compound 5-o-tolylpent-2-ene (OTP). This requires a stronger base site than that associated with supported alkali metal oxides. Many catalysts have been investigated which use an alkali metal supported on a basic carrier, including alkaline earth oxides,'60*161 basified and zeolites. 165 The mechanism for this reaction is similar to that seen for side-chain alkylations and alkenylations of alkylbenzenes using both monoolefins' and di-olefins. 168-169An o-xylene molecule undergoes an initiation step that results in the substitution of an alkali metal for a benzylic hydrogen, followed by butenylation to the benzylic carbanion intermediate. Undesired alkenylations can take place at a side-chain over the base site or at a ring position if acidic sites are present on the catalyst. These additional alkenylations cause the selectivity to the desired OTP product to decrease. By supporting the alkali metal clusters in carriers containing restricted pore sizes, like zeolites, the selectivity to the desired OTP product has the potential to be increased. Several factors must be considered to synthesize an effective catalyst for the production of OTP from o-xylene and 1,3-butadiene. First, the trans isomer of 5-o-tolylpent-2-ene is more desirable than the cis isomer. 165 Second, multiple alkenylations can take place at the side-chain, thus decreasing the selectivity to OTP product. Also, alkenylation of the aromatic ring can occur on residual acid sites of the catalyst. Finally, 1,3-butadiene can oligomerize over strong base sites and deactivate the catalyst. Various zeolites modified by the incorporation of alkali metals in the pores were investigated as potential catalysts that could enhance the selectivity a n d or the relative rate to the trans form of the OTP product.52Extra-lattice metal clusters were evident on each of the alkali metal-loaded zeolites that were effective for the side-chain alkenylation of o-xylene with 1,3-butadiene to form OTP, suggesting that these clusters might be the active species for the reaction. For each reaction in which OTP was produced, the solution color indicated that alkali metal formed from the decomposition of the supported azide may have leached into solution to form an organoalkali complex. Thus, catalytic activity was likely due to metal species in the solution as well as those associated with the zeolite (possibly through reaction with the zeolite). Nevertheless, these studies showed interesting ways to generate active alkali species for carrying out alkenylation. 5411557166v167

4.7 Miscellaneous Reactions - In addition to the reactions described above, solid bases also catalyse other reactions, including amination and dehydrocyclodimerization of conjugated dienes (Figure 2.1 1). I7 Suzuka and Hattori have studied the dehydrocyclodimerizationof 1,3-butadiene to make ethylbenzene over alkaline earth 0 ~ i d e s . Ueda l ~ ~ et al. have reported the use of alkaline earth oxides and ZnO for higher alcohol synthesis by condensation of methanol with C2-C5 primary alcohols (Figure 2. 12).17' Alkali-exchanged zeolites can catalyse the ring transformations of y-butyrolactone and tetrahy-

Catalysis

66 Amination

CH3 'NH

+

/

CH3 Dimethylamine

CHf=CH-CH=CH2

1

Butadiene

Solid Base

Dehydrocyclodimerization

F2H5

2(CH2=CH-CH=CH2)

Solid Base ____)

Butadicne

@

+ H2

Ethylbenzene

Figure 2.1 1 Amination and dehydrocyclodimerization of 1,3-butadiene catalysed by solid bases

R-CH2-CH2-OH

+ CH30H

[ R=H,Alkyl]

Primary alcohol

Basic Metal Oxide ____)

R-CH-CHTOH I CH3

+ H20

Methanol

Figure 2.12 Solid base-catalysed synthesis of higher alcohols from methanol and primary alcohols

drofuran, to y-thiobutyrolactone and tetrahydrothiophene, respectively (Figure 2.1 3). 72 The activity of alkali-containing zeolites for the base-catalysed formation of ethylene carbonate from ethylene oxide and carbon dioxide has recently been tested.95 Among the ion-exchanged zeolites (Na, K, Cs), the cesium form of zeolite X exhibited the highest activity for ethylene carbonate formation. The catalytic activity of a zeolite containing occluded cesium was even higher than that of a cesium-exchanged zeolite. The presence of water adsorbed in zeolite pores promoted the rate of ethylene carbonate formation for both cesium-exchanged and cesium-impregnated zeolite X. These results suggest that this reaction may be a promising probe for solid base sites, but additional work is needed to understand the role of adsorbed water on the reaction kinetics.

2: Catalysis by Solid Bases

67

+ H2S-

0 + H20

0 ybutyrolactone

y-thi obutyrolactone

tetrahydrofuran

tetrahydrothiophene

Figure 2.13 Ring transformation reactions catalysed by alkali-exchanged zeolites

5

Conclusions

Given the wide variety of reactions that are base-catalysed, the usefulness of solid bases as catalysts for chemicals production will continue to increase. As industrial interest in solid base materials grows, additional investigations into how basic sites are created and how the strength of these sites can be tailored to perform a desired transformation with high selectivity are needed.

Acknowledgments This work was supported by the Department of Energy (Basic Energy Sciences, Grant DEFGOS-95ER14549).

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118 V. Dimitrov, T. Komatsu and R. Sato, J. Ceram. SOC. Jpn., 1999,107,21. 119 Y.Okamoto, M. Ogawa, A. Maezawa and T. Imanaka, J. Catal., 1988,112,427. 120 J. Xie, M. Huang and S. Kaliaguine, Appl. SurJ: Sci., 1997, 115, 157. 121 J. Zhu, Y. Wang, Y. Chun and Y . Hu, Huaxue Wuli Xuebao, 1998,11, 178. 122 R.B. Borade, M. Huang, A. Adnot, A. Sayari and S. Kaliaguine, Stud. SurJ Sci. Catal., 1993,75, 1625. 123 A. Hattori, H. Hattori and K. Tanabe, J. Catal., 1980,65, 245. 124 H. Matsuhashi and H. Hattori, J. Catal., 1984,85,457. 125 M.J. Baird and J.H. Lunsford, J. Catal., 1972,26,440. 126 M. Mohri, K. Tanabe and H. Hattori, J. Catal., 1974,32, 144. 127 A. Beres, I. Hannus and I. Kiricsi, React. Kine?. Catal. Lett., 1995,56, 55. 128 K. Shimazu, H. Hattori and K. Tanabe, J. Catal., 1977,48,302. 129 U. Meyer and W.F. Hoelderich, J. Mol. Catal. A:Chem., 1999, 142,213. 130 K. Arata, J.O.J. Bledsoe and K. Tanabe, J. Org. Chem., 1978,43, 1660. 131 H. Handa, Y. Fu, T. Baba and Y . Ono, Catal. Lett., 1999,59, 195. 132 K.S. Kim and M.A. Barteau, Langmuir, 1988,4, 533. 133 M.I. Zaki and N. Sheppard, J. Catal., 1983,80, 114. 134 0 . Koga, T. Onishi and K. Tamaru, J. Chem. Soc., Faraday Trans. I , 1980, 76, 19. Jpn., 1977,50,2579. 135 M. Ai, Bull. Chem. SOC. 136 A. Gervasini and A. Auroux, J. Catal., 1991, 131, 190. 137 J.H. Zhu, C. Yuan, Y. Qin and Q.-H. Xu, Microporous Mesoporous Mater., 1998, 24, 19. 138 H. Hattori, Y. Tanaka and K. Tanabe, J. Am. Chem. SOC.,1976,98,4652. 139 J.I. Di Cosimo, V.K. Diez and C.R. Apesteguia, Appl. Catal. A:Gen., 1996, 137, 149. 140 E. Dumitriu, V. Hulea, C. Chelaru, C. Catrinescu, D. Tichit and R. Durand, Appl. Catal. A:Gen., 1999, 178, 145. 141 W.J. Ji, Y. Chen and H.H. Kung, Appl. Catal. A:Gen., 1997, 161,93. 142 Y.C. ChangandA.N. KO, Appl. Catal. A:Gen., 2000,190,149. 143 G . Zhang, H. Hattori and K. Tanabe, Appl. Catal., 1988,36, 189. 144 A. Philippou and M.W. Anderson, J. Catal., 2000, 189, 395. 145 F. Prinetto, D. Tichit, R. Teissier and B. Coq, Catal. Today, 2000,55, 103. I46 R. Unnikrishnan and S. Narayanan, J. Mol. Catal. A:Chem., 1999,144, 173. 147 M.N. Bennani, D. Tichit, F. Figueras and S. Abouarnadasse, J. Chim. Phys. PCB, 1999,%, 498. 148 J.I. DiCosimo and C.R. Apesteguia, J. Mol. Catal. A:Chem., 1998, 130, 177. 149 K.K. Rao, M . Gravelle, J.S. Valente and F. Figueras, J. Catal., 1998, 173, 115. 150 C.O. Veloso, J.L.F. Monteiro and E.F. Sousa-Aguiar, Stud. Surf. Sci. Catal., 1994,84, 1913. 151 R.M. Martin-Aranda, M.L. Rojas Cervantes, A.J. Lopez-Peinado and J.d.D. Lopez-Gonzalez, J. Mol. Catal., 1993,85, 253. 152 A. Corma, V. Fornes, R.M. Martin-Aranda and F. Rey, J. Catal., 1992, 134, 58. 153 A. Corma, R.M. Martin-Aranda and F. Sanchez, J. Catal., 1990, 126, 192. 154 H. Pines, J.A. Vesely and V.N. Ipatieff, J. Am. Chem. SOC.,1955,77, 554. 155 H. Pines and V. Mark, J. Am. Chem. SOC., 1956,78,4316. 156 H . Pines and L. Schaap, J. Am. Chem. SOC.,1958,80,3076. 157 L.R. Martens, W.J. Vermeiren, D.R. Huybrechts, P.J. Grobet and P.A. Jacobs, in Proceedings of the 9th International Congress on Catalysis, Chem. Inst. of Canada, Ottawa, eds. M.J. Phillips and M. Ternan, 1988,420 pp.

72 158 159 160 161 162 163 164

165 166 167 I68 169 170 171 172

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S. Velu and C.S. Swamy, Appl. Catal. A:Gen., 1994,119,241. P.M. Morse, C&EN, 1997,75,8. G.G. Eberhardt and H.J. Peterson, J. Org. Chem., 1964,30,82. N.Fushimi, K. Inamasa and M. Takagawa, US Pat. 5,344,806, 1994. M. Takagawa, K. Kato, N.Fushimi and K. Kedo, US Pat. 5,436,381, 1995. M. Takagawa, K. Inamasa, N. Fushimi, A. Hashimoto and T. Sasaki, US Pat. 5,444,172, 1995. T. Matsumoto, Y. Kumagai and F. Kumata, US Pat. 5,625,102, 1997. C . Dimitrov, Z.I. Popova and F.V. Thyong, Competes Rendus de I’Academie Bulgare des Sciences, 1980,33,353. L. Schaap and H. Pines, J. Am. Chem. Soc., 1957,79,4967. H. Pines and L.A. Schaap, Adv. Catal., 1960,12, 117. H. Pines and N.C. Sih, J. Org. Chem., 1965,30,280. H. Pines and J. Oszczapowicz, J. Org. Chem., 1967,32, 3183. H. Suzuka and H. Hattori, J. Mol. Catal., 1990,63,371, W . Ueda, T. Ohshida, T. Kuwabara and Y. Morikawa, Stud SurJ Sci. Catal., 1993,75,2821. Y. Ono, Stud. Surf. Sci. Catal., 1980,5, 19.

3 Solid Sorbents for Catalytic NO, Removal BY MASATO MACHIDA

1

Introduction

Interest in the application of heterogeneous catalysts for NO, control has increased markedly in the past two decades. Increasingly stringent regulations on NO, emissions from mobile and stationary sources require the development of more efficient NO, control techniques. One difficulty of catalytic processes is that the NO, concentration is far less than 1 vol % in almost all applications, leading to serious inhibition by 0 2 , CO2, H20 and/or SO,. From the viewpoint of efficiency and running costs, catalytic processes are not always the most appropriate for the removal of dilute levels of NO,. Alternatively, NO, removal has been attempted by use of solid sorbents, which have several advantages, including ease of material handling, simplicity of system design, and low maintenance requirements. Cyclic sorptioddesorption processes can be useful for the continuous separation of NO, diluted in air. Sorptive NO, removal utilizes adsorption, absorption and/or solid-gas reactions.' Adsorption on solid surfaces can be divided into physical and chemical adsorption. Physisorption equilibrium is rapid and reversible. This type of adsorption occurs as result of nonspecific intermolecular forces, such as the condensation of vapor to liquid, and is usually effective for removal of vapors (gases below their critical temperature). This means that physical adsorption is less selective to specific gas species, and the amount of adsorption greatly exceeds monolayer capacity. On the other hand, chemisorption usually occurs as a result of the specific interaction of the adsorbate molecule with the adsorption site, leading to marked selectivity to the specific gas species. The amount of chemisorption is limited to less than the monolayer capacity. Absorption (or sorption) means bulk-type reactions between the solid sorbent and NO,. Conventionally, this has been extensively studied by use of wet processes, e.g., NO oxidation to NO2 followed by absorption into alkali solutions. However, the use of oxidizing agents, and the handling and disposal of wet spent materials, contribute to the cost of wet sorption methods. By contrast, solid sorbents can be easily regenerated by means of thermal- or pressure-swing cycles, or by treating the solid with a reducing gas. This review provides updated information on solid sorbents for NO, removal, which includes carbons, zeolites, alkaline solids and metal oxides. Catalysis, Volume I5 0The Royal Society of Chemistry, 2000

73

Catalysis

14

There are a wide variety of possible NO, sorptioddesorption processes which can be integrated with catalytic reactions in new concepts.

2

Materials for NO, Sorption

Carbonaceous materials and zeolites are widely used in commercial gas separation processes such as gas bulk separations and gas p~rification.~-~ Bulk separation refers to an adsorbate concentration greater than about 10wt% in the feed, while purification refers to an adsorbate concentration of no more than a very few wt% in the feed. De-NO, is therefore a purification process, and high selectivity for NO, is required. In this section, recent research on the wide variety of solid sorbents for nitrogen oxides is reviewed. In most, but not all cases, these materials are not yet available for commercial use. However, it is quite useful to understand various solid-gas processes and materials available for the sorptive removal of NO,. 2.1 Carbonaceous Solids - It is well known that the adsorption of NO onto porous carbonaceous solids is temperature dependent. At temperatures well below ambient, mainly physisorption occurs, whereas chemisorption occurs to a significant extent at above ambient temperature*v9 and is generally accompanied by formation of surface oxides and release of N2. The largest possible sorption capacity can be attained when adsorbate molecules fill the micropores of the solid.l o This type of enhanced physical adsorption is a dominant process for vapors, i.e., microporous solids have a high adsorption rate and capacity for vapors by micropore filling. Generally, micropore filling is not effective for supercritical gases such as NO whose critical temperature (180 K) is less than the adsorption temperature. However, Kaneko and co-workers' -19 have revealed that the external surface modification of activated carbon fibers (ACFs) with chemically active substances introduces some specificity in the micropore filling of NO, i. e. they have proposed chemisorption-assisted micropore filling. Figure 3.1 shows NO adsorption isotherms over FeOOHand Fe203-dispersed ACFs at 3O"C.l7 Dispersion of FeOOH provides the micropore filling of NO, and the decomposition into Fez03 markedly enhances the micropore filling of NO into ACF up to 320 mg g-' at 30°C. This corresponds to 85% filling of the micropore volume. All isotherms consistent with the Langmuir equation exhibit remarkable hysteresis;11*17 the adsorbed NO cannot be removed by evacuation at ambient temperature. The incorporated NO can be reversibly recovered as NO by heating above 200°C. The initial adsorption rate of NO over Fe203-dispersed ACF was approximately 70 times higher than that over the unmodified ACF sample. A more interesting feature is selectivity, i.e. micropore filling of NO can remove 93% of NO from 300 ppm NO in the presence of 02, S02, C02 and H2O.I4 The mechanism of this type of NO adsorption is probably related to the chemisorption and physisorption characteristics. Gaseous NO molecules first chemisorb onto Fe203dispersed around the entrance of the slit-shaped micropores, and then

.

3: Solid Sorbents for Catalytic NO, Removal

75

NO Pressure / mmHg Figure 3.1 Adsorption isotherms of NO over various activated carbon samples at 30°C.17 (Top curve) FezOj-dispersed A C E (second from top curve) FeOOH-dispersed ACE (third from top curve) ACE (bottom curve) activated carbon.

migrate to fill the micropore of ACF as dimerized (NO), even above ambient temperature.l29'* The formation of the dimer as an intermediate in chemisorption onto activated carbon is also supported by kinetic measurement by Teng and Suuberg.20 Stencel and co-workers21*22 reported that, in the presence of gaseous 02, activated carbons display a high capacity to adsorb NO, which is useful for combustion flue gas treatment. Oxygen promotes the conversion of NO into NO2 and the chemisorption as nitrites (C-NOZ) within the micropore. The rate of NO oxidation is increased by the catalysis of activated carbons compared to homogeneous oxidation over silica materials.23 The carbon is stable in gas containing C02, 0 2 and H20 without being consumed by oxidation at temperatures below 200 "C. By contrast, the presence of SO2 inhibits the oxidation of NO and the subsequent storage of nitrites. Inhibition due to SO2 occurs in many NO, sorption systems, including almost all combustion applications, because SO2 and NO/N02 compete for the adsorptiodactivation sites on the solid surface. For activated carbons, the removal of adsorbed NO2 and SO2 and regeneration of the adsorption sites can be accomplished by temperature induced desorption. 2.2 Zeolites - Iwamoto and c o - w o r k e r ~ ~have ~ - ~studied ~ the adsorption properties of NO on various metal ion-exchanged zeolites in a fixed-bed flow apparatus. The amount of reversible and irreversible NO adsorption, qrevand qirr, respectively, which were measured at 273 K on various cation-exchanged ZSM-5, are summarized in Table 3.1 (ZSM-5 is denoted as MFI in Table 3.1). Based on the amount of NO adsorption, the zeolite adsorbents can be classified into two groups, i.e., ion-exchanged samples with alkali, alkaline earth and lanthanides, and those with transition metals.25 For zeolites with transition metals, qirr is larger than q r e v with exception of Zn- and Ag-exchanged samples. In contrast, qres was greater than qirr for the zeolites containing alkali,

76

Catalysis

Table 3.1 N O adsorptionproperties of various cation-exchangedMFI zeolites25 Cation

Degree of exchangeIYo

Content of cationlwt%

Amount of NO a&orbedalcm3 g-' Reversible Irreversible

Na Ca Sr Ba Mg cu

l00b 54 105 80 46 157 90 90 127 68 96 62 41 8 7 100

2.8 1 1.32 5.45 6.44 0.69 5.90 10.85 3.06 4.20 2.41 3.79 2.12 0.87 0.43 0.40 0.13

0.16(0.006)c 1.8l(0.246) 2.7 l(0.195) lSO(0.143) 0.69(0.109) 4.28(0.206) 3.38(0.150) 1.52(0.131) 1.19(0.069) 1.93(0.112) 1 .O l(0.078) 0.52(0.061) 0.38(0.101) 0.34(0.496) 0.25(0.388) 0.12(0.004)

Ag

co

Mn Ni Zn Fe Cr Ce La H

0.00(0.oO0)c 1.56(0.212) 0.20(0.014) lA(0.137) 0.22(0.035) 1 4.90(0.71 6) 0.54(0.024) 19.69(1.693) 5.81(0.339) 6.64(0.727) OSO(0.039) 3.08(0.362) 1.16(0.308) 0.34(0.496) 0.2q0.372) 0.32(0.011)

Adsorption time = 45 min, desorption time = 60 min; concentration of NO = 997 ppm; adsorption temperature = 273 K; adsorbent weight = 0.5 g; flow rate = 100 cm3 min- I . Concentration of NO = 1910 ppm. In parentheses are numbers of NO molecules per cation introduced into zeolites. a

alkaline earth and lanthanides. The qrev value decreased in the following sequence: transition metals > alkaline earths > lanthanides > alkaline metals > proton These results indicate that reversibly adsorbed NO is associated with transition metal ions, such as Cu and Co. For instance, in the case of the Cu-ZSM-5 sample, since both the qrev and qirr were proportional to the exchange level of copper, the ratio of qrev/Cuand qi,,/Cu was constant. From an IR study on the Cu-ZSM-5 sample, the dominant species of reversible adsorption was attributed for the most part to NO+ whereas those of irreversible adsorption were attributed to NO+, N02+, N02- and N03-.25-27 The amount of NO adsorption was also influenced by the zeolite structure. They also found that qrev/CU and qirJCu greatly changed with the zeolite structure and decreased in the following order: ZSM-5 > offretite/erionite > mordenite > L-type > ferrierite > X - t y ~ e . * ~ This order is consistent with increasing A1 content in the zeolite structure.

2.3 Alumina - Almost all NO, emitted from typical combustion processes is present in the form of NO. Since NO is unreactive and insoluble in aqueous solutions, a possible strategy is to oxidize NO to NO2 and then use wet sorption into alkaline solutions. y-Alumina and various alkaline compounds are potential candidates as alternative solid sorbents for dry NO2 removal systems operating at ambient t e m p e r a t ~ r e . ~The * . ~ active ~ sites on the alumina

3: Solid Sorbents for Catalytic NO, Removal

77

surface have been described as basic (hydroxy and 02-vacancies) and acidic (unsaturated A13+ ions as Lewis and protonated hydroxyl ions as Brrnsted) sites? The reactivity of transition aluminas to NO2 is significantly influenced by hygroscopic nature of the s ~ r f a c e At .~~ ~ ~relative ~ 50% humidity, for instance, activated alumina adsorbs water vapor to produce surface hydroxy ion species.33 Therefore, NO2 sorption in this case involves the reaction between NO2 and surface hydroxys, which is explained as a result of the following disproportionation of the surface OH-.

[OH-],

+ 3N02

-+

[NO3-],

+ HN03 + NO

(3.1)

The reaction is also confirmed by infrared spectral analysis, which shows NO2 absorbs to alumina as nitrate or nitrite species.34Despite such high sorption capacity for NO2, the use of y-alumina would be very limited in cases where SO2 was also present since SO, reacts readily with alumina and the reaction is irreversible. This is reflected by the higher decomposition temperature of aluminum sulfate (770 "C) compared to aluminum nitrate (ca. 200 "C). 2.4 Alkaline Solids - Impregnation of alkaline compounds improves the reactivity of the porous materials to NO,. Unlike adsorption processes, the reaction between NO, and alkaline solids is not only limited to the surface, but also spreads into the bulk, so that a large amount of NO, can be sorbed. The alkaline solids include carbonates and hydroxides of alkaline and alkalineearth metal^,^^-^^ calcium silicate,40Mg0,41942and these materials supported on porous oxides such as alumina. These inexpensive but readily available materials with efficient sorptive characteristics would be useful in NO, removal applications. A number of studies to control NO, emissions from jet engines report that a promising sorbent was obtained from MgO supported on However, the ease of a naturally-occurring brittle mineral, ~ermiculite.~~ degradation of the pellets and low surface area make them impractical. Considering its large surface area, cost, and chemical as well as thermal stability, y-A1203is a candidate support material for alkaline solids. Hoflund and c o - ~ o r k e r shave ~ ~ studied NO2 and NO sorption properties of KOH supported on y-A1203,which was prepared by impregnation and precipitation. Figure 3.2(a) shows typical effluent concentration curves of NO and NO2 for MgOhenniculite and A1203, which are measured in a conventional flow reactor. The NO2 concentration leaving the bed increased monotonically with time, and the apparent capacity for NO2 can be measured by analysing the areas under the curves. However, an NO concentration higher than that in the feed suggests that it is not a simple adsorption process; NO2 reacts with MgO to form NO as shown in the following reaction. p42

Thus, the net capacity of these materials for NO2 could be determined by subtracting NO formed from the apparent capacity for NO2. Similarly, the NO2 sorption onto y-A1203produced NO in accord with the reaction (3.1). Nelli and Rochelle35 also found that NO2 adsorbs on various nonalkaline

Catalysis

(b) n 3

'

-

*

*

*

'

-

-

*

-

'

*

.

-

-

€?NO

3'mm 0

100

200

Time (min)

300

Time (min)

Figure 3.2 Bed outlet N O and NO2 concentrations as Q function of time for (a) MgOl vermeculite, y-AI2O3 and (b) KOH-treated y-A1203 at 200 0C.36Reaction conditions: residence time Is, bed volume 25 cm3, 500 ppm N 0 2 , 50 ppm NO, N2 balance.

solids including y-alumina and suggested that adsorbed water and NO2 react to produce nitric acid and NO as follows: 3N02(g) + H20(1) -+ HNO3(l) + NO(g)

(3.3)

Figure 3.2(b) shows the effluent NO, concentration for KOH-treated y - a l ~ m i n a .All ~ ~ three beds yield very low outlet NO2 concentrations as compared to MgO/vermiculite. KOH-precipitated y-alumina showed the maximum NO2 removal of 1.1 mol mol-KOH - I . Treatment with KOH delays and reduces the formation of NO while increasing the amount of NO2 absorbed five-fold. In this reaction, potassium would be converted from hydroxide into nitrate and/or nitrite, which caused 56% loss in specific surface area. The addition of 3 vol% water vapor to the feed gas stream significantly enhances the adsorption. The NO,-sorbed KOHly-alumina can be regenerated by washing with water to remove all of the nitratehitrite, or by heating above 400 "C to decompose the nitratehitrite. 2.5 Transition Metal Oxides Containing Alkaline Earths - The NO,-adsorption properties of transition metal oxides have been studied with regard to their catalytic activity for conversion of N0.43-54Of these studies, the highest NO chemisorption was reported for CuO, NiO, Fe2O3, Fe304and some mixed oxides such as LaMn03. However, the amount of NO chemisorption was not sufficient (< 50 mg g-l) for practical use even though a wide variety of support materials were examined to enhance the surface area. As a result, the oxides containing alkaline elements were studied for NO sorption. Mixed oxides containing Ba species as well as other alkaline-earth, alkali, and

3: Solid Sorbents for Catalytic NO, Removal

79

100

80 E

Q Q

60

\

0

z

40

20

0 Time / h Figure 3.3 Variation of the concentration of NO at the exit of the reactor in the stepwise heating process. 56 Sample: YBa2Cu3 O,,, Reaction conditions: Space velocity 12 000 h- I , 50 ppm NO, 8% 0 2 , N2 balance.

lanthanide elements possess a characteristic affinity for gaseous NO; they produce nitrate or nitrite in the presence of NO and 0 2 . Most of the gassolid reactions between NO, and metal oxides are strongly temperature dependent; NO absorption and desorption corresponds to the formation of nitrate (low temperature) and decomposition of nitrate (high temperature), respectively. Layered cuprates containing Ba or Sr have been extensively studied for use as NO s ~ r b e n t s ? - ~Misono ~ and c o - ~ o r k e r s have ~ ~ . ~reported ~ the rapid uptake of NO as well as CO into YBa2Cu30,,.After pre-evacuation at 300 "C, the sample absorbed ca. 2 mol mol-oxide-' of NO at the same temperature (Figure 3.3). The NO molecules thus absorbed were almost completely recovered as NO by heating above 400°C. The absorption ability is strongly dependent on the pre-evacuation conditions; the NO uptake at 300°C increased to 2.3 mol mol-oxide-', and the recovered gas was a mixture of NO and N2 when the sample was pre-evacuated at 30°C. A large quantity of gaseous CO was also incorporated into YBazCu30,,, but the absorbed CO molecules cannot be recovered below 500 "C probably because of the stronger interaction. More striking is that, on exposure to NO/CO mixtures at 300°C, NO was taken up initially, then CO adsorption and N2 formation took place, suggesting that the catalytic reaction, 2 N 0 + 2CO -+ N2 + 2C02, proceeded over Y B ~ Z C U ~ OThe , , . ~absorption-assisted ~ catalytic reaction showed a highturnover frequency comparable to that over Rh/A1203 catalysts, which is one of most active de-NO, catalysts. NO absorption into Y B ~ ~ C U ~was O , , also examined by Arakawa and A d a ~ h i who , ~ ~proposed the following solid state reaction mechanism: 2YBa2Cu307 + 2 N 0

--+

3CuO + YzBaCuO5 + Ba(N02)z + 2BaCu02 (3.4)

80

Catalysis

Table 3.2 Removal of NO and NO2 by Ba-M-0 systems59 M

NO removal (YO) 0 2 1o%b

0 2 OYOa

Cr Mn Fe co Ni cu

8.0 2.8 6.9 18.6 1.4 100.0

18.6 7.9 10.7 33.3 2.5 100.0

NO2 removal (%) Crystal phase 0 2

17.0 30.2 15.1 34.5 21.9 100.0

BaCr04 BaMn03 + Ba3Mn208 BaFe03-x BaCo03-, NiO + BaNiO2 + Ba3Ni308 BaCuO2.1 + Ba2Cu30~+~

All samples were calcined at 750°C in air. a 0.1% NO, N2 balance, S.V.=6OOO h - l . 0.OYhNO, 10% 02,N2 balance, S.V.= 6000 h - I . 0.07%NOz, 10% 02, N2 balance, S.V.= 6000 h - I .

BaCuO2 + 2 N 0 + CuO + BaN203

(3.5)

Reaction (3.4) can be confirmed by X-ray diffraction after the termination of NO absorption of ca. 2.5 mol mol oxide-' at 3OO0C, which showed that YBa2Cu307 disappeared with simultaneous appearance of Y2BaCuOs and BaCu02. Actually, however, since NO absorption into a single BaCuO2 phase is negligible, reaction (3.5) should be expressed as a gas-solid reaction with a Ba-Cu-0 site to produce barium nitrate or nitrite, which is accelerated in the presence of CuO. The author and c o - ~ o r k e r s ~ have ~ - ~studied ~ various Ba-M-0 systems for the removal of NO and NO2 (Table 3.2). These mixed oxides showed NO absorption activity to some extent at 200"C, which was accelerated in the presence of 0 2 . The highest activity for NOIN02 absorption was observed for the Ba-Cu-0 system, which completely removed 0.09-0.1% of NO, after 20 min of the reaction.59In the cases of NO/N02 removal in the presence of 02, a large amount of NO, was liberated from samples above ca. 500°C. The amount of liberated NO,, which was over 10 times larger than that estimated from monolayer adsorption onto the surface, was equal to the cumulative amount of NO removed. The Ba-Cu-0 system contained two different types of nonstoichiometric compounds, BaCu02. and B ~ C U O and ~ . ~Cu0.59*61 , The NO absorption results from the solid-gas reaction between NO and these binary oxides to produce Ba(N03)2/Cu0 mixtures. The role of CuO in the gassolid reaction between NO and B ~ C U Owas ~ . studied ~ by the catalytic reaction of NO over CuO as shown in Figure 3.4.59When an NO102 mixture was fed to CuO, NO was converted into NO2 with a rise in temperature up to 250°C. CuO had no contribution to the absorption of NO or NO2. The BaCu02.1 sample, which showed much less catalytic activity for NO oxidation, completely removed dilute NO2 (O.OSO/) up to 400 "C. These results indicated that the NO absorption into BaCu02.1 follows the catalytic NO oxidation by CuO as shown in Figure 3.5. The adsorption rate is thus enhanced in the presence of 0 2 , but significantly inhibited by C02, which reacts with Ba species to produce an inactive surface layer of BaC03. This problem could be overcome by employing manganese oxide (MnO2 or Mn203) as a preoxidation catalyst. The

3: Solid Sorbentsfor Catalytic NO, Removal

150

250

350 150

81

250

33

TemperaturePC Figure 3.4 (a) NO oxidation over CuO (a)and B a C ~ 0 2(O), . ~ and equilibrium conversion (solid line);59 reaction conditions: S.V. 6000 h-I, 0.09%NO, 10% 0 2 , N2 balance. (b) NO removal by BaCu02.1 (W) and CuOlBaCu02.1( 0 )and NO2 removal by BaCuO2.1 (a);reaction conditions: S.V. 6000 h-I, 0.09% NO or 0.08% N 0 2 , 10% 02,N2 balance. NOIN02 removal after 20 min of the reaction is shown.

cuo+

0.4502 2N02

2N0-1-02

I

Figure 3.5 Schematic illustration of NO absorption into CuOIBaCuO2.1.59

higher catalytic activity of MnO2 for NO oxidation seems effective in reducing the C02 i n h i b i t i ~ n . ~ ~ Pitchon and c o - w o r k e r ~have ~ ~ studied ~~ the removal of NO, by sorption/ desorption on barium aluminate in comparison with a bulk BaO sample. The NO, sorption in both cases is based on the formation of nitrate on the surface, so that only NO2 can be removed from the gas phase. The bulk BaO also reacts with C02 to produce very stable surface carbonate, which deactivates the sites for nitrate formation. By contrast, Ba0-A1203 contained no strongly bound carbonate, therefore allowing the formation of nitrates, which can be thermally decomposed. The NO2 sorption into Ba0-A1203 calcined at 800 "Cis dependent upon the specific surface area, and the largest reversible sorptioddesorption was 24 mg g- *.They have also reported NO, adsorptioddesorption capacities of a series of perovskite oxides (AB03) with A = Ca, Sr, Ba and B = Sn, Zr, Ti? In the presence of excess 02, only NO2 was selectively adsorbed whereas NO is never trapped. The NO2 uptake decreased with the sequence Ba > Sr > Ca for A, and Sn > Zr > Ti for B, suggesting that NO2 sorption is governed by the B-0 bonding energy and the electropositivity of the element A. The presence of water vapor not only increases the amount of NO2 trapped, but also shifts the

82

Catalysis

temperature of desorption towards higher temperature. The beneficial effect of water vapor upon NO, storage capacity can be seen in many oxide systems containing alkaline or alkaline-earth metals. It seems that the sorbed water molecules stabilize the nitrate ions through hydration and facilitate diffusion of the gas-solid reaction. Also, superficial hydroxides formed on exposure to water vapor may take part in the NO, sorbing reactions.66

2.6 NO, Intercalation Compounds - The author and c o - ~ o r k e r s ~have ~-~~ studied the NO-sorption properties of the substituted double layered cuprate, La2-,Ba,SrCu206. Unlike reported solid oxides containing Ba in the preceding section, the layered cuprate absorbs NO via intercalation into the layer structure without precipitating Ba(N02)2 or Ba(N03)2. Figure 3.6 compares the HREM structural image with the incident beam parallel to the l a ~ e r . ~ ~ ? ~ ' The structure in Figure 3.6(a) is characterized by stacking layers with a periodicity along the c-axis close to 1.2 nm, showing the contrast of two Cu05-sheets (dark) between interlayers containing Ba (bright). In the image after NO uptake [Figure 3.6(b)], the stacking periodicity expanded to 1.4 nm. Since the image of the structure unit sandwiched between the interlayers is basically unchanged, the lattice expansion can be attributed to an increased thickness of this interlayer region. The increased bright contrast at the interlayer probably indicates the incorporation of NO molecules between the copper planes, which is consistent with the result of XRD. Water and NO are accommodated together into the Ba-containing interlayers between copper planes. Such characteristic reactivity to NO is associated with the superstructure including the ordering of Ba and oxygen v a ~ a n c i e s . Another ~~?~~ remarkable feature of the NO, intercalation compound is concerned with the thermal deintercalation process as shown in Figure 3.7.67970971 Part of the NO incorporated into La2 _,Ba,SrCu2O6 is liberated dissociatively as 0 2 at > 600°C and N2 at > 800°C. This is in striking contrast to the other NO absorbents based on the formation of Ba(N02)2 or Ba(N03)2, from which absorbed NO is completely recovered as mixtures of NO/N02. The qualitative relationship between NO uptake and each desorption suggests the presence of two different NO sites in the layer structure. One is the site bound to Ba, where NO is converted into NO2- ions, from which NO can be reversibly desorbed on heating (< 600°C). Another site is an anion vacancy in the Cu05 pyramid layer, being occupied by NO strongly bonded to Cu. Instead of this NO ligand, lattice oxygens are eliminated as 0 2 above 600°C. A further increase of temperature results in cleavage of the N-0 bond and subsequent elimination of N2 while leaving the NO oxygen in the lattice. This system was applied to thermal swing sorptioddesorption cycles of dilute NO in a conventional flow reactor between 250 "C (sorption step) and 500-700 "C (desorption step).71 Stable and reversible sorption/desorption cycles were successfully achieved in the presence of water vapor (0.7%), which readily regenerates the interlayer hydroxys in a sorption step. In this case, the intercalated NO was immediately condensed out of the stream with considerable evolution of N2 at the beginning of each desorption step. Unfortunately, 972

3: Solid Sorbents for Catalytic NO, Removal

83

Figure 3.6 HREM images of Lu,,5Bao,5SrCu206on exposure to (a) water vapor (20 kPa, 60 "C,48 h) and subsequenify to (b) N O (1 3 kPa, 250 "C, 160 rnir~).~~ Incident electron beam parallel to layers. The values shown in images are stacking periodicity of the layer structure.

however, the thermal-swing sorptioddesorption was significantly inhibited in the presence of C02.71

2.7 Non-alkaline Solid Oxides - Metal oxides containing alkaline earth or rare earth elements are active for removal of a large amount of NO, absorption, whereas deactivation of absorbents is expected because of the formation of surface carbonate in the presence of COZ. The alkaline earth elements serve as strong absorption sites for C 0 2 due to their strong basicity. Thus, attempts have been made to develop materials for NO, absorption which do not contain care earth and alkaline earth components, One possible way to design such solids is the combination of NO oxidation catalysts and oxides with intermediate basicity. have studied the reversible sorption of nitrogen Eguchi and oxides in Mn-Zr oxide, prepared by coprecipitation of nitrate precursors, in the presence of 0 2 . The sorption capacity at 200 "C was quite dependent upon the

Catalysis

84

lb 200 I

N2-\

NO

400 I

600

800

I

I

600

800

I

a 200

400

Desorptiontemperature I *C Figure 3.7 TPD profiles from L.ul.4Bao,aSrCu,0, (a) before and (b) afer absorbing 0.5 mol/mol of NO at 250 "C.'* Heating rate, 10 "C min- in He.

'

Mn :Zr ratio and showed the maximum (0.133 mol mol-Zr - I) with Mn :Zr = 1, where the sample is composed of an amorphous phase with a surface area of 185 m2 g-*. From FTIR study, NO is oxidized by manganese oxide to NO2 and then sorbed as NO3-. The removal of NO was not significantly influenced by the presence of C02 and was promoted by the presence of H20. Figure 3.8 shows the typical concentration profile in the effluent gas from Mn-Zr oxide.75 At 200°C, NO in a gas feed was completely removed for the first 60 min. Subsequent heating up to 400°C in a flowing He led to reversible NO, desorption. In addition to the thermal desorption, the reduction treatment is an Figure 3.9 shows alternative method to regenerate the NO,-sorbed the effluent NO, concentration when the atmosphere was switched to a reducing atmosphere after NO absorption at 300°C. Sharp desorption of NO, was observed immediately on exposure to C3H8 or C2H50H.The recovery of NO, was 86% (c3H8) and 55% (C2HSOH),whereas the remainder was N2, which was produced as a result of the following reactions;

C2HSOH + 2N03(s) + N2 + 3H2O + 2CO2

(3.7)

It is noted from the comparison of Figures 3.8 and 3.9 that the desorption proceeds more rapidly in a reducing atmosphere than in the case of He. Cyclic sorption-desorption characteristics of Mn-Zr-O and Pt/Zr02-A1203 were also examined with dynamic changes in equilibrium oxygen partial pressure of < p02 < 1 by use of C3H8-02-NO-He.77 They have successfully

3: Solid Sorbents for Catalytic NO, Removal

85

.~ 120

T i Imin

180

Figure 3.8 Amount of NO sorbed in Mn-Zr oxide ( M n : Zr = 1) and desorbed at 400 "C. Hatched areas indicate amounts of N O sorbed and desorbed.75 Sorption conditions: 900 ppm NO, 10% 0 2 , He balance. W/F ~ 0 . gs 6 ~ r n - ~200 , "C, 60 min. Desorption conditions: from 200 to 400°C and at 400°C in He, W/F = 0.6 gs crnW3.

0

10

Timchin

20

30

Figure3.9 Amount of N O sorbed in Mn-Zr oxide ( M n : Z r = 1) and desorbed in reducing atmosphere. Hatched areas indicate amounts of N O sorbed and desorbed.75 Sorption conditions: 900 ppm NO, 10% 0 2 , He balance, 300 "C, 60 min., W/F = 0.6 gs cm-3. Desorption conditions: 1500 ppm C3H8or 1500 ppm C ~ H S O HHe , balance, 300 "C,W/F = 0.6 gs ~ m - ~ .

demonstrated bifunctional catalysis that can promote catalytic NO, reduction in a reducing atmosphere and sorptive NO, removal in an oxidizing atmosphere. More recently, the author and co-workers have studied NO adsorption on various binary oxide systems, which possess the combination of manganese oxide as an NO-oxidation catalyst and lanthanide oxide as an NO,-sorbing

Catalysis

86

0

2000

3Ooo

4Ooo

I0

lime on streadsec Figure 3.10 EfJIuuent NO, concentrationfrom MnO,-CeO, system. (a) 0.08% NO, He balance, at 30 "C;(b) 0.08%NO, 2% 0 2 , He balance, at 30 "C;(c) 0.08%NO, 2% 0 2 , He balance, at 150 "C;(d) 0.08%NO, 2% 0 2 , 10% CO,,0.7% H20, He balance, at 150 "C, W/F = 0.24 gs ~ r n - ~ . 78979

component. The formation of a fluorite-type solid solution in the Mn0,Ce02 system was found to be effective in accelerating NO s o r p t i ~ n . ~ ~ ~ ~ Figure 3.10 shows the typical eniuent concentration profiles of N0,from Mn0,-Ce02. The NO, removal is most efficient at ambient temperature in the presence of 0 2 , and inhibition by C02 is not significant. The cumulative NO, uptake was increased by decreasing the reaction temperature and/or by increasing the O2concentration, indicative of chemisorption via oxidation of NO/N02. Mn oxide appears to promote the oxidative adsorption of NO as bidentate, monodentate, and ionic nitrates. Since ca. 60% of sorbed NO, at ambient temperature is readily desorbed by heating up to 200"C, the sorbent after saturation can be easily regenerated by employing a thermal swing operation as shown in Figure 3.1 l(a). In a thermal swing process, isothermal heating at 30 "C for a sorption step and at 500 "C for a desorption. step was repeated with supplying gas mixtures of 0.08% NO, 2% 0 2 , and He balance. In the sorption step, almost complete sorptive removal of NO, continued. At the end of each sorption step, the temperature was raised to 500°C and NO, stored in the solid was immediately condensed out of the stream. The reversible sorption-desorption cycles could be maintained with no signs of deactivation. The NO,-sorbed Mn0,-Ce02 could also be regenerated by injection of H2 pulses at 150°C as shown in Figure 3.11(b).80.8' After saturation of NO, sorption at 150"C, the injection of H2 pulses immediately gave a steep drop of the effluent NO, concentration, which was accompanied by the evolution of N2. The regenerative capacity of NO, sorption increased with a number of the H2 pulses and finally exceeded 50% of the maximum NO, uptake (0.15 mmol g-l). Assuming the stoichiometric reaction, 2N02 + 4H2 -+ N2 + 4H20, this corresponds to ca. 70% selectivity of H2 in every injection to the reduction of NO, stored.

3: Solid Sorbentsfor Catalytic NO, Removal 500.

e

87

a

o

0.20

$ 0'16 0.12

:

CI

c

3 0.08 ..........,......,

E O.O4

0.16

20

40

60 80 Time on streamlmin

100

120

Figure 3.1 1 Cyclic regeneration of NO,-sorbed Mn0,-Ce02 system. 78 - 8 1 (a) Thermal swing NO, sorption-desorption cycles over MnO,-CeO2, 0.08%NO, 2% 0 2 , He balance, W/F=0.50 gs cm-3; (b) effect of Hzpulses on the efJluent NO, from 1 wt% PdJMnO,-CeO2 at 150 "C.Each H2 pulse (shown as arrows) was 1.0 cm3, gas feed 0.08 vol% NO, 2 vol0/0 02,He balance, WJFz0.24 gs cm-3 (50 cm3 min-I).

2.8 Heteropoly Compounds - The NO, sorption on other inorganic reactions has been reported. Belanger and have recently studied the sorption of NO2 on heteropoly acid, 12-tungstophosphoric acid (H3PW 12040-nH20, HPW). Although the solid acid has a non-porous low surface area, NO2 penetrates easily into the bulk of this ionic solid as do other gaseous polar molecules such as ammonia, pyridine and m e t h a n ~ l . While ~ ~ . ~ these ~ NO2 molecules react with the hydrogen-bonded water molecules in the structure to form HN03, which is released to the gas phase, an equal number of protons are retained in the solid. In contrast, nitric oxide is not taken up by the solid HPW in the absence of NO2. However, when NO is contacted with HPW containing previously sorbed NO2, a molar quantity of NO equivalent to that of the adsorbed NO2 is adsorbed. They have also revealed that the sorption of NO2 into a microporous derivative, ammonium tungstophosphate (NHPW)

Catalysis

88

1 1 2 ' 3 1 4 ' 5 ' 6 ' 7 ' 8 ' 9 ' PUCSENUMBm

14

12 10 8 6 4

2

0

PULSENUMBER

Figure 3.12 Concentration of N2, N02, N20, 0 2 and NO for each pulse of NO2 in the efpuent of a microreactor containing stoichiometric NH4P W held at 150 'Cg3 Mass in reactor 0.075 g, pulse size 16.5 pmol of NO2, HefIow 15 cm3 min- I .

was accompanied by formation of N2, with ammonium ions (Figure 3.12).83

0 2

and N20 as a result of interactions

WS+-O-N02NH4++ W6+-0 + [NH4N02] + N2 + 2H20 + WO

(3.8) This suggests that N2 is desorbed into the gas phase while 0 2 apparently participates in a surface oxidation process. The formation of N20 appears to result from the decomposition of ammonium nitrate during the process.

3

Regeneration of NO, Sorbent

Unlike liquid systems, it is extremely difficult to move solid sorbents from one place to another without damaging the particles. Therefore, continuous gas

3: Solid Sorbentsfor Catalytic NO, Removal

P1

Pressure

89

Pz

Figure 3.13 Principle of pressure or thermal (temperature) swing adsorption'.

separation requires at least two independent sorbent beds in parallel, which are alternatively used for sorption and regeneration steps3 The resultant mechanical complexity limits the practical application of those processes, but recent attempts have been made to overcome this problem. The solid-NO, interactions described in the preceding sections obey chemical equilibria which depend on temperature, pressure, and atmosphere (reducing or oxidizing). Thus, the regeneration of NO, sorbents can employ thermal (temperature) swing, or pressure swing operations, or cyclic oxidation-reduction treatment -3,86987

3.1 Pressure Swing Process - Figure 3.13 shows schematically the principle of pressure and thermal swing adsorption (PSA and TSA, respectively).' In the PSA process, adsorption takes place at high partial pressure and desorption takes place at low partial pressure. Pressure swing adsorption (PSA) has widely been applied to various processes, such as the separation of oxygen or carbon dioxide from air and purification of hydrogen. The application of this process to remove NO, requires sorbents with a high capacity for reversible sorption. However, for conventional adsorbents, such as activated carbons, silica gels and zeolites, the NO, adsorption capacity is not sufficient for practical uses. Figure 3.14 represents the typical breakthrough and elution curves when the adsorption and desorption runs were repeated.2s In Figure 3.14, the solid and broken lines indicate the concentration of NO with and without adsorbent, respectively. The amount of NO adsorbed corresponds to the area denoted as a,, whereas that of NO desorbed is b,. When a, is to be equal to b,, reversible sorption is possible.

Catalysis

90

0

30

60

90 120 Time / min

150

180

210

Figure 3.14 Typical breakthrough and elution curves of NO:25 solid line, with adsorbent; broken line, without adsorbent; Ca initial concentration of NO; a,,, N O adsorbed; b,, N O desorbed.

3.2 Thermal (Temperature) Swing Process - The TSA process utilizes differences in sorption capacity at different temperatures (Figure 3.13). The sorbent bed operates isothermally during sorption steps, whereas, in regeneration steps, the bed is heated to give energy to desorb NO,r or to decompose NO,containing compounds. The temperatures for sorption-regeneration steps are cycled by use of external heater or by alternating feed gas temperatures. This process is quite useful for thermostable sorbents, the NO,-sorption isotherm of which is strongly temperature-dependent. Since most NO,-solid reactions, with exception of simple physisorption, are strongly dependent upon the reaction temperature, a TSA process is expected to give rise to much larger NO, sorption capacities than PSA processes. On the other hand, heating and then cooling of a solid sorbent clearly requires much more time than changing the pre~sure.~ 3.3 Reduction-Oxidation Cycles - A oxidation-reduction cycle can be applied not only to regenerate the sorbents but also to reduce any accumulated NO, to N2. In an oxidizing atmosphere at relatively low temperatures (< 300 "C),NO is converted into N02, which is much more reactive. NO2 then reacts with solids to produce nitrate or nitrite precipitates. In a reducing atmosphere, however, nitratehitrite will be decomposed and reduced to N2 over suitable catalysts such as noble metals. This combination of adsorption and catalytic reduction is now attracting attention as a novel application of NO,-sorption process. Several examples can be seen in commercialized deNO, processes as described in the following section.

4

Practical Applications of NO, Sorption

4.1 Flue Gas Treatment for Stationary Sources - This section introduces several examples of commercially available de-NO, processes using solid

3: Solid Sorbents for Catalytic NO, Removal

91

sorbents. The NOXSO process has been developed as a dry regenerable flue gas treatment system by use of solid sorbents that simultaneously remove 90% of SO2 and 70-90% of NO, from flue gas generated from a coal-fired c o m b u s t ~ r . ~The ~-~ sorbent ~ is prepared by spraying Na2C03 solution onto the surface of y-alumina. Both sodium and alumina contribute to the capacity to absorb SO2 and NO, from flue gas as described below: Na2C03 + A1203 + 2NaA102 + C02 2NaA102 + H20

+ A1203 2NaOH + SO2 + 1/202 + Na2S04 + H2O 2NaOH + 2 N 0 + 3/202 -+ 2NaN03 + H20 2NaOH + 2N02 + 1/202 + 2NaN03 + H 2 0 + 2NaOH

(3.9) (3.10) (3.11) (3.12) (3.13)

The spent sorbent can be regenrated by heating to decompose the nitrate, and the resultant concentrated stream of NO, is returned to the boiler with combustion air. No significant increase of NO, concentration in the boiler flue gas results because of the reversibility of NO, formation in the boiler. The SO, removed is reduced by CH4, CO or H2 to produce condensed SO2 and H2S, which are converted into elemental sulfur. The sorption is conducted in a fluidized-bed adsorber and the spent sorbent is pneumatically transported in batches from the adsorber to a fluidized-bed sorbent heater and then to the moving-bed regenerator, where the used sorbent is regenerated by the reducing gas. The combination of catalysts and a NO, sorbent can be also seen in the SCONO, process for flue gas treatment of gas-fired turbine^.^' The process utilizes a single catalyst for the removal of both NO, and CO and their reduction by alternating two different cycles. In oxidation-absorption cycles, the SCONO, catalyst simultaneously oxidizes CO to C02, hydrocarbons to C 0 2 and H20, NO to NO2, and then absorbs NO2 onto its surface through the use of a potassium carbonate absorber coating as shown below.

co + 1/202 + c02

(3.14)

NO + 1/202 + NO2

(3.15)

CH20 + 0

2 + C02

+ H20

2N02 + K2CO3+ C02 + KN02 + KN03

(3.16) (3.17)

During this cycle, the potassium carbonate coating reacts to form potassium nitrite and nitrate, which are then present on the surface of the catalyst. When all of the carbonate absorber coating on the surface of the catalyst has reacted to form nitratehitrite mixtures, NO, will no longer be absorbed and the catalyst enters the regeneration cycle. The regeneration of the SCONO, catalyst is accomplished by passing a dilute hydrogen reducing gas across the surface of the catalyst in the absence of oxygen as shown below:

Catalysis

92

T

I

Stoichiometric A@

M: NOx storage component

Reduced to nitrogen Figure 3.15 Schematic mechanism of the NO, storage-reduction catalyst.' I 7

KN02 + KN03 + 4H2 + C 0 2 + K 2 C 0 3+ 4H20 + N2

(3.18)

The hydrogen gas reacts with nitrites and nitrates to form H2O and N2. Carbon dioxide in the regeneration gas reacts with potassium nitrites and nitrates to form potassium carbonate, which is the absorber coating that is on the surface of the catalyst before the oxidation-absorption cycle begins.

4.2 Automobile NO, Storage Catalysts - One excellent application of a NO, sorbent in a catalytic de-NO, process can be seen in the automotive NO, storage catalyst, which is first developed by T o y ~ t a for ~ ~gasoline - ~ ~ engines operated at lean-burn conditions. Since then, many researchers have studied this type of new strategy with the use of NO,-sorbing material^.^^-^*^ In a lean burn gasoline system, common three-way catalysts are able to efficiently oxidise HC and CO but are unable to reduce NO,. On the NO, storage catalyst (Figure 3 . 1 9 , however, NO, is stored on the catalyst during lean-burn conditions. The NO, storage catalyst usually contains noble metals for oxidation of hydrocarbons and CO, and a NO,-sorbing compound, typically alkaline earth elements such as Ba, which are capable of adsorbing NO, as follow^:^^-^^ NO + O2 + Ba

Pt -+

Ba-NO,

(3.19)

To regenerate the catalyst, short periods of rich conditions are employed during which exotherm leads to an increase of the catalyst surface so that the stored NO, is thermally released and subsequently reduced to N2 on the noble metal catalysts. The reactions in this case are: HC + 0

Pt 2 -+

Ba-NO, NO,

C02 + H20 + heat

+ heat -+

+ HC + 0

Pt 2 +

Ba + NO,

N2 + C02 + H20

(3.20) (3.21) (3.22)

The sorption-regeneration cycle is repeated periodically by controlling the air : fuel ratio to lean or rich conditions. One problem concerning this process

3: Solid Sorbentsfor Catalytic NO, Removal

93

is catalyst deactivation in the presence of l 6 This system works without any deterioration in Japan where extremely low sulfur fuel is readily available. In Europe and the US, however, owing to the higher fuel sulfur causes, the resultant SO, in exhaust competes more strongly than NO, for the alkaline site on the catalyst, leading to loss in NO, storage capacity. The development of sulfur-resistant NO,-storage catalysts is now in progress. I 17*1 *

5

Conclusion

Because of the limited sorption capacities of solid NO, sorbents, their application has been limited to the treatment of extremely low levels of NO, in exhaust gases. The recent development of useful regeneration processes has opened novel applications of sorptive NO, removal. There is great interest in using solid sorbents as an alternative to catalytic processes for automobiles and flue gas treatment, because solid-NO, interactions to produce nitrates allow these processes to operate in a strongly oxidizing atmosphere at low temperatures. As pointed out in the latter part of this article, more recent studies of solid sorbents are directed toward the combination with de-NO, catalysts. Catalysts, which have played a key role in de-NO, technology, may be limited in their ability to satisfy the increasingly stringent requirements for pollution control. With the use of solid sorbents, dilute NO, can be removed and concentrated to allow the eflicient reduction in a downstream catalytic converter. Besides such a dual-bed operation, NO,-sorbing catalysts having integrated structures of NO, sorption sites and catalytically active sites would be possible. Provided that these bifunctional surface structures are stable under reaction conditions, the combination of NO, sorption and catalysis may give rise to a synergistic effect, thus paving the way to development of NO,sorbing catalysts for selective de-NO, processes at low temperatures. On the other hand, it should be also noted that the practical application of sorptive NO, removal technologies is sometimes limited by the interactions of the sorbents with SO,, which is an inevitable component of many exhaust gases from combustion sources. Acidic gases, such as SO, and CO2, compete for basic sites on solid sorbents; the higher stability of sulfates and carbonates compared with nitrates will cause rapid deactivation. The hindrance effects of COZ can be avoided by the optimum selection of NO,-sorbing materials, whereas those of SO, are much more severe. Overcoming this problem will be a key to the use of catalytic NO,-sorbing materials.

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2 3

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Catalysis

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6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30

31 32 33 34

35 36 37 38 39

3: Solid Sorbents for Catalytic NO, Removal 40 41 42 43 44 45 46 47 48 49 50 51 52

53 54 55 56 57 58 59

60 61 62 63 64

65 66 67 68

95

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4 Partial Oxidation of Light Alkanes in Short Contact Time Microreactors BY P. AGHALAYAM, Y.K. PARK AND D.G. VLACHOS 1

Introduction

Catalytic partial oxidation of natural gas has now been known for more than half a century, and recent reviews have been reported. 1-3 However, the success in catalytic partial oxidation of methane (a major component of natural gas) to syngas in a short contact time microreactor with Pt or Rh catalyst^^-^ has pioneered an interesting alternative to the commercial process of steam reforming over Ni catalysts. Considerable interest has subsequently been triggered, as syngas has important industrial applications,8 but its formation is the most expensive step in natural gas conversion to liquid products. The oxidative dehydrogenation of ethane to ethylene in short contact time reactors has created further interest in this The latter gas is the basis for polyethylene, a process of tremendous industrial significance. Since autothermal operation of both syngas and olefin production through partial oxidation is p ~ s s i b l e , ~and ~ ~ ~there " is no need to burn natural gas to heat up the reactor as happens in conventional reforming, these processes lead both to significant energy savings and lower COZ emissions as compared with traditional methods. Obviously, the environmental and energy impact, along with the potential for further optimizing conversion and selectivities in partial oxidation of light hydrocarbons in short contact time reactors opens up new horizons for chemical synthesis.l 2 Following up on the pioneering work of Schmidt and c ~ - w o r k e r s , ~ ~ ~ ~ ~ numerous experimental studies have been carried out to delineate the reaction mechanism leading to optimum reactor performance. Because these reactors operate under rather severe experimental conditions (short contact times, high temperatures, high pressures and fuel-rich reactant mixtures), extrapolation of kinetic data from surface science experiments does not adequately describe the processes taking place under these conditions. To ensure compatibility with downstream processes and because natural gas is easier to compress than hydrogen, it is desirable to carry out such reactions at high pressures, up to 100 atm. It has previously been shown13 that gas-phase methane ignition temperatures drop monotonically with increasing pressure (even though this result is intuitive, batch experiments show that this is not true for all fuels14). At pressures of ca. 10 atm, the homogeneous ignition temperatures are below 1300 K (typical operation temperatures of monolithic short contact time Catalysis, Volume 15 0The Royal Society of Chemistry, 2000

98

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

99

microreactors), and thus, it is possible that homogeneous gas-phase reactions contribute to the overall reaction paths. This is almost certain at higher pressures and for larger hydrocarbons for which the onset of gas-phase reactions is observed at lower temperatures. Construction and operation of high pressure microreactors are difficult due to high cost and safety considerations. Furthermore, experimental reactor optimization can be time-consuming, counter-intuitive due to strongly non-linear responses, and potentially dangerous due to the possibility of explosions as flammability limits are crossed; in particular, the fuel-rich flammability limit expands with increasing pressure. l4 These considerations underline the need for predictive mathematical modeling based on fundamental fluid mechanics, multicomponent transport and detailed chemistry as a tool in guiding experiments and optimization. The desire to extrapolate models to conditions with no available experimental data (e.g. high pressures) and to deal with selectivities of multiple paths underscores the need for 'elementary' reaction mechanisms which offer the best promise to achieve such a goal. While a vast set of experimental data has been obtained, apparently contradicting features have been reported. As examples, there has been debate on whether direct or indirect partial oxidation of methane occurs in syngas reactors.'~~ Furthermore, in experiments on the influence of flow velocity on syngas selectivity, a minimum in the H2 selectivity as the flow velocity is increased was observed,15 or a monotonic decrease in H2 selectivity'6 was found, although both experiments were conducted at similar operating conditions. Mathematical models have advanced but are not sufficiently predictive to aid reactor scaleup by modeling the interplay of chemistry with heat and mass transport. This is in part due to strong heat and mass transfer effects in these reactors that disguise kinetic processes. This, along with highly complicated flow patterns (possible lack of fully developed flows in these very short reactors), make modeling and understanding of these processes challenging. This review summarizes recent experimental and theoretical advances in catalytic partial oxidation of small alkanes, focusing on platinum and rhodium (with some comparison to nickel) and attempts to provide a unifying mechanistic interpretation of such processes along with current limitations of fundamental understanding. The global chemical reactions for methane oxidation discussed in this review are listed in Table 4.1.

2

CH4 Partial Oxidation on Platinum and Rhodium Catalysts

Platinum is used in various catalytic processes, commonly in complete combustion (e.g. the automotive catalytic converter). For partial oxidation, a great deal of research has been spurred by the recent discovery that different supported metals (Ru, Pd and Rh) and lanthanide oxides can give high selectivities to syngas (a mixture of carbon monoxide and hydrogen) by the partial oxidation of methane.'6v17 A further step forward has been the discovery that excellent conversions and selectivities to syngas can be obtained

Reaction - 192.0 - 8.5

Complete combustion Direct partial oxidation Steam-reforming COz or dry-reforming Reversed Boudouard Water-gas shift 49.0 59.0 - 130.0 10.0

AH (kcal mol - I )

Name

Table 4.1 Global chemical reactionsfor methane oxidation

6.8

- 131.9

34.0 40.9

- 20.6

- 191.5

AG (kcal mol- ')

0 0

L

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

101

in autothermal, short contact time reactors using Pt, in various reactor configuration^.^ Owing to higher selectivity to syngas and enhanced robustness and stability compared to Pt, a number of recent studies have focused on Rh5 and Pt-Rh alloys as well. Various catalyst and reactor configurations have been explored, including metal gauzes (single and multiple layer^),^. 8-20 foam and extruded monoliths with the catalyst impregnated in the pores,5-7~10~*5~21-28 sponge^,^^^^^, and fluidized bed^.^^^^' Table 4.2 presents a compilation of representative results of various research groups on Pt and Rh catalysts. We discuss next methane conversion and selectivity to syngas obtained with these catalysts under different operating conditions and reactor geometries. 2.1 Unsupported Catalysts - In a detailed study of a single platinum gauze with a CH4:02:Ar ratio of 2 : 1 : 10 (stoichiometric to CO and H2), the catalyst was found to be poor for the production of synthesis gas.18 In particular, even temperatures as high as 1200 K with 0.1-0.5 ms contact time led to relatively low oxygen conversion (50-75%) and hydrogen selectivities (< 10%). Furthermore, the selectivity to CO and H2 dropped with time-onstream, whereas the oxygen conversion increased. Examination of temperature profiles as a function of catalyst use indicated that although CO is a primary product, formed at the position of the gauze, it was readily oxidized to C02 near the exit of the reactor, presumably catalysed by the Pt evaporating from the gauze that coated the downstream reactor walls. Alloying Pt with Rh leads to overall better stability, conversions and selectivities. In particular, with a single Pt/lO%Rh gauze and at similar conditions as above, up to 33% methane conversion, with 80% oxygen conversion, 96% selectivity to CO, and 34% selectivity to H2 have been found.19 No significant changes in the results were observed with time-onstream. Higher selectivities and conversion were achieved with higher reactant inlet and gauze temperatures. With multiple layers of gauzes, almost complete 0 2 conversion was achieved, along with improved selectivities to CO and H2 (90 and 40%, respectively), as shown by Schmidt and c o - ~ o r k e r s .Furthermore, ~ at least three layers of gauze are necessary for the reaction to go to completion, and the conversion of methane decreases as the inlet fuel :oxidant ratio increases (consistent with the findings of Heitnes Hofstad et aZ.19).For a relatively small increase in flow velocity (ca. 8-20 cm s-l), only a small decrease in reactant conversion and a slight increase in CO and H2 selectivities were ~ b s e r v e dIn .~ comparison to the results obtained with this small variation in flow velocity, the strong dependence of conversion and selectivities on operating temperature and the number of gauze layers indicates that for a single wire gauze reactor, optimal conversion of reactants and selectivities to syngas are mainly limited by the catalyst surface area available for reaction. Among other unsupported catalysts, porous platinum and rhodium sponges (made of pure metal) have been examined by Mallens et aZ.,29930in transient kinetic studies, also known as Temporal Analysis of Products (TAP). A fixed

Not reported.

Sponges Rh Pt Single gauze Pt Pt/lO%Rh Gauze pack Pt/lO%Rh Foam monoliths 4 wtY0 Pt 4 wt% Rh 0.1 wt% Pt 0.1 wt% Rh 0.3 wt% Rh Fluidized bed Pt Rh

Catalyst

1 .o 1 .o 40 40 75

-1200 -1 200

-1050 -1050

-1 00 -1 00

0.5

NRt

-1 100 -1 100 -1 200

0.21 0.21

100

1 00

Contact time (ms)

-1200 -1 200

1000 1000

(K)

Temperature

2.0 2.0 2.0

-2

56%N2 56% N2

10% N2 10% N2 77% Ar 77% Ar 90% He

65%N2

1 .o

1.6 1.6 2.0

77% Ar 77% Ar

None None

(“/3

Diluent

2.0 2.0

2.0 2.0

CH4:0 2

50 85

90 96 20 80 80

95

21 33

1 00 95

X ( CH4) (Yo)

70 90

95 95 10 90 96

90

88 96

90

100

S( CO) (Yo)

60 95

80 85 5 90 90

Bharadwaj and Schmidt3’ Bharadwaj and Schmidt3’

Schmidt and Schmidt and Heitnes Hofstad et al. l9 Heitnes Hofstad et al.19 Slaa et

Hickman and Schmidt’

Heitnes Hofstad et Heitnes Hofstad et al.19

9 34 42

Mallens et Mallens et

Reference

100 95

V O )

S( Hz)

Table 4.2 Selected summary of the experimental data for syngas production under different reactor operating conditions and reactor geometries

L

0 t3

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

103

catalyst bed reactor was used, with residence times of the order of 100 ms at temperatures between 900 and 1100 K. Mass spectrometry was employed to measure the product species at the outlet. Consistent with the gauze experiments, high conversions and syngas selectivity were observed, with better reactor performance observed with Rh than Pt.

2.2 Foam and Extruded Monoliths - These are both supported catalysts, with the metal impregnated in the pores of the alumina support. Foam monoliths typically consist of a-alumina in open, sponge-like structures, coated with the metal catalyst. Extruded monoliths, on the other hand, have many straight parallel channels in the basic ceramic support structure (Cordierite), coated with the metal. Thus, the differences in performance of these two systems arise mainly from differences in flow pattern and catalyst loading. High selectivities to synthesis gas have been obtained using Pt-foam monol i t h ~ ,with ~ , ~complete ~ conversion of oxygen and high conversion of methane for fuel-rich compositions between 17.4 and 29.6% CH4 : air (stoichiometric to CO and H2). The contact times were of the order of 1 ms, with autothermal temperatures between 1250 and 1400 K. Preheating the reactants was found to increase the selectivity to syngas. Maximum selectivities were obtained at an optimum inlet fuel composition of ca. 20% CH4: air. Using oxygen instead of air was found to lead to an increase in syngas selectivity and methane conversion for CH4:0 2 ratios between 0.8 and 2.0.5 Comparing Rh to Pt, Hickman and Schmidts have found that Rh-foam monoliths give comparable selectivity to CO as Pt-foam monoliths (ca. go%), but higher selectivity to hydrogen (up to 60% for Pt and up to 90% for Rh) and higher conversions of methane. Similar results were also obtained by Heitnes et al.27 for comparable contact times (of the order of ca. 10 ms). In addition, by varying the contact time (from ca. 1 to 40 ms), Heitnes et al. showed that Rh is able to maintain higher conversions and selectivities to syngas at shorter contact times, compared with Pt. At similar contact times (ca. 5-40 ms) but with lower metal loadings of Pt (0.1 wt%), poor conversion of reactants and selectivities to CO and H2 (all < 20%) were obtained.21 However, even under these seemingly adverse conditions, replacing Pt with Rh leads to selectivities of CO and H2 greater than 80Y0.~' Thus, like gauzes and sponges, supported Rh catalysts also give better selectivities to syngas than Pt. While relatively low surface areas are adequate for these reactions, both the number of gauze layers (see above) and metal loading results indicate that a minimum of catalyst sites is necessary for complete 0 2 conversion. 2.3 Fluidized Beds - Fluidized beds can give higher mass and heat transfer rates as compared to fixed catalyst beds. Therefore, in the transport-controlled regime, higher rates of heat and mass transport can be achieved leading to enhanced performance. To explore such possible advantages, Schmidt and coworkers used a-alumina spherical beads impregnated with metal (loadings of < 0.5% wt, to prevent aggl~meration).~' Using air as an oxidant, up to 90%

104

Catalysis

selectivity to CO and H2 at a methane conversion of 85% has been obtained with a Rh catalyst, at contact times of the order of 50-200 ms. Improvements in both selectivity and conversion have been obtained by using oxygen instead of air. However, the selectivity to syngas for Pt was not as high, with C02 and H20 forming in comparable amounts to CO and H2. Overall, reactor performance for fluidized beds was comparable to monolith reactors.

2.4 The Effect of the Support and Pore Diameter - A decrease in the pore size of the a-alumina foam support from 20 ppi (pores per inch) to 80 ppi leads to an increase in H2 selectivity (from 80 to 90%) and CH4 conversion (from 62 to 80%) in an autothermal monolithic reactor.28Thus, smaller pores appear to be better in terms of yield (pressure drop is obviously another consideration). Washcoating prior to the deposition of the metal leads to a small increase in the syngas yield (mainly due to an increase in methane conversion).28Furthermore, a zirconia-tetra alumina (20% ZrO,; 80% A1203)support was found to be very stable and to lead to higher conversions and selectivities than alumina.28Comparisons between metal gauze packs and foam and extruded monoliths with alumina-supported metals indicate that the foam monoliths give higher fuel conversion and selectivities to syngas over a wide range of inlet compositions and require lower metal loadings7 Overall, while the effect of the support on the reactions is minimal, its presence nevertheless adds stability to the system and would be essential for industrial applications. 3

Influence of Operating Conditions

The influence of operating parameters, such as pressure, reactor temperature, reactant pre-heat, flow velocity and diluent have been studied experimentally by various researchers. In addition, analysis of product distribution under different initial conditions through TAP experiments has provided further insight into the role of different adsorbate coverages and specific species' partial pressures in the selectivities to syngas, under conditions which have been speculated to be representative of steady-state conditions. The role of different operating conditions in reactor performance is summarized below.

3.1 Temperature - Increasing the reactor temperature, or the inlet pre-heat, increases the conversion of the fue15-19.30 and the selectivity to syngas, even when complete conversion of 0 2 is observed, for both Pt and Rh. Based on overall stoichiometry, for complete 0 2 conversion, a higher conversion of CH4 is associated with a higher selectivity to syngas. As an example, the experimental data of Mallens et al.30 for Rh sponge catalyst are shown in Figure 4.1 (a), depicting the enhanced reactor performance with increasing temperature.

3.2 Flow Velocity and Contact Time - The influence of flow velocity has been studied experimentally for the monolith reactor at relatively low contact times

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

750

800

850

900

950

1000

105

1050

Temperature (K) Figure4.1 Comparison of experimental syngas selectivity of Mallens et al.30(a) to the simulated resultsfrom the thinfilm plugflow model (b). Reasonable qualitative agreement is observed.

of ca. 1 to 75 ms by Slaa et aZ.15 and ca. 0.2 to 90 ms by Witt and Schmidt.23 The operating conditions for the two groups were similar; both used a nearstoichiometric ratio to syngas with an average reactor temperature of ca. 1200 K. However, the catalyst loadings and pore diameters were different, with Witt and Schmidt using pores of 0.5 mm diameter and ca. 5 wt% Rh,23 and Slaa et al. using pores greater than 3 pm and ca. 0.3 wt% Rh.I5 Despite the differences in reactor conditions, the qualitative behavior is similar. For example, Slaa et al. found that with decreasing contact time from ca. 75 ms to 1 ms, the conversion of 0 2 decreases from ca. 100% to 5% and the selectivity to CO drops from ca. 95% to 10%. This change in selectivities is not purely an effect of 0 2 breakthrough. In particular, while the selectivity of CO decreases monotonically over the range of contact times studied, the conversion of 0 2 remains at around 100Y0up to a contact time of ca. 30 ms, below which it decreases sharply to only ca. 5%. In this regime of incomplete 0 2 conversion,

Catalysis

106

GHSV (hrl) 1 ' " ' ~ " ' " " " " " ' ~

Figure 4.2 Conversion of CH4 and 0 2 and selectivities to CO, C02, H2 ani H20 as a function of $ow velocity in Rh-coated monolith experiments (from Witt and Schmidt23) (a) and a thin _film plug $ow reactor model (b). Reasonable qualitative agreement is obtained

H2 selectivity exhibits a slight minimum followed by a maximum as contact time decreases, with the opposite trends seen for H20 and CO2. Based on the results of Slaa et al., it seems that for partial oxidation of CH4 in monolith reactors, optimal selectivities to H2 and CO are seen at high enough contact times when O2conversion is complete. The experimental results of Witt and Schmidt are in qualitative agreement with the results of Slaa et al. at the high contact times. The data of Witt and Schmidt for one set of experiments are summarized in Figure 4.2(a). At relatively high contact times of ca. 2 ms or more, 100% of O2 conversion was observed, with high selectivities to CO (ca. 95%) and H2 (> 80%). As the contact time decreases to 0.2 ms, the conversion of O2 decreases to less than 80%' and C02 and H20 are favored over CO and H2. For small reactor diameters (< 11 mm), Witt and Schmidt's selectivities vary monotonically with

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

107

flow velocity. However, a slight maximum and minimum in Hz and H20 selectivities are seen (respectively) as contact time increases for the larger reactor diameter of 18 mm. These maximum and minimum occurred at relatively high contact times with complete 0 2 conversion. An interesting outcome of this comparison is that similar performance can be achieved with different conditions and reactors. This in turn indicates that reactor optimization, including reactor configuration and operating conditions, should be an integral part of scale-up. 3.3 Pressure - The total pressure affects the system in various ways. First, the radial mass transfer coefficient is expected to decrease, primarily because species’ diffusivities are inversely proportional to pressure (other thermophysical and transport properties such as kinematic viscosity also change). Second, surface reaction rates are expected to change, due to a change in adsorption rates (typically adsorption of the ‘surface’ limiting species is rate determining under these conditions). As a result, the coverages of species on the surface are altered, leading to a change in selectivities. Plug flow reactor simulations at constant temperature showed a significant drop in selectivities to syngas with increasing pressure, caused mainly by the latter effect.32Third, for a fixed mass flow rate, pressure is expected to influence the contact time, since gas velocities are inversely proportional to pressure. Fourth, the potential for the onset of gas-phase reactions is increased, since higher pressures promote gas-phase reactions of methane. * 3*33 As a result, gas-phase ignition temperatures decrease and the upper flammability limit (of interest to partial oxidation of methane) increases as the pressure rises. In particular, experimental data for the homogeneous combustion of natural gas-air mixtures indicate that the upper flammability limit increases from a composition of ca. 14% natural gas in air at atmospheric pressure to ca. 60% at approximately 70 atm.14 Such effects are typically highly undesirable because of safety and the potential for decreased selectivity to ~ y n g a s Even . ~ ~ though this behavior is intuitive, the gas-phase ignition temperature of most fuels does not decrease monotonically with increasing pressure over a wide range.14 In that regard, methane is the exception rather than the rule. Finally, consideration should also be given to the heat management in a reactor. Since pressure affects the overall flow velocity (and thus, heat transfer coefficients), reaction rates and gas-phase heat capacities, autothermal operation temperature is also expected to vary with pressure. Despite the importance of pressure on reactor performance, limited information has been published, primarily due to safety and high costs associated with high pressure experiments. Dietz and Schmidt have studied the influence of pressure at fixed mass flow rates, over the range of 1.4-5.5 atm. Overall, a weak dependence of CH4 conversion and CO and H2 selectivities has been observed, with CH4 conversion and CO selectivity decreasing slightly with increasing pressure.24Further experimental studies on the effect of pressure are needed since this is a crucial parameter for potential commercialization.

108

Catalysis

3.4 Dilution - Dilution of reactants is another parameter that influences reactor performance. Simulations and experiments of H2 oxidation over Pt35336 have shown that dilution can be used effectively to control autothermal reactor temperature (for both safety and product formation) at fixed fuel : oxidant ratios. In general, dilution is expected to influence mass transfer coefficients through a change in species’ diffusivities, to lower the autothermal reactor temperature, and to decrease the flammability regime. Whereas dilution is used in most laboratory experiments to approach isothermal operation and assure safety, pure 0 2 appears to give higher selectivities to ~ y n g a s . ~ The influence of N2 dilution on product distribution has been studied experimentally for autothermal monolith reactors by Witt and Schmidt.23At a constant space velocity and over a N2 dilution range of 4-45%, they have found that the exit temperature of the autothermal reactor varied from ca. 1750 to ca. 1600 K as N2 dilution increased. Conversion of both O2 and CH4 decreased with increasing dilution. Furthermore, the selectivity of H20 increased (from around 33 to 57%) at the expense of H2. In contrast, C02 and CO increased slightly with N2 dilution (a surprising result). Since selectivities and conversions depend strongly upon temperature, it is difficult to differentiate the influence of temperature from dilution alone. However, several features indicate that temperature cannot be the sole cause for the change in reactor performance. First, while the difference in operating temperature was relatively small, only ca. 150 K for a substantial dilution change, the change in the selectivities to H20 and H2 was quite dramatic. Second, while the temperature of the reactor decreased, the selectivity of CO increased with increasing dilution, an opposite trend to the experiments involving reactor temperature variation. These results indicate that dilution could be a viable strategy to control the relative selectivities of H2 and H20. 3.5 Inlet Fuel Composition - For fuel-rich mixtures typically of interest for syngas production, the selectivity to syngas goes through a maximum at a CH4 :O2 ratio of ca. 1.1 for Pt and ca. 1.4 for Rh. As the inlet fuel composition is increased (for reactant inlet at room temperature), the conversion of the fuel decreases mon~tonically.~ Since 0 2 is the limiting reactant, complete 0 2 conversion is generally observed given sufficient contact time. An interesting feature is that the optimal selectivity to CO and H2 does not necessarily occur at the expected 2 : 1 CH4: 0 2 ratio? Furthermore, the maxima in CO and H2 selectivities shift to higher CH4 :0 2 ratios as the inlet reactant temperature increases. Several features contribute to these phenomena. First, gas-phase reaction stoichiometry is not directly applicable to catalytic systems. For reactors in which both axial and radial gradients in species concentrations exist (ie. the monolith reactor), the surface stoichiometric point37(i.e. the relative reactant concentrations just above the catalyst) dictates the overall stoichiometry of the surface chemistry. The latter is in turn determined by the relative diffusivities of the reactants and the conversion near the surface up to each longitudinal position. Second, the overall stoichiometry depends also on the reactor temperature, since the latter affects the catalytic

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

109

adsorption, desorption and surface reactions rates which alter gaseous species’ concentrations. Consequently, changing the inlet temperature of the reactants or the CH4 :0 2 ratio affects the autothermal temperature, shifting the location of the maxima in CO and H2 selectivities. For fuel-rich mixtures, the autothermal operation temperature decreases with increasing fuel composition. Thus, while increasing the inlet fuel composition favors the formation of syngas (because it leads to a more oxygen-starved surface), the lower temperatures favor complete oxidation products. This competition in mechanisms results in a maximum in syngas selectivity that differs from the expected CH4 :0 2 ratio of 2. Finally, the extent of the carbon-forming reactions on the catalyst, and the temperature range over which these reactions are dominant, is expected to vary with the inlet fuel composition (as discussed in Section 5). This will in turn influence the selectivities obtained, since the dominant reaction paths may change in the presence of surface carbon.

3.6 Influence of Different Surface Coverages - Because of the differences in surface concentrations of reactants in the longitudinal direction with respect to the surface stoichiometric point, it is expected that the coverage of adsorbed oxygen is a key in determining the product selectivities. This hypothesis is further supported by the transient TAP experiments of Mallens et al.30 and Buyevskaya et al.,38who pretreated a Rh catalyst surface with a CH4,02, or a reducing atmosphere, and observed the effect of different surface coverages on the transient product formation at catalyst temperatures of around 1000 K. For CH4 flowing over an initially 02-treated catalyst, Buyevskaya et al. observed first 100Y0selectivity to C02 followed by a shift in the main carbon product from C02 to CO, indicating that an oxygen-rich catalyst surface leads mainly to C02 formation, whereas an oxygen-lean surface leads mainly to CO formation. Mallens et al. obtained the same qualitative behavior using pulsed reactant experiments. Using gas chromatography, a small C02 peak was observed prior to catalytic ignition of CH4-air mixtures over a Pt consistent with model predictions,M indicating complete combustion products under high surface oxygen coverage and low temperatures.40Consistent with 0 2 pretreatment experiments, flow of CH4-02 over a reduced surface led to preferential production of CO over C02 by 3 : 1. In addition, Buyevskaya et al. have performed experiments with flows of O2 or C02 over a carbon-containing catalyst (CH4-pretreated). They found that under their operating conditions, 0 2 is unreactive with surface carbon species, whereas C02 was converted to CO. This result indicates that, for Rh, carbon acts as a poison for 0 2 adsorption, but it is reactive with C02 through the reverse Boudouard reaction. In summary, both the steady state and the TAP experiments indicate that coverage of O* on the catalyst surface is a key factor in determining partial versus complete oxidation of CH4. The role of O* is further discussed below.

110

Catalysis

Table 4.3 Simplified mechanisms of surface reactions in the partial oxidation of methane Reaction

Activation energy (kcal mol- I )

CH4+5* + C * + 4 H * 0 2 + 2* --* 2 0 * c * + o* + c o * + * c o * + co + * c o * + o* + c02 + 2* 2H* + H2 + 2* 2H* + O* + H20 + 3* 2 0 * -+0 2 + 2*

0 29 32 13 21 24

4

Reaction no.

5

85

Analysis of the Effect of Operating Conditions

4.1 Model - In order to perform a more detailed analysis of the coupling between the mass transport and surface reactions as a function of reactor operating parameters, a steady-state thin film plug flow reactor (representing one channel of an extruded monolith) was modeled. The governing equations considered are i = 1,...,mg

k,j((

-

4) = -<(8,cs),

i = 1,...,mg

j= I

4

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4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

111

Table 4.4 Activation energies (kcal mol-') of suflace reactions for different catalysts calculated through bond-order conservation ~

Reaction

----

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0 2

CO formation c*+o* co*+* co* c o + * C02 formation c o * +o* c*+co2*

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H20

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~

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Ni

51.0

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15.2 0.0

20.3 0.0 5.3

24.1 5.5 19.9

27.9 10.0 28.3

21.0

21.0

23.3

reaction (Rl), the proposed value by Hickman and Schmidt4 of 5 kcal mol-' is used, whereas for the global reaction of H 2 0 formation [reaction (R7)],the highest activation energy in the sequence of H 2 0 formation steps, namely that of the H*+ O* + OH* reaction (see Table 4.4) is used. The pre-exponential factors of all reactions are taken to be 10" s-', with 0.1 and 0.001 as the sticking coefficients of 0 2 and CH4, respectively. It has recently been shown that a skeleton mechanism with BOC-computed activation energies and rough estimates of reaction pre-exponentials is an adequate starting point for constructing surface reaction mechani~rns.~' Numerically, a marching scheme is employed where an integration of the bulk concentrations along the length of the reactor ( z ) is done using an Euler scheme, followed by Newton's method to compute the surface coverages and the gas-phase concentrations just above the surface. For the value of kc in laminar flow, the Sherwood number Sh can be approximated as42

Sh = 1.62(Re)1/3(Sc)'/3

(4.5)

where Re is the Reynolds number and Sc is the Schmidt number. In dimensional form, this expression can be simplified to give an expression for the radial mass transfer coefficient as a function of the flow velocity, tube diameter d, and species diffusivity Di 213

kci= 1.62~"~($)

(4.6)

112

Catalysis

The role of the most important parameters (contact time and temperature) in reactor performance is discussed next. 4.2 Flow Velocity - An example of modeling results depicting the effect of flow velocity on reactor performance is shown in Figure 4.2(b). The range of contact times, channel diameter, reactor length (1 cm) and other operating conditions were chosen in accordance to Witt and Schmidt’s23experiments [Figure 4.2(a)]. Experimentally, the longitudinal temperature was fairly constant at ca. 1000 K for the range of flow velocities considered. We have therefore simulated an isothermal reactor at 1000 K. Simulation results are in qualitative agreement with experiments (no refinement of reaction pre-exponential factors was attempted). In particular, important features such as the transition from complete 0 2 conversion to 0 2 breakthrough, the decrease in CH4 conversion, and the increase in H20 and C02 selectivities at the expense of CO and H2 selectivities as a function of increasing flow velocity are all well described by the model. Apart from the direct partial oxidation path included in the analysis here, the indirect path to syngas (by reforming) may also affect the syngas selectivities. This issue is discussed below in Section 6. To delineate the effect of flow velocity, Figure 4.3(a) plots the surface coverages of dominant species versus residence time (to scale the effect of higher velocities) for different inlet flow velocities and Figure 4.3(b) shows the corresponding concentration of O2 in the bulk and adjacent to the surface. The results show that the surface near the entrance of the reactor is dominated by 0*,because the sticking coefficient of 0 2 is higher than that of CH4 and the adsorption of CH4 is activated. Along the length of the reactor, the coverage of O* decreases due to surface reactions, causing depletion of 0 2 in the gasphase, eventually leading to a C*-covered surface. The high coverage of O* (owing to its high activation energy for desorption) indicates that the adsorption of CH4 and the desorption of O* play significant roles. This point is further discussed in Section 6. This example is an illustration of the fact that the inlet composition alone is not sufficient to determine the limiting surface reactant (see also Section 3.5 on surface stoichiometric point). Both simulations and experiments indicate two flow velocity regimes, namely a low velocity one with complete conversion of 0 2 and a high velocity one with an incomplete conversion of 0 2 . The different selectivities in the two regimes can be explained by 0 2 breakthrough only in part (even if the reactor is long enough, significant drop in syngas selectivity is found with increasing flow rate). At higher flow velocities, O* dominates over the entire length of the reactor, reducing adsorption of CH4. Such an abundance of O* favors complete combustion products. In essence, higher velocities have the effect of preferentially increasing 0 2 near the surface (by increasing the radial mass transfer coefficients), increasing the rates for complete oxidation and increasing the length needed for complete 0 2 conversion. This result differs from the conventional understanding of reactor behavior, where the system is generally thought to be mass transport-controlled following catalytic ignition, and an increase in species mass transport leads to an overall increase in

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

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Q)

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reactivity of the system. Obviously, this general concept is not true for negative order catalytic kinetics observed in this case. Furthermore, our analysis indicates that increasing the length of the reactor should only partially overcome the decreasing syngas selectivities with increasing flow velocity. 4.3 Temperature - Figure 4.l(b) shows the effect of temperature on selectivities using the thin film plug flow model. Qualitatively, the increase in H2 and CO selectivities seen experimentally [Figure 4.1(a)] with increasing temperature is well described by the model. However, quantitatively, the selectivity to Hz, along with its relatively strong dependence on temperature are not as well described. Sensitivity analysis on the effect of temperature indicates that the enhanced performance with increasing temperature is caused by an increase in

114

Catalysis

CH4 adsorption due to its activated nature, an increase in the desorption of O* that blocks sites and inhibits CH4 adsorption, and a relative increase in the desorption of CO* compared to CO* oxidation to COZ. Finally, H* also has two competing paths, desorption (R6) and oxidation (R7). Note that reaction R7 has a slightly higher activation energy than R6, and increasing temperature should actually decrease H2 selectivity. The relative insensitivity of H2 selectivity with temperature shown in Figure 4.l(b) is then due to the competition of H20 formation with the enhanced CH4 adsorption that results in more H*. 5

Bifurcation Behavior of Methaneoxygen Mixtures Near Catalysts

The important bifurcation characteristics of catalytic oxidation systems consist of ignitions, extinctions and autothermal points. More complex dynamics such as oscillations, chaos and pattern formation on the catalyst have been the subject of extensive research43and will not be covered here. Identifying and understanding these features as a function of reactor operating conditions are crucial, because they define the windows of operation for optimal performance and safety.44

5.1 Influence of Fuel Composition - An experimental bifurcation diagram of CH4-air oxidation over a Pt foil is shown in Figure 4.4,39 depicting catalytic ignition, extinction and homogeneous ignition temperatures as a function of normalized CH4 composition. Catalytic ignition (open circles) separates the unreactive from the reactive temperature space, and represents reactor startup, a necessary first step in either complete or partial oxidation. For example, in the study of Heitnes Hofstad et al., no catalyst activity was observed until a temperature of ca. 1100 K.l8*I9 Similar behavior was also reported by Dissanayake et aZ.45 Catalytic ignition has been extensively studied for CH4-air mixture^^^.^' as well as diluted CH4-O2 m i ~ t u r e s ~over ~ 9 ~Pt~ wires and foils. No such information is available yet for Rh catalysts. In general, the ignition temperature is seen to decrease with increasing inlet methane composition, i.e. the fuel has a promoting influence on ignition (this is generally true for paraffins over most of the compositional regimes0). As a result, fuel-rich mixtures, of interest for partial oxidation, are more easily ignitable. The relation of bifurcation data obtained using wires and foils to practical reactor configurations is important. The ignition temperatures are seemingly unaffected by differences in reactor geometry, dilution level and flow rate used in various experiment^.^' Such an insensitivity indicates that ignition is primarily governed by the kinetics (site autocatalysis) of the oxidation process rather than, for example, reaction ex~thermicity~~. However, we should note that catalyst pretreatment and site density (more of concern in the experiments of partial oxidation where catalyst loading varies between groups) can affect ignition temperature for some fuels (for example, see H2’*) and preheating could lower ignition temperature, especially when reaction exothermicity plays a role (this is not the case for

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CH4). Furthermore, it may be concluded that the competitive adsorption of the reactants on the catalyst surface is a prerequisite for ignition, as also indicated by sensitivity analysisw and the lack of hysteresis for dilute mixtures when a non-competitive surface reaction mechanism is e m p l ~ y e d . ~Such ' a conclusion has also been reached with a simplified chemistry rnodeLs0 The terminal fuel-lean limit of ignition shifts to higher methane compositions as the dilution level increase^.^ This feature may be attributed, however, to extinction, which unlike ignition is thermally affected. Once catalytically ignited, decreasing the catalyst temperature eventually leads to catalytic extinction, often associated with reactor shutdown. For CH4-air mixtures, the extinction temperature is higher than the ignition temperature, and has a maximum near the stoichiometric composition to C02 and H20. Realistic reactor startup and shutdown is expected to be more complex than the behavior shown in Figure 4.4, which applies only to stationary bifurcations. Startups and shutdowns involving slow introduction of one reactant can lead either to coking or transient crossing of a flammability limit with undesirable safety consequences. For these reasons, in the experiments of Schmidt and co-workers, ignition of ammonia was first employed to avoid possible flammable tartu up.^ Under some conditions, a catalytically ignited system is capable of sus-

116

Catalysis

taining combustion without any external heat supply, an operation known as autothermal. Such operation is desirable, because it minimizes both operating cost and capital investment necessary for reactor heat exchange. The autothermal temperature of H2-02 mixtures has been found to be sensitive to reactant flow velocity, dilution and heat losses.36 For these reasons, autothermal behavior was not observed for the experimental results shown in Figure 4.4. However, under different operating conditions over a Pt foil, autothermal points have been detected in a small compositional regime between about 10 and 18% CH,-air, for a contact time of the order of ca. 1 s.46 In contrast, monolithic reactors, with millisecond contact times, exhibit autothermal operation for significantly more fuel-rich mixture^.^ This is not surprising as heat losses are lower for a monolith, and fast flows lead to large heat and mass transfer coefficients expanding autothermal behavior that is primarily transport-controlled (except near the end composition^^^). Figure 4.4 shows that for some compositions, and in the temperature range of ca. 1250- 1450 K, homogeneous ignition (open triangles) can occur. Ignition of gas-phase reactions can lead to an undesirable increase in temperature (leading to catalyst evaporation, sintering or melting) and poor selectivities to syngas. Upon homogeneous ignition, we have experimentally observed both flame propagation and flashback, indicating that safety is also a major concern. Figure 4.4 indicates that the temperature window for partial oxidation is confined between catalytic extinction (ca. 900 K) and homogeneous ignition (ca. 1400 K). For sufficiently fuel-rich conditions, although the experimental data indicate no homogeneous ignition in terms of hysteresis or a visible flame, the onset of gas-phase chemistry still occurs, potentially leading to lower selectivities to syngas. The role of gas-phase chemistry is further discussed in Section 5.4. Whereas catalytic ignitions and extinctions and coking of the catalyst under fuel-rich conditions have been reported in a variety of reactor geometries, we have recently observed a new feature in CH4 oxidation (which was apparently not reported before), namely a second hysteresis associated with the ignition of coke or carbon. Figure 4.5(a) depicts an experimental one-parameter bifurcation diagram showing the change in catalyst (Pt foil) temperature as a function of current (or heat) supplied, for a relatively fuel-rich reactant mixture of 19% CH4 in air. At relatively low temperatures of ca. 800 K, catalytic ignition and extinction are seen. Further increase in power to the foil leads to visible carbon formation on Pt. However, at ca. 1200 K, there is a second ignition, leading to higher temperatures and carbon burn-off. Hysteresis is observed, as a second extinction occurs at a different current input. This carbon ignition occurs mainly in the fuel-rich regime and exhibits a maximum with varying inlet composition as shown in Figure 4.5(b). Analysis indicates that this second . hysteresis occurs primarily from the change in the heat loss of the system, as radiative emissivity is high for a carbon-covered Pt surface as opposed to a bare Pt surface. Preliminary gas chromatography data39indicate that when the catalyst is coked, COz is the primary carbon-containing product and is reduced with increasing temperature to form CO (possibly due to the reversed

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Boudouard reaction). On the other hand, upon carbon ignition, CO is the primary carbon containing product. This change in product selectivities indicates that more detailed carbon chemistry needs to be considered in reaction mechanisms to capture the complete behavior of CH4 oxidation. In order to construct a full bifurcation diagram through numerical simulations, steady-state equations for continuity, mass, momentum and energy have to be solved. Detailed reaction mechanisms for the description of both the gasphase and the surface chemistry, and multicomponent transport models are essential. Furthermore, the presence of turning points and multiple steadystates precludes the use of regular numerical solvers. Mathematical techniques

118

Catalysis

to compute bifurcations and pass around turning points have to be incorporated. In a series of studies, a dynamically adaptive, multiple-weights arclength algorithm has successfully been applied for this purpose, and bifurcation diagrams for catalytic and non-catalytic systems have been obtained for a variety of fuels.37~40~53~54 Such algorithms overcome limitations of inaccuracy, the possibility of overlooking additional multiple solutions, and computational intensity of time integration solvers.5 5 ~ 5 6 Among the bifurcation features in the catalytic oxidation of methane, catalyst ignition temperatures have been well captured by numerical calculations in literature, and detailed analysis was done to determine the important step^.^^,^^ Competitive adsorption of methane and oxygen, and oxygen desorption were found to be the only important steps. Simple analytical criteria, that can predict the experimental data well have also been d e ~ e l o p e d . On ~~.~~ the other hand, little is known about catalyst extinction and autothermal behavior. Recent analysis shows that the chemistry of CO may play a crucial role in determining extinction temperatures for fuel-rich CH4-air mixtures5' Additional features, such as 0*-assisted H-abstraction from methane and alternative routes to CO formation (other than the direct reaction of C* and O*),have to be further examined in this context. 5.2 Influence of Pressure and Flow Velocity - The effect of pressure (ca. 0.1-2.5 atm) on catalytic ignition of CH4-air mixtures over a Pt foil has recently been studied. The results indicate that the catalytic ignition temperature is relatively independent of pressure, except at low pressures near the limit of hysteresis (cusp point) where the ignition temperature increases with decreasing pressure. 39 If these results can be extrapolated to sufficiently high pressures, then gas-phase ignition temperatures can be comparable to catalytic ignition. Furthermore, the autothermal regime expands with increasing pressure. Experiments at even higher pressures are desirable. The influence of flow velocity on catalytic bifurcation points has also been studied experimentally in the same reactor geometry. Catalytic ignition was found to be independent of flow velocity, for approximately an order of magnitude change in flow velocity. On the other hand, the catalytic extinction temperature increased with increasing flow velocity, by up to ca. 150 K. This latter result indicates that an increase in pressure decreases the temperature window for partial oxidation.

5.3 The Inhibiting Role of the Catalyst in Gas-phase Ignition - Catalysts are known to inhibit gas-phase ignition. This feature has been observed as early as in the 1920s for natural gas59 and verified more recently in the oxidation of methane48 and hydrogen60 over various catalysts, with platinum having the most inhibiting effect (there are no available data for Rh). Some of the possible causes of this inhibition are the depletion of the limiting reactant, catalytic formation of flame retardant species, and the recombination of chemically active gaseous radicals on the catalyst to form stable molecules.54 For example, in analysis of hydrogen oxidation near platinum surfaces using

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

119

detailed gas- and surface-phase chemistries and multicomponent transport, it has been concluded that the inhibition of the gas-phase ignition is mainly due to the catalytic production of H20 (which in turn increases the rate of the gasphase ‘chain termination’ reaction H + 0 2 + M + H 2 0 + M) and the depletion of the limiting r e a ~ t a n t . ’Such ~ steam-induced inhibition is not expected to play a significant role in methane ignition, when selectivities to syngas are high and temperatures are above 1000 K, due to the high speed of the competitive ‘chain branching’ reaction H + 0 2 + OH + H at higher temperatures. Furthermore, dilution and fast flows are expected to diminish this phenomenon. Simulation results given in Figure 4.6(a) using an available literature reaction mechanism show that the homogeneous ignition of CH4 and the fuel-rich ignition limit (the point where hysteresis ceases) are inhibited by the presence of a catalyst (Pt). These features are in qualitative agreement with experimental observations [Figure 4.6(b)].47,48-59 Note that such inhibition is less important for mixtures sufficiently fuel-rich than the stoichiometric. The similarity of results between H2 and CH4 systems and the generality of mechanisms indicate that this catalyst-induced retardation of the onset of gaseous chemistry is possibly a generic feature of fuel combustion. Gas-phase ignition can also be promoted by catalyst radical formation and desorption to the gas-phase, as both experiments and simulations i n d i ~ a t e . ’ ~However, ?~’ the latter phenomenon appears to be dominated by the inhibition mechanisms outlined here. From a practical point of view, the catalyst expands the temperature window by ca. 200-400 K before the gas-phase chemistry becomes significant. Furthermore, calculations using homogeneous chemistry alone (a much simpler task as heterogeneous mechanisms are still under development) could be used as they provide a conservative limit for gas-phase ignition temperature.

5.4 Importance of Gas-phase Reactions - The contribution of gas-phase reactions has been one of the most debated issues in the partial oxidation of CH4 in short contact time microreactors. Ignition of the gas-phase chemistry is typically undesirable for the synthesis of useful chemicals (the oxidative coupling of methane over oxides is a noticeable exception). At high temperatures, gas-phase methane oxidation is unselective, producing C2 species such as C2H4 and C2H6, various oxygenates, and complete combustion products.34 The small fraction of these species formed in the experiments of Schmidt and co-workers support the view that gas-phase reactions are not significant under typical operating condition^.^ Moreover, simulations have been performed32~40~5’~55,56~62-64 with detailed gas-phase chemistry models, with and without surface chemistry. At atmospheric pressure, all numerical studies have found insignificant contribution of gas-phase reactions. In the study of Veser and F r a ~ h a m m e r the , ~ ~induction time for ignition of the gas-phase reactions was found to be at least two orders of magnitude greater than that for ignition of the catalytic reactions. This indicates that the reactants are depleted by surface chemistry prior to gas-phase ignition (a form of catalyst inhibition). These results should be viewed with some caution because the mechanism does

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I-*--- -.-

U

1200

0.20

*_.---.

Coward and Guest, 1927 Ni Surface

0.30 0.40 CH4/ ( CH4+02)

0.50

Figure 4.6 Model-predicted CH4 homogeneous ignition temperature as a function of normalized composition for a stagnation point flow reactor at atmospheric pressure and an inverse strain rate of 5 ms. Results for only the homogeneous chemistry (inert surface, dashed line) and the homogeneous-heterogeneous chemistries (Pt surface, solid line) are shown (a). For the gas-phase chemistry, the Gas Research Institute M e ~ h - 1 . 2 ‘is~ used, ~ whereas for the Pt catalytic chemistry, the mechanism of Aghalayam et aL5’ is used The experimental data of Coward and C u e d 9 for a relatively inert Ni surface and the data of Gr&$fn and Pfefferle48 for an active Pt surface are also shown (b). The Pt catalyst inhibits both the homogeneous ignition temperature and the fuel-rich ignition limit.

not predict a proper catalytic ignition (similar caution is applicable to other modeling studies due to various limitations in reaction mechanisms discussed below). However, given that the catalytic ignition temperature is much lower than the homogeneous ignition temperature (at these conditions), the results should be qualitatively correct. Furthermore, the authors note that increasing pressure increases gas-phase reaction rates more rapidly than surface reaction rates.63 This is evident as gas-phase ignition of CH4 decreases monotonically

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with increasing pressure, observed both through simulations by Vlachos and c o - ~ o r k e r s ’in~ -a~perfectly ~ mixed reactor and experiments by Law and coworkers for a counterflowing diffusion flame.66Thus, at conditions of industrial importance (high pressures), gas-phase reactions may play a role in catalytic partial oxidation, and it is necessary to explore this factor in more detail. The importance of gas-phase chemistry in monolith reactors has been studied by Vlachos and c o - ~ o r k e r sby ~ ~varying the pressure and pore diameter. Their results indicate that for large pore diameters (ca. 1 cm), gasphase reactions become competitive to catalytic reactions under the typical syngas operating conditions shown in Table 4.2. As the pressure increases, gasphase reactions dominate, leading to lower syngas selectivities at smaller pore diameters. Simulations indicate that short contact times have interesting implications regarding safety. A decrease in the contact time could enhance catalytic autothermal operation, due to increased heat and mass transfer, while increasing gas-phase ignition temperature more than the catalytic ignition temperature, and reducing gas-phase flammability (a more safe operation).

6

Direct versus Indirect Path to Syngas Formation

In one of the early quantitative experiments of CH4 partial oxidation to syngas, Prettre et aZ.67reported a temperature profile along the reactor length using a supported Ni catalyst which showed an initial sharp increase in temperature near the inlet, followed by a decrease downstream. The authors speculated that initially a strong exothermic reaction occurred, followed by either less exothermic or even endothermic reactions. As shown in the introduction, complete oxidation of CH4 is exothermic, whereas the steamreforming, water-gas shift, and dry-reforming reactions are all endothermic. Thus, the concept of ‘indirect’ syngas formation was first introduced, where the primary products, C 0 2 and H 2 0 , may then be converted to CO and HZ.In the ‘direct’ partial oxidation scheme, on the other hand, the primary products, CO and H2, are directly obtained from the reaction of methane and oxygen, and may further be oxidized to CO2 and H20 by unreacted oxygen. A similar reactor temperature profile was also observed for Pt catalysts.27 Whereas the indirect path to syngas was mainly advocated in earlier times, the high conversions and selectivities in the Schmidt and c o - ~ o r k e r s short’~ contact time gauze and monolithic reactors that differed from equilibrium predictions suggested that the direct path is dominant. Following this latter work, significant debate has appeared in the literature as to the nature of the dominant path leading to syngas. Foulds and Lapszewicz’ and Tsang et nl.3 have reviewed earlier findings on various catalyst supports and operating conditions and concluded that neither the direct nor the indirect partial oxidation mechanisms can explain all the existing experimental data. They proposed that both paths are possible, and the relative importance of each is determined by the catalyst, the support, and

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Catalysis

the operating conditions. More recent experiments and simulations outlined below indicate that the direct path is more relevant for short contact time reactors, as the reforming reactions are typically much slower than the oxidation reaction^.^ In experiments with Pt-Rh gauzes with very short contact times (0.2 ms), Heitnes Hofstad et aZ.19 have used feeds consisting of CH4 and H20, and H2 and C02 at a temperature of 1300 K. No conversion to CO and H2 was found in either case, whereas when CH4 and 0 2 were used, high conversions of methane and selectivities to syngas were obtained (see Table 4.2), indicating that the steam-reforming and the water-gas shift reactions are not dominant in syngas production, Mallens et aL30 have observed in their experiment involving the simultaneous pulsing of CH4 and O2 onto Pt or Rh sponge catalysts that CO and H2 emerge at the outlet prior to C02 and H20. This indicates that syngas is the primary product. The contact time in that study was relatively high (ca. 100 ms). Similar conclusions were reached by Buyevskaya et aZ.38 using various feedstocks and Rh treatments (see Section 3.6). Another interpretation of the non-uniformities in temperature profiles is possible, where the initial sharp rise in the temperature may be attributed to fast exothermic reactions (total and partial oxidation), whereas the subsequent drop in the temperature of the reactor may be due to heat losses. In twodimensional adiabatic monolith simulations, a sharp increase (overshoot) in temperature is seen at the location where catalytic ignition occurs, as catalyst sites are generated by desorption of one of the adsorbates. Such a scenario occurs when adsorbate coverages are high (catalyst blocking by O* in the case of CH4) and the reactions are fast compared to adsorption of the limiting surface reactant (CH4 as shown in Figure 4.3). Under such conditions, the reaction zone is confined to very near the ignition location, where high coverages of O* occur and the spike in temperature is associated with a transition from the pre-ignition zone to partially combusted CH4.32Furthermore, equilibrium calculations of methane conversion and syngas selectivities, based on the total oxidation, steam-reforming, and water-gas shift reactions, have been done27 using measured monolith outlet temperatures from experiments. At low contact times (5 ms), equilibrium-predicted compositions compared poorly with the experimental data, although better agreement was obtained at higher contact times (20 ms), consistent with the hypothesis that endothermic reforming reactions are slow. Among Pt, Rh and Ni, several differences are observed. For example, Pt exhibits lower CH4 conversion and selectivity to syngas than Rh at comparable reactor temperatures, both for monolith30 and wire gauze reactor^.'^-'^ On the other hand, Ni exhibits comparable selectivities and conversion to Rh. However, Ni catalyst reactors are characterized by high temperature peaks near the inlet67-68and the formation of coke downstream at reactor temperatures between ca. 720 and 1170 K,68which apparently does not seem to hinder CH4 conversion or syngas formation. Some of the differences seen in experiments among these catalysts can be explained based on energetics of detailed surface reaction paths without the

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use of global steps listed earlier or a change in elementary chemistry. For this, the BOC theory of S h u s t o r o ~ i c h has ~ ~ -been ~ ~ used to compute activation energies of potentially important steps of CH4 oxidation. Similar activation energies for Ni and Pt have previously been computed by Shustor~vich,~' and for Ni, Pt and Rh clusters by Liao and Z h a r ~ g While . ~ ~ quantitatively these computed activation energies may sometimes differ from experimentally measured values, nonetheless they should depict the correct qualitative behavior, by providing insight into how activation energies of particular pathways change among catalysts. In addition to the activation energies, the surface coverage also has a strong influence on the reaction paths, and will be elaborated below. The values shown in Table 4.4 have been computed based on heats of feature is that chemisorption reported in the l i t e r a t ~ r e . ~ ? ~One ~ ~striking ~~ most of the activation energies predicted are below ca. 30 kcal mol- l, with the exception of O* desorption. At typical reactor operating temperatures of greater than 1000 K, surface reactions should be quite fast in comparison to oxygen desorption. In particular, the data show that the activation energy for O* desorption increases in the order Pt c Rh c Ni. Thus, under identical reactor conditions of catalyst temperature and partial pressures of reactants, one would expect higher surface coverages of O* on Ni than on Pt or Rh, especially near the inlet of the reactor where little oxygen has been consumed (assuming that the relative fluxes of CH4 to O2 are similar for the three catalysts). Furthermore, while the formation of CO* on Ni from C* + O* has a relatively high activation energy, the subsequent oxidation step of CO* + O* -+ C02* has a much smaller activation energy than the desorption of CO*. The combination of the low activation energy pathway to C02* formation and the abundance of O* suggest that Ni initially oxidizes CH4, mainly to C02, and strongly consumes oxygen, leading to highly fuel-rich conditions downstream that give rise to coke formation. Hence, the high selectivity of CO on Ni is mainly affected by secondary reactions involving COZ, coke and H20, supporting more the 'indirect' pathway. This explanation is consistent with the experimental r e s ~ l t s . ~ ~ * ~ * The higher selectivity to CO on Rh compared to Pt can also be explained using the computed BOC activation energies. On Pt, C02* formation from CO* + O* has a very low activation energy compared to O* or CO* desorption, implying that with an appreciable coverage of O*, C02 formation is relatively easy. In contrast, on Rh, the oxidation of CO* to C02* has a much higher activation energy that is comparable to the desorption of CO*. As a result, the CO selectivity is higher on Rh. Regarding the selectivity of H20 versus H2 on Pt and Rh, a main difference is in the activation energy of H 2 0 formation via the 973774

20H* -P H20* + O* and H* + OH*

+ H20*

(4.7)

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Catalysis

paths, with activation energies for the associative desorption of H being comparable. The work of Vlachos and co-workers4 for Pt-catalysed oxidation of H2-02 mixtures has shown that the H 2 0 formation path depends strongly on the surface adsorbate coverages. Under high H* coverages, reaction (4.8) is the main path, whereas for surfaces deficient in H*, reaction (4.7) is the preferred path. Under typical operating conditions for CH4 partial oxidation, only OH* has been observed experimentally by Walter et aZ.,75but the H* coverage has not yet been established. One may expect that under 0*-rich surface conditions, path (4.7) dominates, whereas under fuel-rich surface conditions, path (4.8) is important, Irrespective of the specific path, the predicted BOC activation energies indicate that Rh has a higher activation energy than Pt for both H20 formation pathways, with reaction (4.7) exhibiting a significantly higher activation energy of 19.9 kcal mol-', compared to 5.3 kcal mol-* for Pt. This would at least qualitatively indicate that H 2 0 formation rates are lower on Rh than Pt, leading to higher H2 selectivities, in agreement with an early suggestion by Schmidt and c o - ~ o r k e r s . ~ The above conclusions have been reached based solely on the energetics of the reactions, without any consideration being given to the reaction preexponential factors. Furthermore, the activation energies shown in Table 4.4 do not include interactions between adatoms, which depend strongly upon surface adsorbate coverages and operating condition^.^' Finally, the adsorbate coverages dictate surface reaction paths (there is an inherent coupling between gaseous concentrations and surface coverages, and the distinction has only a mechanistic merit). Besides the effect of O* coverage and possible various catalyst states (see Sections 4.2 and 7.2.2), the CH,*/H* coverages are also crucial. The relative stoichiometry of surface species, O* versus CH* and H*, is determined from surface reactions, the blocking by O* (desorption), and the adsorption fluxes of reactants. The activation energy of CH4 adsorption is higher for Pt than Rh (ca. 10 kcal mol-' versus ca. 5 kcal mol-1),4 leading to relatively low coverages of C* and H* on Pt and favoring complete combustion products. In fact, this latter mechanism has been suggested as responsible for the superiority of Rh compared to Pt.30 Sensitivity analysis using the thin film reactor model discussed above (Section 4) seems to support this idea partially. In particular, based on the calculated BOC activation energies and assumed reaction pre-exponentials, the adsorption flux of CH4 and the desorption of O* are about equally important to the activation energy of CO* desorption in determining CO selectivity. For H2, because the overall H 2 0 formation activation energy is comparable to H* desorption, the adsorption flux of CH4 and the desorption of O* are the most important factors influencing selectivity. Finally, even though the activation energy of desorption for O* is higher on Rh than Pt, the higher activation energy of CH4 adsorption on Pt,4 along with the energetics discussed above, offsets the surface stoichiometry leading to superior performance of Rh for comparable operating conditions. The sticking coefficients of reactants are also important.

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Figure 4.7 A schematic of the catalytic reaction mechanism of CH4 with 0 2 . Adsorption of CH4 and 0 2 to the surface is followed by sequential oxidation of C* with O* to form CO and CO,. Desorption of H* to H2 competes with the H 2 0 formation path.

7

A Quantitative Reaction Mechanism for Oxidation of Methane

In addition to partial oxidation, complete catalytic combustion of methane is also an important process because of its potential for efficient energy production with relatively low pollutant emission^.^^^^^ However, partial and complete combustion typically operate under different conditions, making the development of a single consistent reaction mechanism that is suitable for both processes a formidable challenge. While global reactions (as discussed above) give some physical insight into the process, they cannot quantitatively capture all of the important chemical interactions between the reactants and product gases and the catalyst surface. For Rh- and Pt-catalysed oxidation of CH4 with 02, several surface reaction mechanisms have been proposed in the l i t e r a t ~ r e . ~ ~ ~However, ~-~ all of these mechanisms have some inherent drawbacks (see ref. 51 for specific examples). Below, we discuss the specific features of these mechanisms as well as some generic deficiencies. 1956363

7.1 Currently Proposed Mechanisms 7.1.1 Selectivity Mechanisms - The detailed surface reaction mechanism proposed by Hickman and Schmidt4 has been constructed mainly by compiling kinetic parameters (reaction pre-exponentials and activation energies) for elementary reactions from surface science measurements reported in the literature, with the rest of the parameters adjusted to reproduce their experimental data. A schematic of this reaction mechanism is shown in Figure 4.7. Briefly, the mechanism describes the adsorption of CH4 on the metal surface as a single activated step, resulting in the formation of C* and 4H*. These adsorbed species react with O* (formed by the dissociative adsorption of 0 2 ) to form the main oxygenated products, CO, C 0 2 and H 2 0 . Adsorbed H* can alternatively undergo associative desorption to form H2. The same reaction steps were assumed for both Pt and Rh catalysts, with different kinetic parameters. A coverage-dependent activation energy of desorption for CO was used only for Rh. Modeling results using this reaction mechanism showed reasonable agreement to their experimental data obtained from Pt- and Rh-

126

Catalysis

coated monolith^.^ However, this reaction mechanism cannot predict catalytic ignition and extinction, which are vital steps in reactor startup and shutdown.51In addition, for a different set of syngas experiments by Alibrando et al., the model predictions were in poor agreement with experimental selec~ ~ recently proposed a surface reaction tivity.78 Veser and F r a ~ h a m r n e rhave mechanism to reproduce the selectivity data of Hickman and Schmidt that includes competitive adsorption of reactants (i.e. all the gas-phase species adsorb onto the same surface site), modified kinetic parameters, and additional intermediate surface species. There was close agreement with experiments. However, this mechanism does not accurately predict catalytic ignition. 7.1.2 Catalytic Ignition Mechanisms - Bui et al.40 have analysed the Hickman and Schmidt reaction mechanism4 for catalytic ignition of CH4 over Pt. They have concluded that the assumption of non-competitive adsorption of CH4 and 0 2 is the main reason for the lack of catalytic hysteresis. Another detailed reaction mechanism has been proposed by Deutschmann et al.56 to fit Pt catalytic ignition temperatures. The major reaction paths of this mechanism are similar to Hickman and Schmidt's with some exceptions. For example, Deutschmann et al.'s mechanism also assumes competitive adsorption of CH4 and 0 2 , with CH4 chemisorption being non-activated, and described by sequential thermal H decomposition steps. Furthermore, coverage-dependent desorption rates of O* and H* have been included. In terms of mechanism performance, the competitive adsorption enabled prediction of catalytic ignition, in agreement with experimental values (significant deviations appear at fuel-rich condition^^',^^). However, the mechanism fails to predict the conversions and selectivities observed at high temperatures and fuel-rich conditions typical for syngas production, and catalytic extinction. The mechanism of Aghalayam et aL5' is currently the only one that can simultaneously predict catalyst ignition, extinction and syngas selectivities.

7.1.3 Other Proposed Reaction Pathways - Other researchers in the partial oxidation field have also proposed reaction pathways based on their experimental results, providing a qualitative picture of the overall process. For example, based on product distribution through TAP experiments and surface species observations through DRIFTS experiments, surface reaction paths which include CH, species, carbonate and the interconversion of atomic oxygen between adsorbed oxygen and metal oxide have been ~ u g g e s t e d . ~ ~ . ~ ~ Mallens et al.30 have also speculated that adsorbed oxygen plays an important role in the oxidation of CH4. In particular, they have postulated that the formation of CO occurs through Mars-van Krevelen redox reactions between adsorbed atomic oxygen and the metal catalyst.

7.2 Limitations of the Existing Detailed Surface Reaction Mechanisms - The primary reason for the lack of a single comprehensive and reliable reaction mechanism for CH4 oxidation is that most of the kinetic parameters have been obtained from surface science experiments, which are often performed at

4: Partial Oxidation of Light Alkanes in Short Contact Time Microreactors

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idealized conditions of low pressures and/or single crystals. In the cases where experiments are performed under realistic conditions (higher pressures and polycrystalline or supported catalyst), reaction parameters are usually fitted to only a single type of experiment, limiting the regime of applicability of existing mechanisms. Furthermore, strong heat and mass transfer effects and complex (undeveloped) flow patterns strongly disguise kinetics, making quantitative analysis of these experiments challenging. Currently, several important features observed experimentally have not been incorporated as yet into available detailed reaction mechanisms (such as the ones proposed by Hickman and Schmidt4 and Deutschmann et aZ.55),and will further be discussed below. 7.2.1 Adsorption of CH4 - The adsorption of CH4 is often written as a single step to produce C* and 4H*.4 This is a relatively important step as it can influence both the catalytic ignition temperature and the selectivities to syngas. Surface science experiments of CH4 adsorption on Pt( 11 1)80-82indicate that the primary products in CH4 adsorption on Pt are CH3* and H*. The experiments using deuterated methane indicate that tunneling effects can be significant at lower temperatures. In addition, the theoretical predictions of Shustorovich and Bell7' using the BOC theory indicate that the adsorption of CH4 to CH3* and H* is activated by ca. 6 kcal mol-I for Pt. More surprisingly, the subsequent thermal decomposition steps of H from CH,* are predicted to be activated by ca. 20 kcal mol-'. Besides thermal decomposition steps, experimental data indicate that the formation of coke (presumably C* and CH*) is enhanced by the presence of 0 2 . As an illustration, Figure 4.8 shows the catalyst temperature as a function of time for four different inlet reactant mixtures.39 Before every experiment, a H2-air mixture was combusted over the catalyst at ca. 1000 K for 15 minutes. In each experiment, the flow velocity and the power provided to the Pt foil were kept constant. In addition, the high fraction of N2 ensures that there is little difference in the specific heats and heat transfer characteristics from case to case. Hence, differences in the foil temperature can be attributed to the amount of coking, since carbon increases heat losses through radiation, due to its higher emissivity compared with bare Pt and endothermic reactions. Whereas for N2, the foil temperature does not change with time, a small fraction of 0 2 in a N2-CH4 mixture results in a significant decrease of foil temperature. These results indicate that O* (or some 0-containing species such as OH*) helps in the abstraction of H* from CH,* species. In fact, the spontaneous decomposition of CH4 following adsorption is quite slow in comparison to CH4 decomposition in the presence of 0*,as shown in Figure 4.8. Such oxygen-assisted CH,* decomposition is consistent with BOC calculations predicting lower activation energies than pyrolysis.39The important role of O* in decomposing CH,* has also been proposed by experimental studies of CO2 reforming of CH4 over supported Pt83 and Ni84 catalysts. These results indicate that two CH4 decomposition mechanisms can occur, depending upon reactor operating conditions. Under oxidizing conditions (surface 0*-rich), CH,* decomposition would likely occur through 0*-assisted paths. However,

Catalysis

128

I

800’;

’

30

1

I

1

I

60 90 Time (min)

1

I

120

I

150

Figure4.8 Catalytic foil temperature as a function of time for four different inlet mixtures indicated. The conditions are identical to Figure 4.4, with the same current input for all four experiments. Visible carbonformation is not observed for 5% CH4 in N2. Fractional addition of 0 2 to CH4-N2 mixtures causes visible carbon formation with a concomitant decrease in catalyst temperature, indicating the important role of 0, in the catalytic decomposition of CH4 (data from reJ: 39).

under reducing conditions (surface 0*-starved), pyrolysis should be the dominant pathway. 7.2.2 The Role of Adsorbed Oxygen - Both the steady-state and the TAP experiments indicate that oxygen on the catalyst surface is a strong factor which ultimately determines partial or complete oxidation of CH4. However, describing this essential feature in terms of reaction kinetics remains difficult. Part of the reason is that many different forms of oxygen can exist on the catalyst surface. As discussed above, various researchers have speculated upon redox cycles involving CH4 reducing metal oxide sites and oxygen re-oxidizing the reduced metal atom.30.75Surface science experiments indicate that depending on the temperature range, adsorbed oxygen can exist in different forms. For example, below about 160 K, oxygen adsorbs onto Pt in its molecular state as a peroxo-like species (022-).85986At temperatures greater than around 160 K, the oxygen molecule dissociates into its atomically adsorbed In the temperature range of 600- 1 100 K, platinum oxides form, mainly PtO,, whereas both PtO and Pt304 have also been speculated to exist.89 In some cases, at temperatures of ca. 1000 K, ‘dissolved’ or ‘subsurface’ oxygen has also been observed and shown to depend on the impurities in Pt crystals.90Mallens et al. reported chemisorbed oxygen as well as Rh2O3 through X-ray diffractometry and X-ray photoelectron spectroscopy measurements under oxidizing condition~.~~

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For a realistic partial oxidation reactor, spatial gradients in temperature and species' concentrations are expected as reactants are consumed and products are formed. This means that even though the inlet feed composition is fuel-rich, various parts of the reactor may experience different catalyst state, such as bare catalyst, catalyst oxide, dissolved oxygen within the lattice, or chemisorbed oxygen with different kinetics. As a result, a different CH4 decomposition path should dominate at different spatial positions (O*- and OH*-assisted reactions near the entrance and pyrolysis downstream), as the 0 2 is being depleted along the length of the reactor. It would then be desirable to elucidate the catalytic chemistry as a function of different types of oxygen and oxides present, the coverages of other adsorbates which may potentially alter reaction energetics and pathways, and rates of interconversion of various oxygen forms. 7.2.3 Adsorbate-Adsorbate Interactions - The influence of adsorbate-adsorbate interactions on energetics is another well known phenomenon. For example, surface science experiments have shown that the activation energy for associative desorption of adsorbed atomic oxygen on Pt depends on its surface ~~ coverage, decreasing by ca. 8 kcal mol-I for 0.25 m ~ n o l a y e r .While researchers have already introduced this concept in oxidation reactor scale m ~ d e l i n g , ~its~ application .~~ has been rather simplistic, allowing for only linear interactions for a single type of adsorbate. One important consequence of adsorbate-adsorbate interactions is that they can potentially change the activation energies of most surface reactions, as shown by S h u s t o r o ~ i c hand ~~ Park et aL4' Currently, only the adsorbate interactions of H*-H*,92993 0*0*,91 and 0*-0H*94-96 have been well quantified, under specific conditions. The interactions between each adsorbed species and all other adsorbates need to be analysed and incorporated into reaction mechanisms for a more complete and reliable representation of catalytic reactions. While this feature might be difficult to capture realistically in a mean-field type of approach, there has been some progress using multi-scale models.97 7.2.4 Other Limitations - Several other aspects for further work include catalyst aging, carbon or coke formation on the catalyst, thermodynamic consistency and hydrocarbon coupling reactions. ( 1) Catalyst aging is a commonly encountered feature in

experiment^.^^'^-'^^^^

For some systems, such as H2-02 oxidation over a Pt foil, Fernandes et al.36have found that reactivity increases with catalyst use. On the other hand, Slaa et al.15 reported decreasing conversions of CH4 and O2 in a monolith reactor under partial oxidation conditions. Heitnes Hofstad et a1.,18 using a Pt wire gauze for partial oxidation, reported decreasing conversion of CH4 and CO selectivity, with a concomitant increase in 0 2 conversion and C02 production with increased catalyst use. Changes in surface morphology, evaporation or reactive etching of the catalyst, or poisoning of the catalyst are possible explanations for the changing behavior of the catalyst that need to be explored further.

130

Catalysis

(2) Deposition of carbon or coking is an important factor in partial oxidation reactors under fuel-rich conditions. The formation of coke has been observed for various systems, e.g. Ni catalysts in a monolith reactor6* and Pt foils46(see also Figures 4.5 and 4.8). However, it is not yet fully clear how coke affects reaction paths. For example, Figure 4.5 shows that it is possible for the catalyst to be ignited even with visible coke. In addition, even under conditions where the catalyst is coked, formation of CO through the reversed Boudouard reaction has been ~ b s e r v e d . ~ ~,~~ The removal of coke by gas-phase hydrogen through the reaction C + 2H2 + CH4 is yet another possibility. Formation of multilayers of coke on the catalyst and heterogeneous reactions in the presence of carbon need to be further clarified and incorporated into reaction mechanisms. (3) Reaction mechanisms need to be thermodynamically consistent. Mechanisms for methane, CO and H2 oxidation often violate thermodynamic consistency due to the lack of available thermodynamic databases of specific heats or heats of formation (analogous to the counterparts of gasphase reactions, e.g. ref. 98), especially for intermediates. This is a problem in optimization of activation energies when comparison to experiments is attempted. While the heat of a reaction is made equal to the difference in activation energies of the forward and backward reactions during fitting, it can be shown that the number of linearly independent degrees of freedom is equal to the number of surface species rather than the number of surface reactions. The latter is typically much larger than the former for such complex mechanisms. In addition, the net reaction cycle can produce or consume several kcal per mole of reactants, a serious miscounting as heat transfer is important for these systems, (4) Finally, under fuel-rich conditions and high surface coverages of hydrocarbon fragments on the catalyst surface, coupling reactions that lead to higher hydrocarbons also need to be considered. For example, production of various C2 species such as C2H6, C2H4 and CzH2 has been observed by Witt and Schmidtz3 under partial oxidation conditions, with a combined selectivity as high as about 10%. Details of individual reaction steps are not currently available. A newly proposed approach on the construction of heterogeneous reaction mechanisms overcomes, to a large extent, the last three limitations, and its application to CH4 oxidation is de~irable.~' 8

Partial Oxidation of Ethane and Higher Alkanes

The importance of the reactions of methane stems from the fact that it is the predominant component of natural gas. However, other alkanes, particularly ethane, which is the second largest fraction of natural gas, may play a significant role in the reactions of methane. More importantly, higher alkanes are feedstocks for a number of chemical processes, such as the production of

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ethylene. Thermal cracking of ethane, which is the current commercial process for ethylene production, is highly endothermic and requires a large heat input. Besides, it produces a large amount of C02 (a greenhouse gas), and requires considerable steam input to prevent carbon deposition. Any process improvements or modifications of this reaction would have a large economic impact on the chemical process industry. 8.1 Production of Olefins from Alkanes using Pt and Rh - A number of researchers10~11*22~2s*99-103 have recently attempted to extend the principles of short-contact time autothermal catalytic reactors developed for methane partial oxidation to higher alkanes. In a series of studies, the ability of Pt/Rh gauze catalysts and Pt- and Rh-coated foam monoliths to convert ethane, propane and n-butane to more desirable chemicals has been explored. Some of these results are presented in Table 4.5. For ethane oxidation over Pt gauzes and Pt-coated monoliths, high selectivity to ethylene (> 60%) and conversion (> 60%) have been obtained with contact times of the order of rnilliseconds'O to ca. 50 microseconds.Iol Use of Pt-Sn alloys as catalysts leads to improvements in both selectivity and conversion. Comparable conversions and selectivities to olefins have been achieved for higher alkane oxidation over Pt-based catalysts, such as propane,lo4 pentane and hexane. Io2 However, in contrast to partial oxidation of CH4 to syngas, alloying Rh with Pt leads to decreasing conversion and selectivities, regardless of the fue1.99~101~102 Use of a combined separationreactor for the dehydrogenation of i-C4H 10 has also shown potential for improved conversion and selectivities by shifting equilibrium limitations. Io5 Addition of H2 to the feed has a positive effect on performance.ll-lO1 Regarding the role of H2 in improving selectivities to olefins, it has been proposed that H2 is preferentially oxidized to H20 on the catalyst surface. This has two performance-promoting effects: first, preferential consumption of 0 2 that can lead to deep alkane oxidation; and, second, an increase in reactor temperature, due to the exothermicity of H20 formation, that promotes alkane dehydrogenation. Overall, Pt is an active catalyst for the production of olefins in oxidative alkane dehydrogenation in short-contact time reactors. Among the other catalysts studied, Rh is found to be non-selective to olefins, whereas Pt-Sn alloys gave higher selectivities than Pt. Aside from the catalyst, reactor design (such as heat transfer, distributed feed of reactants, etc.) plays a key role in performance as the experiments with co-feeding of H2 indicate.

8.2 The Mechanism of Alkane Dehydrogenation - The short-contact times in alkane dehydrogenation processes suggest a purely heterogeneous reaction mechanism. In fact, a detailed surface reaction scheme has been proposed and found to give reasonable predictions of data.9 However, various experiments without the catalyst have been conducted for ethanelO' and propane oxidationIoOrevealing that, even in the absence of a catalyst, high conversions of the fuel and selectivity to olefins may be obtained. In fact, the presence of a

Contact time

1-10ms 1-10 ps 50 ps 50 ps 1 ms 1 ms 1-10 ps ca. 1 ms ca. 5 ms ca. 5 ms ca. 5 ms

Catalyst

Pt/A1203 Pt/Rh gauze Pt gauze Pt/Rh gauze Pt/A1203 Pt-SdAl203 Pt/Rh gauze Pt/A1203 Pt/Al203 Rh/A1203 Pt/A1203

Reactant

!% Conv.

70 62 65 60 64 83 58 60 60 40 60

!% S(o1eJins)

Schmidt and Goetsch and Schmidt99 Lodeng et al."' Lodeng et al."' Bodke et al." Bodke et al." Goetsch and Schmidt99 Beretta et Dietz et a1.Io2 Dietz et al. Io2 Dietz et a1.Io2

Reference

Table 4.5 Selected summary of the experimental data for catalytic oxidation of higher alkanes at short-contact times

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133

catalyst seems to be detrimental to olefin production at low conversions, as compared to a non-catalytic system.lO*On the other hand, in the absence of the catalyst, the induction times are found to be high (of the order of seconds), as compared with milliseconds for Pt.lm+lol Calculations were recently carried out with a detailed gas-phase reaction mechanisrn,lo6 and the selectivity to olefins vs. conversion of ethane was compared to experimental results' for various inlet ethane/oxygen/hydrogen compositions. The agreement was poor for all systems studied. The improvement in selectivity due to the addition of hydrogen in the feed was underpredicted, whereas the induction times were overpredicted. Selective alkane dehydrogenation in short-contact time catalytic reactors has thus been proposed to be a homogeneous-heterogeneous process, involving intimate coupling between gas-phase and surface-phase chemistry. Reliable detailed reaction schemes for both gas- and surface-phase chemistry are necessary in order to model this system.

'

9

Conclusions and Outlook

The experimental efforts of various researchers have shown that partial oxidation of CH4 over Rh-catalysts and the oxidative dehydrogenation of C2H6 over Pt-catalysts are viable and are potentially more attractive methods to syngas and ethylene production over the traditional methods of steam reforming and thermal cracking. Major advantages of short-contact time oxidation reactors include autothermal operation with significant economic and environmental impact, high yield, large throughput, small reactor sizes, a more suitable H2 :CO ratio for downstream utilization of syngas, and easy scale-up by increasing the reactor diameter (although better heat and mass transfer management would also be essential for this). These reactors run under rather severe conditions of high pressure and temperature, short-contact time, complex fluid flow, and with strong coupling of heat and mass transfer with surface- and possibly gas-phase kinetics. Such conditions impose serious challenges to both experiments and simulations. Despite the extensive experimental and simulations efforts, our understanding of such processes is far from being complete. In order to realize the potential of partial oxidation in short-contact time reactors, several factors, outlined above, need to be further studied. Some critical aspects include: 0

0

Kinetic studies of very fast, exothermic reactions - Novel reactor configurations are needed which approach realistic operating conditions while allowing deconvolution of gas-phase from surface chemistry and chemistry from heat and mass transfer. Reactor scale-up - A combined effort of experiments and modeling is needed to study reactor behavior at commercial scale operating conditions, including high pressures and large reactor sizes. The engineering aspects of temperature control and catalyst management also need to be considered.

134

Catalysis

Safety - The bifurcation behavior should be elucidated at high pressures for safe and optimal reactor operation. Again, features such as flammability limits and homogeneous ignition are sensitive to reactor geometry, and need to be carefully considered during scale-up to industrial conditions. Surface reaction mechanisms - The major theoretical impediment is systematic and rapid development of elementary surface reaction mechanisms that can be extrapolated, are thermodynamically consistent, and are validated with a wide range of types of experiments and conditions. The outcome of this effort will pave the way for efficient catalyst screening with minimal experimental information and rapid commercialization of such processes. Gas-phase reaction mechanisms - While the gas-phase reactions of small alkanes have been well studied for fuel-lean mixtures, the understanding of the chemistry of fuel-rich mixtures, typically of interest in partial oxidation, is at an embryonic stage. Knowledge and techniques developed in the combustion community can be an asset in overcoming this problem. Computational efficiency - The computational demand of high dimensionality fluid mechanics, multicomponent transport, and detailed chemistry codes prohibits parametric studies and reactor optimization. Techniques need to be developed to accelerate codes by up to two orders of magnitude in order for simulations to be an integral part of experimentation and optimization.

Acknowledgment

Acknowledgment is made to the Office of Naval Research with Dr G.D. Roy through a Young Investigator Award under contract number N00014-96-10786 and to the National Science Foundation (CAREER CTS-9702615) for support of this work.

References 1 2 3 4 5 6 7 8 9 10

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5

Indirect Liquefaction - Where Do We Stand? BY YONGQING ZHANG AND BURTRON H. DAVIS

1

Introduction

Following its discovery by Fischer and Tropsch,’ the desirability of converting synthesis gas to hydrocarbons has undergone many cycles ranging from the view of an immediate, urgent need for commercialization, to periods of benign neglect. In the 1930s, because Germany did not have an internal source of petroleum, but did have a vast reserve of coal, interest in Fischer-Tropsch Synthesis (FTS) was great. This interest is illustrated by the visit in 1937 of Chancellor Adolf Hitler and Albert Speer to the leading coal research center in Germany and their discussions with its leader, F. Fischer.2 A feeling for the perceived concern for the petroleum supply situation at that time is represented in the f~llowing:~ ‘Jersey (now Exxon) sought still other methods for making motor fuels of high quality. For example, it investigated the Fischer-Tropsch hydrocarbon synthesis process, developed by Ruhrchemie, A.G., which converted brown coal into liquid fuel. In 1938 and 1939, patents for this process outside Germany were transferred by Ruhrchemie to Hydrocarbon Synthesis Corporation, in which Standard Oil Development (now Exxon) took 680 shares, Shell and Kellogg 425 each, and I.G. Farben 170. Both Great Britain and France considered the building of plants using this synthetic process for providing aviation gasoline, but they had been unable to accomplish anything definite by the time the war broke out in Europe in 1939.’ Germany developed a significant synfuels industry prior to and during WWII. However, their total peak production was small-scale in comparison to the US consumption of petroleum today. Furthermore, much of the German transportation fuel was derived from coal tars and direct coal liquefaction rather than from FTS. During WWII, the German FTS production was mainly conducted at atmospheric pressure using cobalt catalysts in fixed-bed reactors. Following WWII there was a perceived shortage of petroleum crude. In many countries, crash programs to develop a synfuels industry were initiated. US officials were among those who recognized the potential limits of petroleum crude. Major efforts were directed toward developing the FischerTropsch (F-T) synthesis processes in the US. One of these efforts was led by Catalysis, Volume 15 0The Royal Society of Chemistry, 2000 138

5: Indirect Liquefaction - Where Do We Stand?

139

Dobbie Keith, the engineer responsible for building the Oak Ridge Weapons facility, and his efforts led to the construction and operation of a commercialscale plant at Brownsville, T e ~ a s . While ~ - ~ the subsequent discovery of major petroleum reserves in the Middle East prevented this plant from being a commercial success, it introduced the use of large-scale oxygen plants for the production of syngas. The Brownsville plant utilized a fixed fluid-bed catalyst technology. At the time of the shutdown, the Brownsville plant was approaching a technically viable operation. On a smaller scale, pilot plants for both direct and indirect coal liquefaction were constructed and operated by the US Bureau of Mines in Louisiana, Missouri in 1 9 5 0 ~ The . ~ Louisiana Fischer-Tropsch plant could utilize a version of a slurry-bubble column reactor that was being tested in Germany, primarily in work led by Herbert Kolbel.*

2

Today's Commercial Operations

In general, three operations are needed for a commercial Fischer-Tropsch plant to produce transportation fuels: syngas generation, Fischer-Tropsch synthesis and conversion of light- (C,-C,) and heavy wax (about C18+) products to transportation fuel range or developing markets for these materials. For a process based on natural gas and heavy-product F-T synthesis, the capital cost is considered to be approximately in the range: syngas generation :F-T synthesis : hydrocracking = 50 : 35 : 15. For operation with coal as the carbon source, the capital cost is divided approximately: syngas generation : F-T synthesis = 67 :33. A recent Arthur D. Little, Inc. study provides estimated cost components of a gas-to-liquids unit (Figure 5. 1).9 Natural gas Pdudion

Oxidant

I

Syngas Conversion

Pmducl Upgrading

,

I

I

I Convemion &.20hbl

Upgrading

Total = $17.5Wbbl

-I

$3.mbl

Feedstock cost

I

Figure 5.1 Typical gas-to-liquids (GTL) product cost breakdown for a 100000 bbl day-'plant in North Field, Qatar (from ref: 9).

Catalysis

140

Currently there are three major commercial operating groups. Two of these are located in South Africa: the Sasol operation based on coal as the source of syngas and the Mossgas operation that is based on natural gas. The third commercial plant is operated by Shell in Bintulu, Malaysia and is based on natural gas. Rentech, Inc. claims to have operated '. . . at the commercial scale in 1992 and 1993'.'* However, the Rentech commercial operation at Pueblo, Colorado was to produce approximately 235 barrels of liquid hydrocarbons per day, and the failure to produce sufficient methane limited the operation to a few months. 2.1 South Africa - In South Africa the fear of being boycotted due to apartheid led to major efforts to develop for the country an independence from crude imports. The method used by the South Africans was the Fischer-Tropsch synthesis. This led first to the creation of Sasol and, more recently, Mossgas, and the introduction of a government subsidy. However, since crude is now readily available to South Africa, political and economic considerations have forced the South African Government to begin phasing out the subsidy for transportation fuels. With vision, Sasol anticipated the eventuality of losing its subsidies and has gradually shifted its emphasis from producing only transportation fuels to deriving a significant fraction of its operating profits from the sale of chemicals and petrochemical feedstocks. Today Sasol and Mossgas produce about 130 000 bbl day- corresponding to about 35% of the transportation fuels used in South Africa. While the South African production is impressive, its total output would only satisfy the needs of a medium-sized US petroleum refinery. Put in another perspective, the South African total production could replace only about 1% of the US consumption of crudes.

',

Sasol - Initially, Sasol utilized Arge fixed-bed reactors and iron catalysts. Subsequently they employed circulating fluidized bed reactors. Today the circulating fluidized bed reactors, combined with the fixed fluidized bed reactor that has recently been developed, account for about 90% of the Sasol production." The Arge reactors are being replaced by slurry phase bubble column reactors, and a 2400 bbl day - slurry reactor has operated successfully for almost ten years.12 Sasol still employs iron-based catalysts in all of its reactors. The production at Sasol is approximately 100000 bbl day-'. Of this total, about 10% is produced in the Arge and slurry reactors. Because of the very exothermic nature of the F-T synthesis (FTS), the introduction of the slurry reactor, with its ability to allow rapid heat transfer and isothermal operation, offers a decided advantage. Sasol features two types of product selectivities. Because the FTS is a 'polymerization reaction' using 'C1 monomers', the products fit an AndersonSchulz-Flory distribution (Figure 5.2). This distribution means that only methane can be produced as a pure product, with the ratio of chain termination to chain growth, a,being zero; all other values of a produce a mixture of products (Figure 5.3). The fluid-bed reactors must operate at an a value that is 2.2

5: Indirect Liquefaction - Where Do We Stand?

141

I 0'

10

10

ld

a

4

8

12

16

.

.

.

a

.

20

24

P(= carbon number) Figure 5.2 Theoretical Anderson-Schulz-Flory plots for U-values relevant for the F-T synthesis.

0.2

0.4

0.6

0.8

1.o

a

Figure 5.3 The Fischer-Tropsch reaction leads to a product distribution common to polymerizations from Cl monomers. The value of u (the probability for chain growth) depends on the relative rates of the propagation and termination steps.

142

Catalysis

Figure 5.4 Schematic showing sizes of Sasol fluid-bed reactors (compiled from T. Shingles and D. H. Jones, ChemSA, August 1986, 179- 182 and B. Jager, M.E. Dry, T. Shing1esandA.P. Steynberg, Catal. Lett., 1990, 7, 293).

small enough so that heavy waxes are not produced. Non-volatile waxes would condense on the catalyst particles, causing agglomeration so that the fluid state of the catalyst bed would be destroyed. Thus, the Sasol fluid-bed reactors are operated to produce 'chain-limiting' products. In general, the high temperature operation produces little wax, whereas the low temperature operation produces high molecular weight wax as the dominant product. Sasol operates four types of reactors at the commercial scale: tubular fixedbed (TFB), Synthol circulating fluid bed (CFB), Sasol Advanced Synthol (SAS; fixed fluid bed) and slurry phase (SP). During more than 40 years of operations, the Synthol circulating fluid-bed reactors have grown in size (Figure 5.4) until it would be difficult to scale up to a larger size. Sasol has recently introduced an advanced Synthol reactor (SAS) that makes use of conventional solid-gas fluidization that offers the advantage of requiring only about half the size of the CFB for the same production rate. Because of the decrease in reactor size and the elimination of the need to recirculate large masses of catalyst, the SAS reactor requires only about half the capital cost. The SAS is a much simpler reactor making it easier to operate, lowering operating cost by eliminating the need for catalyst recycle, reducing catalyst consumption by about 40%and reducing maintenance to only 15% of the CFB system. The SAS reactor accommodates higher gas loads and more cooling coils; this allows for larger capacity equipment with the advantages of economy of scale. Whereas the capacity of the CFB reactor in use at Sasol today is 6500 bbl day-', the SAS reactor has a capacity of 11 000 bbl day-' and it is anticipated that it has the potential for producing 20000 bbl day-' (Table 5.1)."

5: Indirect Liquefaction - Where Do We Stand?

143

Table 5.1 Sasol Fischer-Tropsch commercial reactors - capacities (bbl day- ') (from ref: 1 1 ) ~~~

Total installed capacity Capacity per reactor Potential per reactor

CFB

SAS

TFB

SP

1 10 000 6500 7500

1 1 000 1 1 000 20 000

3200 500-700 1550

2500 2500 10000

10000

Slurryqhase

Tubular fixed bed

0

195s

1987

1991

1993

Future

Figure 5.5 Reactor capacity increasefor LTFT reactor (from reJ 13).

A similar increase in size occurred with the tubular fixed-bed reactors as seen with SAS reactors (Figure 5.5) where the output increased from 500 bbl day-' reactor-' in 1955 to 1200 bbl day-' reactor-' in 1991.13 These increases in output may be viewed as evolutionary advances by operators of a commercial plant. In 1993, Sasol introduced, for them, a revolutionary advance in switching from the tubular fixed-bed reactor to a slurry bubble column reactor (SP) with 2400 bbl day-' reactor-' output. It is anticipated that the SP reactor can be scaled to 10000 bbl day-' reactor-', and even more. In a commercial operation, economic incentive drives operators to make scientific and engineering advances that lead to improvements in both the scale and the efficiencies of the operation. Sasol has made significant improvements in the energy efficiency of the Lurgi gasifiers during many years of operating them for the production of syngas from coal. These improvements have in turn allowed for the production of the products of the F-T process at a lower cost. The 16 conventional CFB Synthol reactors have been replaced with eight SAS reactors: four with a diameter of 8 m and four with a diameter of 10.5 m. The size differences were determined by the plant configuration. The first SAS reactor was commissioned in 1996. The Synthol reactors will be decommis-

144

Catalysis

Table 5.2 Selectivity (carbon basis) of Sasolprocesses (fromref 11) Product ~

LTFT

HTFT

~~~~~~

CH4 C2 to C4 olefins C2 to C4 paraffins Gasoline Middle distillate Heavy oils and waxes Water soluble oxygenates

4 4 4 18 19 48 3

7 24

6 36 12

9 6

sioned and left on-site for the immediate future.14 The use of SAS reactors rather than slurry reactors was also determined by the in-place downstream equipment. In general, Sasol operates their reactors in two modes: low-temperature (LTFT), high-alpha; and high-temperature (HTFT), low-alpha. The typical products from the LTFT are about 50% heavy oils and waxes (Table 5.2). The product make shown in Table 5.2 can be fitted to a two-alpha plot with the higher-alpha, heavier product polymerization accounting for about 75% of the total F-T products. The products from the high temperature operation are representative of a low-alpha product distribution, producing more of the light CI-C4 products (about 33% of the total hydrocarbon products) with the dominant fraction of the products in the transportation range. The ideal F-T situation is one in which the Cl-C4 products are decreased significantly (to less than 10%) and, at the same time, have the remaining products in the transportation fuel range. To date, this has not been accomplished in reasonable sized reactors that are operated at steady-state conditions. A significant fraction of the products from the fluid-bed reactors at Sasol are in the C I - C ~range and these are not in the boiling range of transportation fuels. In many locations this would be a severe disadvantage but apparently this is not the case in South Africa today. Since South Africa did not have abundant reserves of natural gas, methane and ethane were salable as town gas. The alkenes in the C2-C4 fraction have been utilized by Sasol to develop a strong position as a supplier of feedstocks for the petrochemical industry. In fact, many of the compounds being utilized as petrochemical feedstocks are valued at 2-4 times their value as transportation fuels. This has led Sasol to reserve Sasol I, using the slurry bubble column reactor and Arge fixed-bed reactors, for the production of chemical feedstocks, including paraffin wax. Until sufficient plants are on-stream to saturate the market, petrochemicals will allow Sasol to reap significant benefits from this part of their operation. Sasol has entered the petrochemical business with great determination and has been able to develop significant new business areas. For example, the Sasol

5: Indirect Liquefaction - Where Do We Stand?

145

Synthetic Fuels operating profit increased 28% for the year ending June 1997 over that of the previous year. Even without including the tax subsidy from the government, Sasol now is able to operate at a profit.

2.3 Mossgas - Mossgas, established in 1989, is a subsidiary of South African Parastatal Central Energy Fund.15 The life of one of the two initial gas fields (FA) used as the source of their feedstock was found to be significantly shorter than originally expected, and was expected to run out in 1997. Consultants were used to advise on the potential sale of Mossgas but, after reviewing the bids, the government decided against privatization. Subsequent efforts to use satellite facilities to tap nearby pockets of gas and improvements in compression in the field are expected to extend the life of FA by 3-5 years. Current estimates indicate that there will be sufficient gas from FA and its satellites to maintain synthetic fuels production at 30200 barrels of refined product per day until 2001. Production from another gas field is expected to provide gas to operate Mossgas until the end of 2005. The onshore Mossgas plant receives more than 5 million m3 day-' gas and associated condensate of more than 9000 barrels a day. The methane reforming plant is considered to be the world's largest. The reactors are of the circulating fluidized bed-type utilized by Sasol. The iron catalyst is manufactured on site. The plant consists of three identical trains. Since an iron catalyst is utilized, some water-gas shift will occur and, even with carbon dioxide recycle, it would appear that the plant would have a surplus of hydrogen. Mossgas has had an operating surplus after the first year of operation even before taking credit for the synlevy and tarriff protection. However, the stateowned plant is uneconomic. l 6 Long-term operations at Mossgas are dependent upon the development of economically recoverable gas reserves in addition to those now available to the company, the development of another source of the synthesis gas, or conversion to another operating process. 2.4 Shell - Shell has developed a process to produce middle distillate (SMDS). 7*1* The process consists of syngas generation, Fischer-Tropsch synthesis, the heart of the SMDS, and selective hydrocracking (Figure 5.6). The synthesis gas is generated using a version of the Shell methane reformer (SMR) that utilizes autothermal partial oxidation, first utilized by Shell in 1956 for the gasification of heavy resid. SMR produces a Hz :CO ratio that is close to 2 : 1, the consumption ratio in the F-T synthesis step using a cobaltbased catalyst. However, the complete balance requires that some hydrogen in excess of this ratio be generated, and this is done using conventional catalytic steam reforming of methane. The F-T synthesis is carried out in fixed-bed reactors configured similar to the multi-tubular Arge reactors that are utilized in commercial operations at Sasol. Shell acquired the F-T technology developed by Gulf-Badger that included work with promoted cobalt catalysts and both fixed-bed and fluid phase reactors. Gulf workers had shown the beneficial effects of the incorpora-

Catalysis

146 SYNTHESlS

SYNGAS

CONVERSION

MANUFACTURE

I

SGP

+ I

HEAVY

NA 0, E PLANTLJ GAS

PARAFFIN CONVERSON

HMU"

CH, t 10 2 - t C 0 2

+ 2H2 --c (-CH2-) + H 2 0

t

DISTILLATES

HEAVY PARAFFINS

Figure 5.6 Shell Middle Distillate Synthesis (SMDS): basic concept (HMU* =hydrogen manufacturing unit) (from ref. 17).

tion of ruthenium in a cobalt-based catalyst and the utilization of Group IIIB and IVB metal oxides. The Gulf workers also showed that an activation technique consisting of a reduction, re-oxidation and re-reduction (ROR) procedure for the promoted cobalt catalysts provided a higher conversion than just a single reduction. The catalyst utilized by Shell in its commercial operation is apparently a silica-supported cobalt catalyst promoted with ruthenium and zirconium oxide, a catalyst composition which presumably falls in the range covered by patents assigned to the former Gulf Oil Corporation. Shell workers provide a comparison of the products obtained from the classical catalyst formulations, those obtained as a result of new catalyst development, and the Shell catalysts. The data indicate that the Shell catalyst has an U-value in the 0.90-0.95 range, indicating a 5-10% probability for chain termination at any carbon number (Figure 5.7). Based upon the product distributions reported by Sasol workers, we calculate that Sasol, operating with a 'classical precipitated iron catalyst' in their slurry reactor which is designed to produce the same product distribution as they obtain in the Arge fixed-bed reactors, produces heavy products with an a-value of 0.95, and this high-alpha pathway produces about 75% of the products. Thus, it appears that even the 'classical iron catalysts' can be modified to produce very high avalues, and one should not limit their view to the need to utilize a cobalt catalyst to produce heavy wax F-T products. The distribution of the products from F-T synthesis is normally considered to follow an Anderson-Schulz-Flory distribution. However, the products produced in the early large-scale plants exhibited a 'two alpha' distribution (Figure 5.8). l 9 With the advent of more sophisticated analytical instruments, plots like those shown in Figure 5.8 have been extended to higher molecular weight products. Most authors show data that support the two (or more)

5: Indirect Liquefaction - Where Do We Stand?

147

PRODUCT. %w

0.75

0.85 0.90 0.95 PROBABILITY OF CHAIN GROWTH CLASSICAL CATALYST 0.80

o

-- ---

N E W CATALYST

SHELL

DEVELOPMENT

CATALYST

-----------t

Figure 5.7 Product distribution in Fischer-Tropsch synthesis (from reJ 17).

alpha product distribution although it is the paraffin fraction that is responsible for this (for example, see refs. 20 and 21). Shell workers,I7 on the other hand, report that, ‘in a few hundred FT synthesis experiments, conducted with various catalyst formulations and a range of operating conditions, the products follow a single ASF distribution with U-values in the range of 0.7 to 0.95’. However, the plot (Figure 5.9) that the authors provide to support this conclusion does not provide a full range of product distributions to cover both sides of the carbon number region where the shift from one U-value to the second occurs. The hydrocracking of F-T products can be effected using conventional hydrocracking processes and catalysts. Typical product distributions of a Fischer-Tropsch product and the products obtained following medium- or high-severity hydrocracking are illustrated in Figure 5.10.l 8 While the data in Figure 5. 10l8indicate that little wax needs to be converted, operations at alpha about 0.95 will produce up to about 50 wt% of product that needs to be hydrocracked to produce transportation fuels. The Shell plant in Malaysia was started up in 1993 and, despite encountering start-up problems, was operated successfully. The process is at the point where it could, in 1994, be ‘. . . considered as technically proved and, subject to local circumstances, commercially viable’.18 The plant was designed to have a production of about 10000 bbl day-’; however, the actual production from the plant over its lifetime has not been widely publicized.

Catalysis

148 100

10

I

0.1

0.02

I

1

5

I 9

I

I

13

17

I 21

CARBON NUMBER

Figure 5.8 Logarithmic plots of moles against carbon number. Hydrocarbons from the Schwarsheide tests compared with those from fluidized synthesis (from re$ 19).

A fire/explosion in the air separation plant damaged the SMDS plant and caused the plant to be shut down since December 25, 1997. The fire/explosion was due to accumulation of debris in the vaporizer of the air separation plant, presumably as a result of the severe forest fires in the vicinity of the plant. Current plans are to repair and modify the plant so that it is back in operation in 2000.22

3

Large Pilot/DemonstrationPlant Operations

3.1 Rheinpresussen-Koppers - A large (1 1.5 tons day- synthesis products; about 110 bbl day-') demonstration plant, which featured a slurry phase reactor, was operated at Meerbeck, Germany during 1952-1953 by Kolbel and co-workers (Figure 5.11).* The feedstock for this work was German brown coal. Work with laboratory and demonstration plants provided the most complete relations for estimation of the hydrodynamic properties of the slurry reactor at that time. For the demonstration plant, the synthesis gas was produced using a

149

5: Indirect Liquefaction - Where Do We Stand? MOLES IN PRODUCT, ARBITRARY UNITS

0

Figure 5.9

5

10

15

20

25

30

35 40 CARBON NUMBER

Typical carbon number distributions (from ref: 17).

PROOUCT CORWOSITION

10r

t h

%w

AFTER IGH-SEVERITY Y OROCRACKING

MEDIUM-SEW R ITY HYDROCRACKING

FISCHER-TROPSCH

0

10

20

30 40 CARBON NUMBER

Figure 5.10 Cmbon number distribution of Q Fischer-Tropsch product before and after selective hydrocracking (from ref: 18).

Catalysis

150

Froduet-water

A

Figure 5.1 1 Flow diagram of the large-scale demonstration plant (Rheinpreussen Process): A = freely separated primary products; B =final gas for recovery of low-boiling hydrocarbons; C = high-boiling primary products; a = compressor; b = gas meter; c = orifice plate; d = sampling intake; e = bubble volume reactor; f = steam collector; g = heat exchanger; h = separator; i = container for separated substances; k =pump; 1 = condenser; rn = CO2 expulsion unit; n =pressure cfilter; o = contact stirring container and container for suspension removed from reactor; p = centrifuge; q = mash oil (from re$ 8).

Koppers water gas generator, with some of the CO2 generated during synthesis being recycled. The reactor was 1.55 m in diameter and 8.6 m high (Figure 5.12). The reactor temperature was held constant by controlling the saturated steam pressure in the cooling system.23The height of the catalyst suspension (about 8 m) was maintained at a constant level either by collecting wax using an external, high-pressure filter (high alpha operation) or by adding higherboiling synthesis products to the reactor (low alpha operation). The synthesis gas enters the bottom of the reactor through a gas distributor with jets about 2-3 mm in diameter. The catalyst concentration, because of the micron particle size, was essentially constant from the bottom to the top of the reactor, and the optimum catalyst loading was reported to be about 10 wt% based upon iron. Complete removal of the heat of reaction of 1000 m3 synthesis gas per hour required less than 50 m3 of cooling pipe surface area. The space-time yield of C3+ products was 38.75 kg m-3 reactor volume h- I . The operators of the plant believed that the ultimate yield of a plant running at higher temperatures may be more than 125 kg m-3 reactor volume h-'.24 When operating in the gasoline (low-alpha) mode, it was reported that C3+ products accounted for 85% of the total products. This is surprisingly high

5: Indirect Liquefaction - Where Do We Stand?

151

Gas exit

Slurry

Gas distribution

Gas inlet

Figure 5.12 Reactor for liquid-phase synthesis (Rheinpreussen Process): 1 = reactor cylinder; 2 = cooling pipe register; 3 = liquid level regulator; 4 = steam drum (from ref: 8).

since most reports indicate that 85% C3+ synthesis products are only obtained with high-alpha operations. The work of Kolbel and co-workers was conducted using primarily unsupported iron catalysts. However, Kolbel and Ralek reported that, 'Basically all the catalysts which are suitable for the F-T synthesis can also be used in the liquid phase.' They also reported that, in contrast to other situations, it is desirable for the catalyst particles to break up during use since this leads to a better dispersion of the catalyst. Catalysts with high content of support are less suitable since they lead to unusually high viscosities. The deactivated catalyst recovered following filtration or centrifugation was reported to be regenerated repeatedly simply by oxidation after which the catalyst regained its initial activity. Furthermore, the regenerated catalyst led to the formation of only about half as much carbon as the fresh catalyst, thereby enhancing the catalyst lifetime and its activity. The formation of methane was also decreased with the reactivated catalyst. About 400 kg of hydrocarbon per kg iron was produced up to the time of catalyst regeneration. Because of limited compressor output at the plant, the gas flow for complete synthesis was limited to 3 m3 kg-' Feoh; however, the authors believed that this could be doubled with no problem. Even when operating at the high temperature needed for the gasoline mode of synthesis, less than 4% of the hydrocarbon product was methane plus ethane. This is a surprisingly low amount of methane and ethane. Based upon reported results and their data, Mobil workers25 reported a correlation between the amount of heavy wax and methane that are produced; the lower the heavy wax production the higher the methane make (Figure 5.13). It is noted that the

Catalysis

152

9

-t

60

IK w ( l n e m a n c + - )

A

X

*:

YtF

e

CT-256-7 CT-256 RWIS 1-S &son ct rf.(19!j4}. O t Circulation Chy (1981) Flrley a d Rav (19661 Kunugi ct 3.(ISSa),SIkri t t at- (1974) Schlesinger et al. 119Sll 0

0

2

4

..

6 8 Methane YicM. Wt X HC

1 10

12

Figure 5.13 Reactor wax yield versus methane yield with data from ref: 8 added (from ref: 25).

data point we added to the Mobil plot for the work of Kolbel and coworkers has a very low methane plus ethane yield compared to the work of others. The results obtained in a 6 liter laboratory slurry reactor were in agreement, except for reactor efficiency, with the 10000 liter capacity demonstration plant. Kolbel and Ackermann reported that the formation of higher-molecular weight products is favored by low reaction temperature, a high alkali content and by the chain-extending molecular build-up of the hydrocarbons present in the reactor or recycled into the reactor.26When much of the fraction boiling to 320°C was recycled back to the reactor about 80%of the recycled fraction was involved in molecular enlargement. This appears to be an exceptionally high fraction of incorporation of recycled olefins, based upon our work with adding I4C-labeled alkenes and alcohols to measure directly the extent of reincorporation. 27-37 It appears that the lack of gas chromatography for analysis of the lower molecular weight products and problems with mass balance during recycle of higher carbon number synthesis products caused very large errors in the mass balance for the operations summarized by Kolbel and Ralek. As noted, the ability to obtain at the same time both low methane and low wax production has not been reproduced by other workers. Furthermore, even a native German has, at best, great difficulty in discerning whether the product yields are based on CO+H2 feed or on C O + H 2 converted. The ReinpruessenKoppers work demonstrated the ability to run a large sized slurry reactor. However, the product distribution was not as reliable as desired.

5: Indirect Liquefaction - Where Do We Stand?

153

3.2 British Fuel Research Station - Work was started at the laboratory scale in 1949 and this led to the operation of a large plant with a gas throughput of 70 m3 h-'.38 The slurry reactor was 0.248 m in diameter and 8.5 m tall, with expanded catalyst slurry occupying the lower 6 m of the reactor (276 liter reaction volume). Gas was supplied through a single jet that had a 19 mm diameter. The operators indicated that the iron catalyst employed during much of the work was unsatisfactory. While the operators had solved the operational problems, the catalyst performance was still a problem when funding for the project was terminated in 1961. 3.3 US Bureau of Mines - During the 1940s the Bureau of Mines undertook extensive studies of direct and indirect coal liquefaction with the view of providing a source of transportation fuels that could supplement the anticipated dwindling supply of petr01eu1-n.~~ Surprisingly, these authors dismiss natural gas from consideration, '. .. since the supply of natural gas was subject to essentially the same limitations as crude oil'. The indirect coal liquefaction pilot plant at Louisiana, Missouri began operation in 1949. Much of the work was devoted to making improvements in the gasification process, the major cost associated with transportation fuels production. The reactor (Figure 5.14) operated at 450-525 "F (232-274 "C) and 300-350 psi but could operate to 600 psi. The reactor was loaded with about 7 tons (6356 kg) of iron catalyst placed on top of iron spheres ranging from 3/8 to 3 inch (0.95 to 7.6 cm) in diameter. These spheres supported the catalyst and acted to distribute the feed gas. The syngas and coolant fluid were mixed prior to entering the reactor and the fluid was fed at a rate to limit the temperature rise. Normally, about 50% of the gas entering the reactor was in the dissolved state. The initial catalyst volume [116 ft3 (3285 liter) of 8 to 18 mesh particles] expanded by 20-30% during feed gadcoolant fluid flow. Since coal was expected to be the source of the syngas, work focused on iron catalysts. The catalyst was reduced with hydrogen at high temperatures (up to 1000°C) and thus the precipitated, as well as the fused iron, catalysts had low activity. It was found that the pumps could handle some catalyst fines during circulation between the reactor and external wax recovery unit. Catalyst disintegration occurred in the demonstration plant to a much greater extent than it did in the smaller laboratory reactors. It was concluded that this was a result of the method of addition of the synthesis gas/coolant stream and that modifications to give better dispersion of the feed gas prior to entering the catalyst region of the reactor would have reduced catalyst disintegration. The operations were terminated in 1953 prior to a demonstration of this conclusion. The Bureau of Mines also operated slurry reactors ranging in size up to a 3 inch (0.076 m) diameter, 10 foot (3.05 m) tall reactor of about 15 liters volume that produced about 20 liters of product a day.40After 124 experiments lasting up to 4 months in duration, they concluded that: (1) catalyst concentrations of 50 to 500 grams of iron per liter of slurry are operable; (2) ratios of synthesis gas of 0.7 : 1.3 H2 : CO can be used; (3) pressures of 300 to 400 psi are preferred; (4) the range of space velocities are limited at one end by catalyst settling and

Catalysis

154 ptoducts 4

synthesisgas and cooling oil

=? For activating catalyst

Figure 5.14 Synthesis reactor at Bureau of Mines, Louisiana, Missouri plant (from ref: 39).

,

at the other by frothing; and (5) certain additives aid in keeping catalyst particles suspended. Although many successful runs were made in the slurry reactor, unpredictable erratic behavior was observed in about half of the tests. The erratic behavior usually resulted in low catalytic activity. While the factors determining the low activity were not clearly defined, one factor was due to catalyst settling out of the slurry and this was considered to be the major obstacle to the successful operation of a slurry reactor. While the work at the Bureau of Mines featured iron catalysts, they did also work with cobalt catalysts. They demonstrated that over a range of alpha values, the liquid and solid products obtained in their experiments agreed very well with those predicted from the alpha of the ASF distribution (Figure 5.15). During the 1970s’ energy crisis, Department of Energy (DOE) officials established a bubble column pilot plant at LaPorte, Texas. Initially this plant was used to develop a process to produce methanol, and later, in order to increase productivity, dimethyl ether. This effort was so successful that Air Products and Tennessee-Eastman have recently brought a commercial plant on-stream. In the 1980s, the plant was expanded to include the capability to conduct the Fischer-Tropsch ~ y n t h e s i sThe . ~ ~ facility has a capacity of about 1 ton day- of Fischer-Tropsch products. The reactor is 22.5 inches (0.57 m) in

5: Indirect Liquefaction - Where Do We Stand?

155

Figure 5.15 Correlation of theoretical (curves) and actual (points) product distributions obtained with cobalt Fischer-Tropsch catalysts (from ref: 40).

diameter and is 28.3 feet (8.6 m) tall. The slurry has a height of about 20 feet (6.1 m). The unit, apart from problems with wax-catalyst separation, has performed to expectations during the initial four runs which have included industrial partners together with the DOE. The first two runs used iron catalysts and the last two runs utilized a Shell proprietary catalyst. The first two runs utilized tangential (cross) flow filters and settling ex-situ the reactor to effect wax-slurry separation; this operation was not very successful because of the fines produced by catalyst attrition. Catalyst attrition during the third run generated sufficient fines to plug the proprietary separation system provided by Shell; however, it appears that the problems encountered in separations have been overcome so that the fourth run was completed as scheduled. 3.4 StanolindlCarthage Hydrocol - Work initiated by Dobie Keith at HRI ultimately led to the construction of a commercial-scale plant at Brownsville, Texas. This plant was designed to convert 90 million cubic feet of natural gas daily into 7000 bbl of petroleum products and 300000 pounds of chemicals. The plant was held by nine companies: The Texas Co. owned close to 50% with The Chicago Corp., Forest Oil Corp., Hydrocarbon Research (HRI), La Gloria Corp., Niagara Share Corp., Stone & Weber, Inc., United Gas Corp., and Western Natural Gas Co. holding the remainder. The total investment was about $50 million (early 1950s). The F-T plant utilized circulating fluidized bed technology and an iron

Catalysis

156

catalyst. The plant was the first to introduce large-scale air separation to obtain oxygen for the partial oxidation of natural gas to produce the synthesis gas. Despite many problems with the early operation of the plant, the operators progressed to the point where the plant was technically verified, and the later success of the process at Sasol confirmed this. However, the Hydrocol plant did not run at design capacity for extended lengths of time since the price of petroleum decreased to the point where the F-T product was not competitively priced. The plant was sold to Amoco in the 1950s who operated it as part of their chemicals production and then shut it down.

4

Standard Oil Co. (NewJersey) - Exxon

As indicated above, Exxon (Standard, New Jersey) has been active in the F-T synthesis area for many years. Their most recent activities have produced their Advanced Gas Conversion Technology (AGC-21) process that is based upon the F-T synthesis.42 The AGC-2 1 technology consists of three integrated processes: syngas generation, conversion, and product upgrading. Synthesis gas generation utilizes a novel fluid bed reactor system that combines partial oxidation and steam reforming reactions in a single large reactor that contains a fluidized bed of catalyst particles. The gasification technology builds upon Exxon’s long history of pioneering work in the application of fluid-bed technology, beginning with fluid-bed catalytic cracking in the 1930s. The refractory lined steel gas-generating vessel is operated at conditions which provide high thermal efficiency and offer economy of scale. The F-T synthesis process utilizes an advanced multi-phase reactor and a proprietary high performance catalyst. One of the reactor designs included in Exxon patents is similar to the Sasol Arge reactors except the cooling fluid is on the shell side and the catalyst slurry is contained in multiple tubes that can be operated to approach plug-flow conditions within each reactor tube (Figure 5. 16).43Exxon workers claim to have met the development challenges of heat removal, scaleup of the multi-phase fluid dynamics of the reactor, maintenance of the catalyst performance, and separation of the liquid wax from the catalyst slurry. The third part of the process involves mild hydroisomerization using conventional fixed-bed reactors that contain a proprietary Exxon catalyst. This work was on a massive scale with the cost through 1993 being more than $150 million. During the 1980-1994 period, more than 150 patents were issued to Exxon that cover a range of technology, and a similar pace of patents and publications continues today. The fluidized bed synthesis gas generation (FBSG) unit was constructed during 1989, and the Fischer-Tropsch unit was completed in 1990. The FBSG unit is 120 feet (36.0 m) tall with a reactor vessel that has a 5 inch (0.13 m) diameter. A variety of catalysts, prepared by commercial catalyst manufacturers, were tested as well as gas distributor configuration and materials of construction. During the last two years of a three year test period, the unit produced syngas for the F-T reactor on a sustained and fully

5: Indirect Liquefaction - Where Do We Stand?

157

16

I9

c .

46

Figure 5.16 Schematic of Exxon slurry phase reactor (from ref: 43).

integrated basis. Testing of the F-T hydrocarbon synthesis process was carried out in a reactor with a four foot (1.12 m) diameter. Productivity of the F-T reactor approached 200 bbl day-', a value greater than initially anticipated. This Exxon demonstration unit was at the productivity scale that Rentech has claimed as a commercial operation. Many of the Exxon patented F-T catalysts are based on promoted cobalt supported on titania catalyst, frequently with a second metal in the 15% range to improve and/or retain surface area of the support under reaction conditions. Since the only patent issued to date that provides operating data for this demonstration plant utilized a cobalt on titania catalyst, it is presumed that the proprietary commercial catalyst is some version of this catalyst. To date, the scientific and patent publications by Exxon workers are the most extensive of any group working in the Fischer-Tropsch area. For the work with iron catalysts, Exxon provided data for operations with both fixedbed and slurry (CSTR) reactors. However, with two or three exceptions, Exxon patents on cobalt catalysts present data that were obtained only in fixed-bed reactors.

158

Catalysis

The patents on Fischer-Tropsch synthesis obtained by Exxon could serve as a model for study of patent coverage. There must have been considerable technical/attorney cooperation or the patent attorneys were very well founded in the science and engineering aspects of F-T technology. In addition, the patent process must have been developed following a defined business plan for the commercialization of F-T technology. In many instances the open literature utilizes directly material from patents, but in other instances the open literature extends the coverage into areas not claimed in the patents. In the latter instances, it is presumed that patent coverage was sought but was not attainable. The dominant themes of the Exxon open literature summary papers are the impact of diffusion and of alkene reincorporation upon the rate and product selectivity of the F-T synthesis. The authors divide the effects into a kinetic factor and into factors that are controlled by the physical characteristics of the catalyst and/or reactor. The specific examples apply to cobalt and ruthenium catalysts although it is implied that the conclusions are general and would apply to iron and other catalysts. The theory and data in these papers are generated for the case of fixed bed reactors although it is shown or implied that the results are general rather than limited to fixed-bed reactors. The kinetic factor is very simple: all cobalt sites have the same activity and selectivity with perhaps the exception of Co-Ru supported on titania (Figure 5.17).44 The activity improvement by Re is indicated to be due to an increase in cobalt dispersion whereas the improvement by Ru is accomplished without an increase in cobalt dispersion. Thus, Ru is considered to preserve active sites of cobalt to a greater extent than Re or other promoters during the F-T synthesis. The other factor relates to diffusion effects upon product distribution by allowing for readsorption of olefins and their reincorporation into the synthesis. Exxon workers have developed a reaction-transport model wherein the effect of diffusional restrictions on readsorption rate is controlled by a dimensionless parameter an where Y nreflects the ratio of maximum diffusion rates to maximum reaction rates for olefins of carbon number, n, with the rate constant for first-order olefin readsorption turnover rate, kr,n,and the effective diffusivity, D,, of aolefins within the support pellets. The term Y nreflects the molecular properties of olefins of a given size and accounts for the observed effect of a rapid decrease with carbon number of the chain termination probability (Figure 5.18).44 The other term, X, includes the pellet radius, R,, the support void fraction, E , the Co sites per unit area, @co, and the average pore radius, rp. This term involves the structural properties of the support and the number of surface Co atoms that form and readsorb olefins. Diffusion-enhanced readsorption and incorporation of olefins increase with increasing carbon numbers up to about carbon number 15. Beyond carbon number 15, diffusion-inhibited chain growth impacts the product distribution. Thus, for small values of X the

5: Indirect Liquefaction - Where Do We Stand?

159

,' CO-RU

26

w

0

om

0.04

om

0.00

at

0.e

0.u

Cobalt Fractional Dispersion Figure 5.17 Effect of cobalt dispersion (ratio of surface Co to total Co atoms), support and alloying on FTS Co-time yields (mole CO convertedltotal g-atom Co-s). Reaction conditions: 473 K, 2000 kPa, H2 :CO = 2.05, Cf selectivity > 80%, 0.17 m m pellet size; A,titania; a, silica; m, alumina; A, titania; 0, silica (from ref 44).

reincorporation of olefins enhances the C5+ production up to a maximum after which further increases cause a decline in C5+ due to the diffussioninhibition of chain growth (Figure 5.19).44 The reverse trend is obtained for methane production. Exxon has an extensive number of patents covering F-T and their associated processes. Some of these patents make claims that are obvious whereas other patents claim areas where the novel advance or improvement is apparently obvious only to those with expert knowledge in patent law. Many of the earlier catalysts covered a variety of iron-based catalysts and the high productivity of these catalysts was demonstrated in both fixed-bed and stirred tank reactors. With two or three exceptions, patents covering cobalt catalysts provide activity data obtained only for fixed-bed reactors. The patent coverage for cobalt catalysts includes preparation, activation, reactor configurations, catalyst rejuvenation and processes for their use. In conformity with the theoretical developments, they have covered rim-loaded cobalt catalysts wherein techniques are claimed for the deposition of cobalt on the exterior rim of a support pellet, thereby limiting and/or controlling the diffusional effects. Patents cover catalyst activation in the reactor using fluid-bed techniques and then adding the start-up solvent. Another patent covers the reduction of a fresh catalyst with hydrogen or a hydrogen-containing gas in the presence of hydrocarbon liquids. A patent covers a 'super' activation treatment wherein the catalyst is initially activated in a fluid bed, passivated and then reactivated in a slurry

Catalysis

160 02

I

1

1

5

I

a n e a

0.15

C

= 0

c-

C

at

FF

0.0s

c 0

Figure 5.18 Bed residence time and carbon number effects on chain termination probabilities to olefins and paraffins (9.5% CO conversion, 2 s bed residence time; ColTiO2, 1 1.7% Co, 0.0 15 dispersion, 0.17 mm pellet size, 473 K,2000 kPa, H2: C0=2.1) (from ref: 44).

Figure 5.19 The effect of structural parameters (x) on FTS selectivity. Diffusionenhanced readsorption (----) and diffusion-inhibited chain growth (-) simulations and experimental data; a, dispersionisupport effects; A,pellet size variations; 0, eggshell thickness variations; (a) Cg+ selectivity; (b) CH4 selectivity (473 K, 2000 k h , HZ:CO=2.1, 5 5 4 5 % CO conversion) ( x values from eqn. 2 of ref: 44 with rp and R, in m, and Oco in surface Co atoms m-2) (from ref: 44).

5: Indirect Liquefaction - Where Do We Stand?

161

reactor. From a scientific point-of-view, it is difficult to appreciate the improvement of this technique over that of the ‘ROR’process patented by Gulf. Patented techniques for catalyst rejuvenation (regeneration) have been obtained that permit the rejuvenation to be carried out during reactor o p e r a t i ~ n .Another ~ ~ ? ~ ~patent also covers the rejuvenation of a catalyst using an external The operation of the slurry reactor pilot plant at Baton Rouge is described in patents; other patents cover various reactor configurations. K ~ r o des ~ ~ scribes a reactor in which the coolant is on the shell side of the reactor and the catalyst is dispersed in the slurry phase within small diameter tubes, each operated to approach plug-flow, in a configuration that resembles the Arge reactor except that the fixed catalyst bed is replaced by slurry (Figure 5.16).43 Benefits are also claimed for the operation of a reactor that includes one or more vertical down corner^.^^ Kim and fiat^^^ claim a processing scheme in which the synthesis gas contacts an olefin-producing catalyst (e.g. iron) so that conversion is limited and then the remaining syngas and highly olefinic hydrocarbon stream contacts a catalyst (e.g. cobalt) that will effect incorporation of olefins into the product. Obviously, space limits the coverage to only a few of the many patent claims, and those included are biased by the writer’s interest and understanding.

5

Pilot Scale Operations

Syntroleum - Syntroleum has been especially active in the business aspect of Fischer-Tropsch commercialization and they appear to have been a major factor in the renewed interest in this process. The company’s work, started in 1984, includes the operation of a 2 bbl day-’ pilot plant beginning in Tulsa in 1990 and a second-generation of unspecified type reactor that was introduced in 1996. They have announced, at times, fixed-bed, fluid-bed, HMX (hybrid multiphase technology) and horizontal reactors. It has been reported that a fluid-bed reactor utilizes a high-alpha cobalt catalyst that operates at 300-500 psig and 375-450°F (191-232°C) and has been operated in the 2 bbl day-’ pilot plant?* There have been several reports that ‘product yields and quality [have been] confirmed’ at the pilot plant scale but the basis of the confirmation are not reported. Syntroleum utilizes air rather than oxygen in an autothermal reforming step that employs a catalyst and a reactor of proprietary design. They do not recycle an appreciable amount of gases to the reformer. They claim that the reactor is simple enough so that it does not need large scale to be cost effective.s’ Work on chain limiting catalysts began in 1994, with partial funding from three oil companies.52Agee reports that, ‘Recent [1997] multiweek test runs in a fluid-bed reactor at the pilot plant yielded a product profile that indicates success. This catalyst promises several additional efficiencies to the process configuration, including: a lower operating pressure for the process, the use of higher capacity fluidized-bed reactors that cannot be 5.1

Catalysis

162 PRODUCT COMPOSITION, Xw

I O r

0

10

20

30

4

CARBON NUM8ER

Figure 5.20 The data are for Shell process (hydrocracking product) and the line is based on data taken by B. Davis from slide shown during presentation by M. Agee.

effectively used with the high-alpha, wax-producing catalyst, and elimination of a hydrocracking step.’52A plot of their reported chain-limiting product distribution, compared to Shell FT wax and the product produced with heavyhydrocracking, is provided in Figure 5.20. They also offer a proprietary, highalpha, highly-active cobalt catalyst with relatively low yields of methane (below 10%). They report that, ‘Pilot test runs with commercially manufactured batches have demonstrated the viability of the high-alpha catalyst system.’52 The product distribution reported by Syntroleum is, at first glance, indeed unique. Based upon the high-temperature F-T products in Table 5.1, Sasol produces products that meet the cut-off requirement shown for the Syntroleum data in Figure 5.20. However, there is a major difference between the lower carbon number yields reported by Syntroleum and by Sasol. The Syntroleum data show a methane product that is about 6 wt% of the hydrocarbon products, in reasonable agreement with Sasol data. However, Syntroleum shows the amount of C2-C4 products decreasing with carbon number to about 1.5 wt% for C4; this is in marked contrast to the Sasol data where C2-C4 products account for 30 wt% of the product and for Syntroleum this is about 8 wt%. One way Syntroleum could obtain such a distribution would be to oligomerize the 20 wt% C2-C4 olefins to produce a C10-C20hydrocarbon fraction. Mossgass does have a process that utilizes a silicalite catalyst to produce alkylate from the C2-C4 olefins; however, the oligomerization step is conducted in a separate reactor and at temperatures that are different from the F-T synthesis. Theoretically, it is possible to obtain a product distribution that

5: Indirect Liquefaction - Where Do We Stand?

163

Table 5.3 Agreements reported by Syntroleum Company

Date of agrmt. Specific agreement

Ref:

TEXACO

12130196

Nonexclusive ‘Master’ Licence Joint work to accelerate development of Syntroleum Process Build 2500 bbl day-’ plant outside US, site to be released 1st quarter 1998, start-up Fall, 1999

53 54

1 2110197

MARATHON

31 10197

Non-exclusive ‘Master’ Licence Marathon provides Syntroleum nonexclusive rights to proprietary technologies

55 56

ARC0

4110197

Non-exclusive ‘Master’ Licence with rights to use in broad geographic areas, including Alaska Joint development; Arc0 to construct 70 bbl day-’ pilot plant, start-up Fall, 1998 to demonstrate new reactor design

57 58

10124197

ENRON

02/24/98

Agreement to develop 8000 bbl day-’ speciality products plant, Wyoming; construction begins late 1998, operational in 2001 Non-exclusive volume license for Enron to license Syntroleum Process outside North America

59

YFP

08/04/97

Non-exclusive license agreement, rights to use Syntroleum Process in broad geographical locations outside North America

60

KERR-McGEE

02/05/98

Non-exclusive license agreement for KerrMcGee to use Syntroleum Process outside North America Syntroleum receive certain hydrocarbon processing technology and related patents from Kerr-McGee

61

resembles that shown for the Syntroleum data in Figure 5.20 but it requires two catalysts and, most likely, two reactors. Syntroleum routinely refers to a chain-limiting catalyst, implying that two catalyst formulations are not used and their simple process scheme does not indicate a product upgrading step. Syntroleum has announced agreements with several producing companies, and include Texaco, Marathon, Arco, Enron (world’s leading integrated natural gas and electricity company and owner of $23 billion energy-related assets), YFP (a wholly owned subsidiary of YFP Sociedad Anonima, Argentina’s largest oil company), and Kerr-McGee. These agreements are reported as non-exclusive ‘Master’ license agreements or non-exclusive license agreements. A synopsis of these arrangements are outlined in Table 5.3. The Arco-Syntroleum joint development project will include the construc-

164

Catalysis

tion of a 70 bbl day-’ pilot plant at Arco’s Cherry Point Refinery near Bellengham, Washington. ARC0 announced the completion and successful start-up of the pilot plant on July 28, 1999. In addition, they announced the successful integration of a new catalyst system and the advanced reactor design. This operation will demonstrate a new reactor design and will be used to produce synfuels ranging from diesel to heavy waxes. It was announced that Texaco intends to p a r t i ~ i p a t eThe . ~ ~ type of reactor was not announced but one would presume that it would involve the HMX hybrid multiphase technology. Arc0 does not appear to have made public their plans other than arrangements with Syntroleum. Enron and Syntroleum contributed $3 million (share supplied by each company was not announced) to fund a detailed engineering study, the purchase of land and other development costs for an 8000 bbl day-’ gas-toliquids (GTL) specialty product plant in Sweetwater County, Wyoming. Construction was scheduled to begin in late 1998 and the plant will be operational in 2001. The products will include synthetic lubricants, drilling fluids and liquid n-paraffins and the plant will produce 50 MW day-’ electricity for sale. Bateman Engineering is the engineering contractor for the project. In addition, a non-exclusive license was granted to allow Enron affiliates to acquire the right to license the Syntroleum Process outside of North America to produce liquid fuels. It is unlikely that the Sweetwater plant will be built in the near future. Because of more favorable gas prices, Syntroleum is now considering sites in northwestern Australia and South America for this plant. Final negotiations are expected to be completed in late 1999 (private communication). The Sweetwater plant has now been announced to be located in Australia. Criterion Catalyst Company and Syntroleum formed an alliance in which Criterion will manufacture and supply Syntroleum proprietary catalysts, presumably to those companies who operate plants using the Syntroleum Process. Syntroleum also obtained from Lyondell Petrochemical Company the right to license and sub-lease a new synthetic wax isomerization process based on Lyondell’s catalytic dewaxing process. The Lyondell dewaxing process was developed by Lyondell, Criterion Catalysts Co., and Zeolyst Int. Syntroleum has agreements with a number of engineering companies. Apparently the first of these agreements was with Brown & Root, a Colorado-based engineering company. In February, 1997, Brown & Root, a subsidiary of Halliburton Co., and Bateman Engineering, based in Denver, Colorado, were authorized to represent, market and license the Syntroleum Process to approved third parties for the production of synthetic fuels from natural gas.62 Shortly thereafter a project development agreement was announced for Bateman to be part of Syntroleum’s plan to build a series of ‘natural gas refineries’ in North and South America for production of synthetic lubricants, solvents and chemical feedstocks. In late 1997, it was announced that three firms (Brown & Root, Bateman and AMEC Process and Energy Limited), were named ‘Approved Process Design Provider’ to Syn troleum.

5: Indirect Liquefaction - Where Do We Stand?

165

b m m ROD vita Figure 5.21 (Left) Schematic of Sasol Synthol reactor" and (right) schematic of Syntroleum chain-limiting reactor (from reJ 63).

A recent international patent application, assigned to Syntroleum, describes a gas conversion process using a chain-limiting reactor.63 This patent teaches that, ". . . syngas is passed through a fluidized-bed catalyst in a chain-limiting reactor until the chain length of the hydrocarbon being absorbed [adsorbed?] on said catalyst reaches a chain length which is substantially equal to the chain length of the desired hydrocarbon product. This saturated portion of the catalyst is then removed from the reactor and is replaced with fresh catalyst. The 'saturated' catalyst is regenerated and recycled to the reactor." This patent does not contain examples with experimental data. This is both surprising and disappointing since Syntroleum claimed, in March 1997, to have operated a pilot plant in this chain-limiting mode during multi-week test runs in a fluid-bed reactor. The reactor described in the patent application does not appear to be unique from the circulating fluidized-bed reactor design used for years at Sasol (Figure 5.21). Fresh catalyst is added at the bottom of the reactor, contacts syngas as it rises in the reactor and the length of the chain of the hydrocarbon being absorbed on the catalyst continues to grow. The retention time of the catalyst in the reactor is such that the hydrocarbon chain on the catalyst will grow as the catalyst moves toward the top of the reactor and will attain the maximum desired length as the catalyst reaches the takeout point. It would appear that the gas flow-rate must be sufficient for the reactor to operate with the catalyst particles in, or nearly in, a plug-flow mode. This will require high gas flows and tall reactors. If the catalyst bed is well-mixed,

166

Catalysis

then catalyst particles throughout the bed will have the same average hydrocarbon composition. In the abstract, patent body, and the claims, the author uses absorb, rather than adsorb. As used in chemistrykhemical engineering terminology, adsorption means that the compound is present on the surface of the catalyst where absorbed means that the compound is taken into the bulk. It is not possible for the hydrocarbon product to become a part of the bulk of the cobalt or support; thus, the only way that absorb can have meaning in the context of the patent is that the hydrocarbon is retained within the catalyst porosity. Likewise, the patent does not define ‘saturated’. It appears that there is a contradiction within the description of the claimed operation of the reactor. Claim 1 is ‘A Fischer-Tropsch process carried out in a chain-limiting reactor for converting syngas to a desired hydrocarbon product having a desired chain length, said process comprising: passing said syngas through a fluidized bed of catalyst in said chain-limiting reactor to convert said syngas to said hydrocarbon product; operating said reactor at temperature and pressure at which said hydrocarbon product remains in its vapor phase while in said reactor; removing said hydrocarbon product while in its vapor phase in said reactor; continuously removing catalyst from said chain-limiting reactor as the chain length of any hydrocarbon product absorbed on said catalyst reaches said desired chain length for said hydrocarbon product; and continuously supplying fresh catalyst in said chain-limiting reactor at the same rate as that at which said catalyst is removed.’ The authors claim that the hydrocarbon product remains in the vapor phase while in the reactor. If, as claimed, the hydrocarbon product is in the vapor phase in the reactor, there will therefore be no ‘absorbed’ hydrocarbon on the catalyst. It will therefore be impossible to remove catalyst at the top of the catalyst bed that contains ‘absorbed’ hydrocarbon of the desired chain length. Even if it is granted that some of the hydrocarbon is ‘absorbed’ with the catalyst, the product using a typical F-T catalyst must follow the typical Anderson-Schulz-Flory distribution. This is illustrated in the following. Take the simpler case of plug-flow of catalyst particles with the residence time adequate to make the longest product a C18 hydrocarbon. Chain termination will occur for some of the catalyst particles as they travel the length of the catalyst bed. Furthermore, the chain termination probability will be that which gives the ASF distribution. For example, assume that a catalyst particle that enters the reactor has a chain growing that is terminated at methane (Cl). When this happens, there will be a free site on the catalyst where CO will adsorb, initiate a new chain that will grow as the catalyst particle passes along the catalyst bed, and the second chain will follow the ASF plot. The only way that the products using the chain-limiting reactor can deviate from the ASF plot that would be obtained in a typical reactor is if a catalyst is utilized that has a termination probability of zero. Stated another

5: Indirect Liquefaction - Where Do We Stand?

167

c

Yn m

a

lo-7

2 4

6

(I

10 I2 H

16 I 6 2 0 2 2 2 4

n Figure 5.22 Effect of time on the distribution of Fischer-Tropsch products obtained from a well-stirred slurry reactor when the catalyst is characterized by a single value of a (V,= 15 cm3; V G = 300 cm3; Go=400 STP cm3 min-'; Mw, = 350 g mol- I; PL = 0.7 g ~ r n - ~a= ; 0.7) (from ref: 64).

way, there will be vapor-liquid (vapor-absorbed phase) equilibrium for each carbon number hydrocarbon. Mass balance requires the hydrocarbon product of each carbon number leaving the reactor to be the sum of the vapor and adsorbed product that is produced by the ASF mechanism. Whether the product leaves the reactor in the vapor or absorbed phase is irrelevant as far as the mass balanced product distribution is concerned. The only difference that the reactor could make, whether operated in the backmixed or plug-flow mode, is that which is well defined in any reactor text where at intermediate conversion levels, a higher concentration of an intermediate product (e.g. alkenes in this case) could be higher in the plug-flow reactor than in the backmixed reactor. It appears that the ASF products will not be formed only if the catalyst residence time is sufficiently short or if appreciable hydrocracking occurs. We conclude that the typical Fischer-Tropsch catalyst will produce an ASF product distribution in any reactor; any deviation from this distribution will be due to obtaining product analyses that do not represent a proper mass balance. For example, one can easily obtain a higher fraction of lower hydrocarbons (chain-limiting operation) due to accumulation of heavy hydrocarbon products in the reactor; this has been illustrated by, for example, Dictor and BellM (Figure 5.22). Data obtained at early times on-stream in the runs conducted by Dictor and Bell readily fit the definition of chain-limiting

Catalysis

168

---- -- -

Chaln Llmltlng Hlgh Alpha

Figure 5.23 Product distribution of chain limiting and high alpha product catalyst (redrawn from a Syntroleum handout).

reactor but this is true only if one ignores the heavier products that are temporarily retained within the reactor. The Syntroleum patent application clearly teaches that the reactor is operated so that the hydrocarbon products are in the vapor phase. This requires that the catalyst used in the chain-limiting reactor has a chain termination probability that is greater than zero. It is therefore concluded that the reactors described in the patent cannot be operated to effect an operation with a chain-limiting hydrocarbon product distribution in the sense of deviating from the ASF distribution or that the patent application does not reveal sufficient information for one versed in the art to be able to reproduce the art described in the patent. The products from a Syntroleum chain limiting catalyst are represented in Figure 5.23 together with their representation of a normal high-alpha catalyst. The limitation of the high molecular weight products for the Syntroleum catalyst is typical of a low-alpha catalyst. However, there is a vast difference between the products shown for the Syntroleum chain-limiting catalyst and the products produced by a conventional low-alpha catalyst. Whereas the volume percentage of products decrease from C1 to about C4 for the Syntroleum catalyst before they begin to increase, the volume percent produced by a typical low-alpha catalyst would increase from C1 to about Cq. For example, an average molecular weight of C4 would be obtained for a catalyst with an alpha of about 0.7. At Mossgas, a silicalite catalyst is utilized to carry out oligomerization of the low molecular weight products to produce gasoline and diesel transportation products. Thus, it is anticipated that the Mossgas products, when the alkenes have been converted to liquid transportation fuels,

5: Indirect Liquefaction - Where Do We Stand?

169

will resemble the product distribution shown for the Syntroleum chain limiting catalyst. A problem that one encounters is the definition of a chain-limiting catalyst. To most investigators, a chain limiting catalyst is one that deviates from the normal Anderson-Schulz-Flory distribution. Furthermore, the distribution deviates negatively from the ASF at some carbon number range of products. In other words, the products follow the ASF distribution up to some carbon number, but above this carbon number the products are present at a lower value than that corresponding to ASF, or are zero. This distribution has also been designated as being produced by a catalyst with a ‘cut-off’. It appears that the groups currently using chain-limiting catalysts do not mean a deviation from the ASF distribution. Rather, it appears that a chain-limiting catalyst, as used by Syntroleum and Carbon Resources, is one that produces a product distribution whose vapor pressure is such that essentially all products are in the vapor phase at the reaction temperature. This type of chain-limiting catalyst has been utilized at the commercial scale for about 50 years. For the circulating fluid-bed operation, liquid phase products cannot be tolerated. The presence of liquid causes the catalyst particles to agglomerate and plug up the reactor because they become too large for the gas to maintain them in a fluid phase. Thus, catalyst agglomeration was one of the early problems encountered by the operators of the Brownsville, Texas and the Sasol fluid-bed reactors. By adjustment of process conditions and catalyst compositions, these plants were able to operate their fluid-bed reactors in the absence of a liquid phase. At Brownsville and Sasol, the catalyst was chain-limiting because it produced such a small amount of high molecular weight product. However, the products follow the distribution expected for a ‘normal low-alpha’ catalyst that follows the normal ASF mechanism. Thus, on the basis of limiting the amount of high molecular weight liquid products, Sasol has operated reactors that contain chain-limiting catalysts, and the output from these catalysts accounts for about 90% of their production (that is, 90000+ barrels per day). The distribution shown by Syntroleum is that expected for the Sasol catalyst with the exception of a different distribution at the lower molecular weight products. The lower molecular weight products represented by the Syntroleum curve in Figure 5.23 deviate significantly from that of the Sasol operation, but the higher molecular weight products do not. The higher molecular weight products of the Syntroleum distribution curve are expected if the catalyst produces products representative of a low (e.g. 0.70) alpha catalyst. Furthermore, the Syntroleum curve is consistent with the conversion of the low molecular weight alkene products to liquid products in the range of about CgCI7.This type of process is apparently operated by Mossgas where they use a pentasil catalyst to oligomerize low molecular weight alkenes and, while they may not produce a product distribution curve that is identical to the one shown in Figure 5.23 to represent the Syntroleum products, the Mossgas product distribution is expected to resemble the one shown for the Syntroleum products. Salomon Brothers analyst Paul Ting asserts that, ‘They [Syntroleum] are

170

Catalysis

creating the gas-industry equivalent of the PC standard’. If Mr Ting refers to business operations with the creation of publicity and obtaining agreements for potential commercial operations, Syntroleum indeed is creating a ‘PC standard’. If, on the other hand, Mr Ting is referring to a definition of the scientific and technical aspects of Fischer-Tropsch synthesis, he is certainly incorrect since Exxon and Shell, for example, have provided a much more complete and detailed description of Fischer-Tropsch synthesis in both their scientific publications and in their patents. 5.2 Gulf Oil - The technology developed by Gulf was transferred to Chevron following a merger of the two companies. Later, Chevron transferredsold the technology rights to Shell and Shell appears to have incorporated much of the Gulf technology into the commercial plant in Malaysia. Gulf workers prepared a series of catalysts that demonstrated a beneficial effect of a minor amount of ruthenium on a catalyst containing a major amount of cobalt for the low pressure synthesis of higher hydrocarbons from synthesis gas.65 It was claimed that the incorporation of ruthenium increased not only the activity but also the average molecular weight of the product. However, problems with the analysis of products were probably responsible for the heavier products obtained with the catalyst containing ruthenium. Gulf workers found that a catalyst that was subjected to an initial reduction followed by an oxidation and then a second reduction (ROR process) yielded a catalyst with a higher activity than when the same starting material was activated using only the same initial reduction procedure.66 Gulf workers also patented a two-stage process67 wherein the first stage contained their cobalt catalyst, promoted with a Group IIIB or IVB metal oxide and ruthenium, and was used to generate Fischer-Tropsch synthesis products. The C5-C9 fraction of the F-T products was then converted to high octane gasoline using a typical naphtha reforming catalyst. While it was stated that the reactor type utilized was not a critical factor, a fixed-bed, with gas downflow, was preferred. Another variation of the two-stage process was also patented68 wherein the Fischer-Tropsch products were converted to highly aromatic high-octane gasoline range products using a silicalite ZSM-5 type catalyst in the second stage rather than the naphtha reforming catalyst. Gulf workers demonstrated their process and their promoted catalyst at the commercial scale of operation. With the fixed bed reactor, one must use small diameter tubes in order to effect sufficient heat transfer to maintain reasonable temperatures along the length of the reactor tube. Thus, Gulf and Badger operated a 40 foot tube reactor (two 20 foot sections in series) for extended periods of time and obtained data under conditions where the exotherm was less than 5 O F (2.8 “C) along the reactor (Figure 5.24).69These data verified the Gulf-Badger process and catalyst at the commercial level since a bundle of identical tubes would be utilized in a commercial plant. The takeover of Gulf by Chevron eliminated the urgency to commercialize the Gulf-Badger process, and the rights to it were eventually transferred to Shell.

5: Indirect Liquefaction - Where Do We Stand?

171

COOLINQ WATER

c

L

-3

MIXER

[

QAS OUT

t

O A S OUT

Figure 5.24 Schematic diagram of the pilot plant remote-gas-to-dieseljixed-bed process (fromre$ 69).

5.3 Rentech - As noted above, Rentech reports that they have operated during 1982- 1983 a commercial-scale gas-to-liquids facility. This facility utilized a Pueblo, Colorado landfill to produce gas sufficient for a plant with a capacity of about 235 bbl day-' of liquid hydrocarbons. This plant has a capacity that is about the same as the reactor for the Exxon demonstration plant. It was soon learned that the landfill was unable to produce sufficient gas and the operation was discontinued within three months of the attainment of full-scale product ion. Rentech has nine patents to cover their technology. The research topics for all of these patents are either the same or very similar and provide a concise summary of Fischer-Tropsch synthesis that can serve as excellent introductory reading for the novice in this area. The claims are more modest in coverage than the body of the patents. Furthermore, it is sometimes difficult to discern when the cited examples are based on theoretical calculations and when they are based upon data obtained by Rentech. Rentech has operated a 6 inch (0.15 m) diameter, 8 foot (2.54 m) tall slurry reactor with an iron catalyst. They report that they have designed, built and tested two 6 foot (1 -83 m) diameter slurry reactors and have produced catalyst batches of more than 7.5 tons (6810 kg) for use in these reactor^.^' The waxslurry separation, as reported in 1992, was accomplished using an external

I72

Catalysis

cross-flow filter. Using this technique, separation was still a problem. They apparently utilized later in their pilot plant studies an external separator that involved gravity settling to concentrate a catalyst-wax slurry for return to the bottom of the reactor; this approach was apparently very similar in design to the one Mobil Oil utilized in the work they conducted for their DOE contracts. On June 16,1999 an agreement was signed to study, during the next 8 months, the integration of Rentech's Fischer-Tropsch and Texaco's gasification processes. The research will concentrate on the use of petroleum coke, heavy resid and/or coal to generate synthesis gas. Rentech's iron catalyst technology will be utilized. Rentech also signed a memorandum of understanding for BC Projectos, Ltd., a Brazil engineering company, to be Rentech's engineering firm of record. Rentech is also trying to interest owners of existing methanol plants, many operating below capacity or moth-balled because of low methanol prices, to add the Fischer-Tropsch plants to utilize existing gasification capacity. Rentech's 280 bbl day-' plant was transferred to Donyi Polo Petrochemicals and shipped from Pueblo, Colorado to Arunachal Pradesh, India. The plant will process natural gas from the Kunchai field that is currently being flared. The plant is expected to be commissioned in 2000. United Catalysts, Inc. (India) will manufacture the catalyst. Chinese Studies - The Chinese government is supporting several projects aimed at developing technology, including Fischer-Tropsch synthesis, to convert coal to transportation fuels. The Institute of Coal Chemistry, Chinese Academy of Science, located in Taiyuan, is responsible for the large-scale reactor studies. During the past few years these workers have completed a number of studies directed toward understanding of the science and engineering aspects of FTS and with the operation of larger pilot plants. They have utilized fixed-bed tubular (50 x 4500 mm) reactors with a total capacity of 2-5 litres with iron catalysts. A pilot plant of 100 ton year-' (25 x 4000 mm multitube) and a demonstration plant of 2000 ton year- (32 x 7000 mm multi-tube) have been tested. The F-T reactor was used as the first stage with the second stage being a ZSM-5 catalytic process designed to convert the F-T products to gasoline range transportation fuels. They also have built and tested a slurry bubble column reactor (40 x 4500 mm reactor). This reactor has been utilized with an iron-based catalyst and has been employed in conjunction with the ZSM-5 catalytic second stage process. Currently, large fixed-bed reactors are being utilized in studies with ultrafine iron catalysts, magnesium-iron catalysts, and supercritical operation with an iron ~ a t a l y s t . ~ ' 5.4

'

Mobil Oil - During the 1980s, Mobil Oil workers constructed and operated, with DOE funding, a slurry bubble column reactor as the first stage and Mobil's ZSM-5 zeolite catalytic conversion of the F-T products to gasoline range fuels as the second stage.25The slurry bubble column reactor was 5 cm (i.d.) x 7.6 m and in normal operation utilized about 1.5 kg of catalyst. The reactor utilized an external wax filter assembly (Figure 5.25). Thirteen runs were made, the first three with a low-alpha iron catalyst and 5.5

5: Indirect Liquefaction - Where Do We Stand?

173

rJ

-06-

V-6

v-t

B Ib

To Wax Recefver

Figure 5.25 Schematic of the external wax$lter assembly (from ref 25).

switching to the high-alpha mode during the third run and the runs thereafter. During the course of these runs, many operational problems were encountered. The diameter of the reactor is at, or perhaps below, the diameter now considered to be free of significant wall-effects which impact the measured kinetics. In spite of any operational problems, these studies provide the most detailed open-literature data for slurry phase operations and have been the basis of all, or nearly all, of the economic and process studies conducted by DOE contractors, such as Bechtel and Mitre. In most instances, the F-T data used in these evaluations are from the last Mobil run using a high-alpha iron catalyst. A summary of these results is given by Fox and Tam.72 'Mobil gave detailed product breakdowns which exhibit a characteristic break in the Schulz-Flory relationship at a carbon number between 20 and 25. In addition, methane production is several times what it would be if the correlation line for carbon numbers 2 to 20 were simply extrapolated back to methane . . . . The a values are constant over the specified range of carbon numbers and the summed weight fractions for each range are adjusted such that the total yield adds up to 1.0. There is a sharp break in the slope of the product distribution curve at C20-C25.Vapor-liquid equilibrium predicts that the C22 component is split almost evenly between vapor and liquid phases. Since the liquid phase remains in the reactor much longer than the vapor phase, it is believed that the longer liquid residence time is responsible for the sharp break in slope, at about C22,to a higher value of a.'

Catalysis

174

I

I

I 1

I I

I I

I

0.1

-- 7

:::56

0 . 01

w f / n 0.0 101

9.49

1

0.00 101 0.000 01 0.0000 10 1

0

10

I

1

I

1

I

I

I

I

20

30

40

50

60

Carbon Number, n

Figure 5.26 Design Schulz-Flory plots based on Mobil Oil data (from ref: 72).

Fox and Tam conclude that for Mobil's F-T catalyst system and reactor configuration, setting the operating temperature fixes the entire carbon number distribution (Figure 5.26). The Mobil product distribution data stand in sharp contrast to the ones reported by Kolbel and Ralek for the operation of the Rheinpruessen-Koppers plant. Whereas very low methane and wax yields were reported for the German plant, Mobil obtained a higher methane make than expected from the ASF distribution. With an iron catalyst the water-gas shift reaction can make an important contribution to the overall consumption of CO. The Mobil operation should be viewed to be representative of processing where the WGS is approaching its equilibrium value. Thus, Mobil has reported WGS ratios that are typically 25 whereas the value at equilibrium is expected to be in the range of 60.

6

Process and Economic Evaluations

The following will present typical reports rather than attempt to be comprehensive. G r e g ~ summarizes r~~ the economic advantage of the F-T synthesis that is utilized within a petroleum refinery. He concludes that the F-T distillate fuels have exceptional properties but that the F-T naphtha is difficult to reform into high octane gasoline. When F-T is viewed as a supplement rather than an alternative to petroleum, the refiner would likely utilize the lighter F-T products in petrochemical applications rather than fuel. The lighter F-T products have been utilized as gasoline fuel in South Africa but until now they have been able to add lead to boost octane number and therefore do not have to depend on naphtha reforming alone to boost octane to acceptable levels. As more and more F-T plants come on-stream, the ability to utilize the lighter fractions as petrochemical feedstocks will become plagued with overcapacity.

5: Indirect Liquefaction - Where Do We Stand?

175

Most of the process and economic assessments carried out by the US DOE have been based on coal as the feedstock. Only recently have these DOE studies been based upon natural gas.74 In the past few years, the DOE effort has concentrated upon producing F-T products within an electric power generating plant or as a plant that co-produces power. Lange and Tijm75have considered many schemes for converting methane to liquid hydrocarbon fuels. They conclude that F-T transportation fuels are not competitive at oil prices less than U S 2 0 per bbl. These authors state that while natural gas offers the advantage of a lower cost feedstock, compared to petroleum, the higher capital cost associated with GTL offsets the feedstock cost and makes GTL unattractive at prices less than about $20 per bbl. They indicate that the analyses of numerous fuel manufacturing plants show that the capital cost correlates with the energy loss as well as with the heat and momentum transfer duty of the process and process segments. These authors conclude, perhaps not surprisingly, that when the various methane conversion schemes are compared, the Shell SMDS process is the cheapest scheme proposed at that time (1996). The authors found that a direct scale-up to a 50000 bbl day-' plant from that of a 10000 bbl day-' plant results in a reduction of capital cost sufficient to lower the cost by about $5 per bbl. However, formidable difficulties are faced in increasing the size of plant components by this five-fold amount. G r a d a ~ s i ,based ~ ~ on his selection of recent literature, concluded that a Fischer-Tropsch plant can generate close to $6 billion in cash flow over the lifetime of the plant, and would have product sales revenue of $200-400 million per year. However, the initial investment will be at risk for at least six years before revenues have returned capital cost, taxes and operating expenses.

7

Potential Commercial Operations

Currently, the interest in monetizing remote natural gas is very high. In some instances this interest is driven by environmental considerations. Methane is at least 30 times as potent for developing the 'greenhouse effect' as carbon dioxide is. Thus, there is the expectation that very severe restrictions will be developed throughout the world that will limit the release of natural gas during recovery of petroleum crude. Even today, some regions require measures such as reinjection of natural gas that is obtained in association with crude. Limitations on carbon dioxide release will also place restrictions on flaring of associated natural gas. In addition, producing countries are beginning to place charges on the total carbon recovered, and not just the petroleum crude. Syngas generation, even with a natural gas feedstock, comprises about 50% of the cost of a Fischer-Tropsch operation. There has been a renewed interest in adapting the rapidly advancing membrane technology to the production of syngas at a much reduced cost compared to partial oxidation or steam reforming. Recently it was announced that Phillips Petroleum would join a syngas alliance77made up of Sasol, Amoco, British Petroleum, Praxair and

176

Catalysis

Statoil. Each of the members will contribute technical expertise, R&D and funding to the project. The original alliance was announced7*to 'expeditiously and aggressively move this [membrane] technology toward commercialization'. The original alliance companies announced that they have enlisted support from universities and government laboratories and would consider other companies or groups. DOE has announced an eight year project to develop technology to convert natural gas to liquid for transport through conventional pipelines.79 Air Products' proposal was chosen to develop ceramic membranes that could sharply reduce the cost of generating synthesis gas. In the first two years, members will develop membrane materials, catalysts and other key components. During the next three years the technology will be scaled up to a 12 Mcfd (thousand cubic feet per day) experimental unit. During the final two year phase the technology will be scaled to 15 MMcfd (million cubic feet per day) precommercial unit at Air Products' industrial gas complex at LaPorte, Texas - the location of the DOE slurry phase F-T reactor. In addition to Air Products, the team consists of Babcock & Wilcox, Cerametec, Eltron Research, ARCO, Argonne National Lab., Pacific Northwest National Lab., Penn State University and the University of Pennsylvania. Several of the companies currently looking to develop commercial FischerTropsch plants are not the well-known petroleum companies. Reema International Corporation, located in Denver, Colorado, signed on December 17, 1997 a Memorandum of Understanding with the National Gas Company of Trinidad and Tobago Limited for a project that would convert gas to liquid transportation fuels. Reema expects to finance, build and operate a plant costing about US275 million that will convert about 100 million ft3 day-' of natural gas into about 10000 bbl day-' of high quality transportation fuel. It is projected that the plant design and construction will require about three years. Exxon announced feasibility studies with Qatar's state-owned oil company to build a commercial plant at the 50 000 to 100 000 bbl day-' scale. The final decision on building this plant appears to have been delayed from the initial date expected for the decision.80Apparently the Exxon affiliate wants Qatar to sell the natural gas for less than they now sell it to petrochemical producers, which is believed to be about US$O.SO/Mcf (thousand cubic feet). Exxon expects to spend several million dollars in the next year to review an Alaskaspecific GTL application of its AGC-21 process. The potential Alaskan plant would be at the 50000 bbl day-' scale. This plant would have to compete for natural gas feedstock with liquified natural gas (LNG) production. An advantage of the GTL process is that the product could be transported through the existing oil pipeline and in this way extend the life of the pipeline as crude production declines. Sasol has been about as active as Syntroleum in forming joint ventures. Sasol together with Qatar General Petroleum Corp. (QGPC) and Phillips Petroleum have agreed to assess the possibility of a joint venture that would lead to a F-T plant with a 20000 bbl day-' facility.8' In April, 1998 Chevron

5: Indirect Liquefaction - Where Do We Stand?

177

1oa

80

2

. I

>

. I

60

c,

0

a

.-a?> 4-

I

40

a?

U

20

0

Gascat**

SiO, Supported Catalyst

Ti4-Suppor ted

Figure 5.27 Activity of cobalt-based catalysts. Reaction conditions; 220 "C, pressure = 450 psig. (** = Energy International catalyst) (from ref: 84).

and Sasol announced that they had reached agreement to pool their resources to begin design and engineering for construction of a 20 000 bbl day- GTL products plant in NigeriaqS2Chevron views this to include the possibility of expanding this new technology to other worldwide applications. Statoil and Sasol have formed an alliance for the conversion of natural gas to synthetic crude oil and liquid fuels by Fischer-Tropsch technology. Sasol's technology for the conversion of gas to fuels and Statoil's offshore and floating production technology for oil and gas are mutually complementary and are a basis for the alliance. Statoil prefers to cooperate rather than continue the development of their own Fischer-Tropsch technology, which had been underway for several years. The two companies will cooperate exclusively in developing floating and offshore applications and on a case-by-case basis for other applications. Wexford P.A. Syncrude Ltd. has, together with Bechtel Corp., made a design/economic assessment of a once-through natural gas Fischer-Tropsch plant with co-production of power.74 Presumably, the detailed design and scale-up technology that serves as the basis of this study is proprietary at this time, and belongs to Syncrude. The company has patented a catalyst that is highly selective for production of liquid hydrocarbons that contains cobalt and molybdenum or molybdenum and zirconium on an inorganic support .83 Energy International, a subsidiary of Williams Field Services, Inc., Tulsa, Oklahoma, announced their GasCat F-T process that is claimed to achieve higher productivity and improved catalytic performance (Figure 5.27) compared to existing Fischer-Tropsch technique^.^^ The process utilizes an

'

178

Catalysis

alumina-supported cobalt catalyst that shows significantly higher activities, among several advantages which include longer catalyst life, higher degree of regenerability, and lower cost than titania-supported catalyst. The process results from more than 20 years of effort by the Pittsburgh-based engineers and scientists, first as the Alternate Energy Development Department of Gulf Oil Corp. in the 1970s, and later as Energy International. The parent, Williams International Co., is exploring opportunities to licence the Gas-Cat catalyst and to co-invest in a grass-roots plant using the process. The liquid hydrocarbon output from a GasCat plant is claimed to be competitive with crude oil at a price of $16-17 per bbl. Energy International has recently completed a three year study, funded at approximately $3 million, on Fischer-Tropsch cobalt catalyst technology for the US Department of Energy. The DOE has also funded a study by Energy International Corp. on a 25000 bbl day-’ floating gas-to-liquids plant featuring Fischer-Tropsch synthesis. Howe-Baker Engineers, Inc., an international company specializing in the design and construction of plants to generate syngas and hydrogen, has announced that they are expanding their syngas operations to include FischerTropsch synthesis for the production of transportation fuels. They anticipate offering small plants suitable for use at the 1000 to 10000 range as well as the larger scale plants. Conoco/du Pont has initiated a large effort to develop Fischer-Tropsch technology for the conversion of natural gas to transportation fuels. To date, few details have been made public. Today, Rentech offers two approaches to the use of their Fischer-Tropsch technology. In the first, the 250 bbl day-’ plant constructed in Colorado has now been shipped to be located in the Kumachi gas field in India.85 Various dates have been circulated for the startup operation of this Indian plant, and discussions for two other plants are underway. Rentech is also negotiating with the Texaco Group, Inc. to establish a business relationship to accelerate the development and licensing of Rentech’s Process technology towards commercializing the technology on a worldwide basis.86 In the second approach, Rentech proposed the gasification of the 15000 bbl day- of heavy resid that a typical 100000 bbl day-’ refinery must dispose of today. A portion of the synthesis gas derived from the heavy resid could meet the refinery’s estimated 50 MW electrical requirement, with the remainder being used for Rentech Fischer-Tropsch technology to produce about 5500 bbl day-’ of liquid hydrocarbon products. As indicated above, an agreement has been reached whereby Rentech and Texaco will develop this technology. Carbon Resources is a Cyprus company that has announced that Automated Transfer Systems Corporation will provide capitalization to establish the commercial design parameters of its ‘SYNGEN’ process and the proprietary Fischer-Tropsch process of the N.D. Zelinsky Institute of Organic Chemistry of the Russian Academy of Sciences. Carbon claims strong prospects for plants from Algerian, Egyptian and Nigerian concerns. A feature of this operation is the conversion of natural gas via high-energy plasma using technology patented by Prof. Albin Czernichowski, University of Orleans

’

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(France). The syngas is generated by passing the natural gas, which may contain high contents of COZ, at a moderate temperature through a plasma arc generated by electricity. This process is claimed to provide substantial saving in capital costs and operating costs over conventional syngas formation. The syngas is generated at low pressure and this could be a disadvantage, except for Fischer-Tropsch processes that operate at a lower pressure. The second stage converts the syngas to hydrocarbons using a proprietary catalyst developed at Novocherkassk Plant of Synthetic Products located in Russia. Prof. A. Lapidus, N.D. Zelinsky Institute of Organic Chemistry, Moscow, is providing direction for Dr V.S. Boudtsov at the Novocherkassk plant. Prof. Lapidus’s published work on Fischer-Tropsch synthesis has emphasized cobalt catalysts that are operated at atmospheric pressure. The proprietary catalyst is claimed to have ‘chain-limiting’ properties and has been reported to have been confirmed in their lab in Orleans, France. The recent history of this company is complicated. The company was acquired by Synergy Technologies Corporation, a Colorado corporation. An agreement was signed in June, 1999 so that Texas T Petroleum Ltd. would acquire 50% of Carbon Resources, Ltd. Two of three early entrance co-production plants (EECP) selected for funding by US DOE will feature Fischer-Tropsch synthesis. The US DOE will fund these projects at the $6-8 million range during about five years, and the industries must provide up to 50% cost sharing. These projects are designed to be the first step toward developing advanced technology modules that would ultimately be integrated into an ultra-high efficiency, near pollution-free energy concept labeled ‘Vision 21’. One of these projects will be conducted by Waste Management and Processors, Inc. (WMPI), located in Frackville, Pennsylvania. The other will be managed by Texaco Natural Gas, Inc., of Houston, Texas. WMPI will evaluate the concept of using coal residue as a feedstock. They plan to utilize a Texaco gasification process and Sasol technology for the Fischer-Tropsch synthesis. In their study, Texaco will use Rentech’s Fischer-Tropsch technology to produce high-quality transportation fuels and electricity from coal and petroleum coke. The 5000 barrel day-’ plant considered for Pennsylvania has won a $47 million tax break from the State General Assembly. Plant developers can take up to 15% of monies that would normally be paid on state taxes to fund the cost of the project. The cost per barrel of daily capacity would be $62 400 and would be decreased to $53000 per barrel with the tax break. Even the lower cost appears to be about double the values quoted by most potential developers. The Texaco project fits into the ‘bottomless refinery’ concept that Texaco is advancing. The refiner should be able to generate their power need, sell excess power, produce a cleaner slate of fuels, run a heavier crude and maximize the value of refinery bottoms if the Texaco project reaches the commercial stage. At this time, there is significant activity and optimism concerning the introduction of a number of commercial operations using Fischer-Tropsch technology to convert natural gas to transportation fuels. An oil analyst for Morgan Stanley Dean Witter, Doug Terreson, anticipates that by 2005 refinery output worldwide will be up from the current 76 million bbl day-’ to 90

Catalysis

180

million bbl day- * and that gas-to-liquids plants could be contributing two to three million bbl day- .86 Based on the current research activity in the US and other countries, there is reasonable expectation that Fischer-Tropsch synthesis will assume a significant role in producing transportation fuels. This is made more likely by the projections that by about 2000 the world production of petroleum crude will peak and then decline from that point onward.87 The pessimist will take the view that they have heard this before - the US geared up in the 1950s and again in the 1970s to use synfuels to replace a predicted shortage of petroleum. The optimist will take the view that the date predicted for US petroleum production to peak (1960s) and then decline was correct, and that supply factors will dictate that Fischer-Tropsch synthesis commercialization will accelerate. The optimistic view must be correct at some date, and there are many reasons to believe that it will be within the next few years. According to a recent announcement, Shell International believes that it has gained much experience with Fischer-Tropsch synthesis during the operation of the Bintulu Middle Distillate Synthesis plant and are now pursuing additional opportunities form commercial projects. A recent coverage in the Remote Gas Strategies newsletter indicates that there is mention of a 50 000 bbl day- plant in Bangladesh. In this article, Jack Jacometti, Manager, Technology and Commercial Support for Shell International Gas Ltd., London states that, ‘We are looking at a variety of opportunities worldwide.’ He continues, ‘I don’t believe introducing new technology is a problem’, implying that Shell will move forward from the fixed bed reactor technology that serves for the Bintulu plant, and it is likely that the new technology will include a slurry phase reactor. 8

Summary of Current Status

There are two contrasting viewpoints of the current state of Fischer-Tropsch technology. On the one hand, there is the viewpoint of the investor and venture capitalist who are looking for situations where there is expected to be rapid growth during the short to mid-term period. On the other hand, there is the viewpoint of the scientific and engineering communities. It is not surprising that the two viewpoints may not agree and, in many instances, may be completely contradictory. The investor outlook currently appears to be very positive. The recent report entitled ‘Fischer-Tropsch Technology: Gas-to-Liquids, Solids-to-Solids, Liquids-to-Liquids’, by Howard, Weil, Labouisse, Fredrichs, Inc., a New Orleans-based investment firm, is representative of the optimistic outlook.88 For the short term, this report indicates that gasification will be utilized in downstream processing as an approach to handle heavy bottoms. It has been reported that Texaco is planning to increase refinery throughput by 40000 bbl day-’ using heavier curde at a 150000 bbl day-’ plant and using a deasphaltizer unit and a gasifier to handle the bottoms. The addition of a FT unit would increase the revenue from the plant significantly, and presumably this would be the short-term goal of the Rentech-Texaco joint effort. The

5: Indirect Liquefaction - Where Do We Stand?

181

report indicates that Exxon has a license agreement with Texaco to use the IGCC process at its refineries and chemical plants worldwide. For the short-term, the report indicates that the requirements to reduce CO2, such as required by the Kyoto Protocol, could make a dramatic expansion of Fischer-Tropsch technology. This could even reach to the gasification of coal and producing both electricity and hydrocarbon products. For the longer term, the report indicates that the gas-to-liquid process provided by the Fischer-Tropsch technology will permit today’s major refining companies to prevent the petroleum-producing countries from acquiring significant fractions of refinery operations. Most of these investment-type organizations indicate that there are five to six companies that are in the forefront of a larger group of potential developers of Fischer-Tropsch technology. For example, Tower89recently cited six such companies: Sasol, Shell, Exxon, Rentech, Syntroleum and Conoco. Nearly all of these organizations include Rentech and Syntroleum in the same grouping as Sasol and Shell and, because of the relative size of the companies, are impressed by the potential for growth by Rentech and Syntroleum. Clearly, Rentech and Syntroleum have been much more effective in publicizing their processes to the public than either Shell or Sasol has done to date. For instance, one entry to a web chat-site devoted to Syntroleum activities indicated that ‘he had heard that some company in South Africa had some activity that was similar to Syntroleum’. From the technical perspective, we would have a much different view of the ranking of current and potential company’s commercial Fischer-Tropsch technology. Sasol has operated three types of reactors at all scales from the small laboratory pilot plant to large commercial scales, and are therefore the most advanced in commercialization. They have operated fixed-bed reactors to produce high wax products for about 40 years, and currently produce more than 5000 bbl day-’ of products with these reactors. In spite of numerous operational problems, the Sasol operators were successful in developing ‘chain limiting’ technology and currently produce about 90 000 bbl day- of products using circulating, and now stationary, fluid-bed reactors. These operators developed their technology to utilize slurry bubble column reactors, scaling from laboratory size to a 3 m diameter large pilot plant and then to a 5 m diameter commercial reactor with an output of 1500 bbl day-’. No other company has reported such extensive operations. Shell has commercialized the fixed-bed reactor at an even larger scale than the Sasol operations, and expects to have a production of about 15000 bbl day-’ when their Bintula plant is back in operation. However, it appears that Shell’s largest bubble column reactor operation has been limited to short runs at the US DOE facility at LaPorte, Texas, and details of this operation have not been made public, To date, it does not appear that Shell has operated the fluid-bed reactors at a large scale. Exxon and Rentech have operated bubble column reactors of 200-300 bbl day- scale and have reported reliable operation at this scale. Exxon’s research and development effort has been reported to have cost more than $300 million

’

’

182

Catalysis

and that they have operated integrated gasification and Fischer-Tropsch units for about two years. On the other hand, Rentech’s operation at this scale has been limited to about three months. In addition to a longer operating experience, Exxon has vast experience in the design, startup and operation of a wide range of proven and of revolutionary process units. Other potential Fischer-Tropsch developers are, based on public information, much less advanced than the above companies. It is always hazardous to make such an assessment since some companies are effective in scaling to a significant size before any information reaches the public. For example, there are rumors that BP-Amoco is near to a commercial operation; however, neither company has a patent position that would indicate any activity beyond the laboratory or small pilot plant level. Amoco, for example, has had a very active Fischer-Tropsch research activity for more than ten years and this included the operation of several large-scale pilot plants. While these potential developers may have adequate, or even superior, technology, it remains to be advanced to a scale that most developers would consider adequate to scale to the size needed for a 10 000 bbl day- plant. Acknowledgement

This work was supported by US DOE contract number DE-AC22-94PC94055 and the Commonwealth of Kentucky. References 1 2

3 4 5

6 7 8 9 10 11 12 13 14 15

16 17 18

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Anonymous, Oil & Gas J., Dec. 30,1996. Syntroleum news release, 12/12/97. Syntroleum news release, 03/10/97. Houston Business Journal, 03131/97. Syntroleum news release, 04/10/97. Anonymous, International Journal for Measurement and Control, Jan., 1998. Syntroleum news release, 02/24/98. Syntroleum news release, 08/04/97. Syntroleum news release, 02/05/98. Syntroleum news release, February 6,1997. G.L. Beer (Inventor), Syntroleum Corp (Applicant), PCT/US97/19722, filed 30 October 1997. R.A. Dictor and A.T. Bell, Ind. Eng. Chem., Proc. Des. Dev, 1983,22 678. T.P. Kobylinski, US Pat. 4,088,67 1, May 9,1978. T.P. Kobylinski, C.L. Kibby, R.B. Pannell and E.L. Eddy, US Pat. 4,605,680, August 12,1986. H. Beuther, C.L. Kibby, T.P. Kobylinski and R.B. Pannell, US Pat. 4,413,064, November 1, 1983. H. Beuther, C.L. Kibby, T.P.Kobylinski and R.B. Pannell, US Pat. 4,399,234, August 16,1983. A.H. Singleton and S. Reiger, Proceedings of the 62nd Annual Convention - Gas Processors Association, 1983,pp. 162-165. C.B. Benham and M.S. Bohn, ‘A decade of research and development in FischerTropsch applications’, Pittsburgh Coal Conference, October 15,1992. Dr. J-L. Zhou, seminar at CAER, February 1998. J.M. Fox, I11 and S.S. Tam, Top. Catal., 1995,2,285. J.H. Gregor, Catal. Lett., 1990,7,317. G.N. Choi, S.J. Kramer, S.S. Tam, J.M. Fox, N.L. Carr and G.R. Wilson, 1997 Coal Liquefaction & Solid Fuels Contractors Review Conference, Sept. 3-4, 1997,Pittsburgh, PA. J.-P. Lange and P.J.A. Tijm, Chem. Eng. Sci., 1996,51, 2379. M.J. Gradassi, in ‘Natural Gas Conversion V’, A. Parmaliana et al., eds., Elsevier, Stud. Surf: Sci, Cat., 1998,119,35-44. Anonymous, C&E News, May 1 1, 1998,p. 10. Anonymous, C&E News, July 28,1997,p. 22. Anonymous, O&G J.,June 23,1997,p. 20. L.K. Aalund, O&C J, Special Issue, April 27,pp. 33-42,1998. Anonymous, The Oil Daily, Sept. 23,1997. Chevron News Release, April 15,1998. G.R. Wilson and N.L. Carr, US Pat. 5,639,798, June 17,1997. A.H. Singleton, O&G J., Aug. 4,1997,pp. 68-72. Anonymous, Oil & Gas J., 1997,95,36. Anonymous, Remote Gas Strategies, December 1997. C. Campbel and J. Laherrere, Sci. Am., March 1998. Gas-to-Liquids News, 1999,vol. 11, no. 1. B. Tower, ‘Strategic Implications of FT Technology on the Oil and Gas Industry’, Proc. of Gas-to-Liquids Processing ’99,San Antonio, TX,March 17- 19, 1999.

1997.

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75 76 77 78 79 80 81 82 83 84 85

86 87 88

89

6 Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts BY AKIFUMI UENO 1

Introduction

Steam reforming of methane into syngas (CO/H2) is the first step in the production of methanol and formaldehyde. Steam reforming is endothermic and requires significant energy, so that the direct conversion of methane into oxygenates in a single step is preferred (see Table 6.1). Despite decades of work, only a few papers have reported the yields of oxygenates higher than 4% based on methane.'-3 This is due to the inverse relationship between methane conversion and the selectivity to oxygenates; a high conversion corresponds to a low selectivity to oxygenates. Since formation of CO and C 0 2 during methane partial oxidation reduces the selectivity of oxygenates, suppression of CO and C02 formation, as well as an increase in methane conversion, is necessary to achieve high yields of oxygenates. CO is formed by the oxidation of methanol and formaldehyde, and C02 is produced mainly by the direct oxidation of methane.4 The oxidation of methanol and formaldehyde might be suppressed by a reactor design in which the residence time of the oxygenates on catalysts will be ~ptimized.~ The direct oxidation of methane into C02 might be controlled by the catalyst design. Thus, the key technologies for the direct conversion of methane into oxygenates (methanol and formaldehyde) are in reactor and catalyst design. Most of the work done before 1985 is reviewed by Pitchai and Klier,6 where Mo03-based catalysts were reported to be one of the most promising catalysts for partial oxidation of methane. Recent work published before 1990 is summarized in a review by Brown and P a r k y n ~elucidating ,~ the interaction between methane and catalyst surfaces in order to clarify the active sites for the selective production of oxygenates. In this decade, considerable work has been published on the formation of silicomolybdic acids (SMAs) by the solid state reactions between M o o 3 and Si02 in MoO3/Si02 catalysts in the presence of water vapor. It is still ambiguous, however, whether or not silicomolybdic acid plays an important role in the partial oxidation of methane to oxygenates. This paper will review the recent work concerning Mo03/Si02 catalysts for direct conversion of methane into oxygenates, and then focus on the generation, thermal stability, and the catalytic activity of SMA dispersed on Si02 surfaces for the direct conversion of methane. Catalysis, Volume 15 0The Royal Society of Chemistry, 2000 185

186

Catalysis

Table 6.1 List of thermodynamicparameters for some reactions of interest AH298 (kcal mol- *)

Reactions of interest

AG298 (kcal mol-

CH4 + H20 -+ CO + 3H2 H2 + C02 + H20 + CO 2H2 + CO + CH30H CH30H + 11202 HCHO + H20

49.1 9.8 -22.2 - 34.9

33.9 6.8 -6.5 - 39.7

CH4 + 1/202 -+ CH30H CH4 + 0 2 + HCHO + H20

- 30.7

- 65.6

- 27.1 - 66.8

-+

2

I)

Partial Oxidation of Methane on MoOJSiOz and Alkali Metaldoped MoOJSi02

The effect of molybdenum content in Mo03/Si02 catalysts on the partial oxidation of methane to formaldehyde was studied by Banares et al. using two kinds of Si02 supports; one with the surface area 86 m2 g-' and the other with 200 m2 g-1.8 The loading of Moo3 was converted into the number of Mo atoms exposed on Si02 unit surface (1 nm2), assuming complete dispersion of Moo3. (This assumption is suspicious because Mo ions coagulate to form Moo3 crystallites before complete dispersion.) The catalysts were used to carry out the partial oxidation of methane at 863 K using gases composed of CH4/ 0 2 (1 : 1 molar ratio) mixture. Figure 6.1 shows that the methane conversion and the yield of formaldehyde increased as the number of exposed Mo atoms increased up to approximately 1 Mo nmA2 for both of the Si02 supports, although the maximum with the lower surface area Si02 support is more pronounced. The methane conversion and the formaldehyde yield decreased as the number of Mo nm-2 was higher than 1.5, suggesting that the structure of active species generated on Si02 supports strongly depends upon the surface concentration of Mo ions. According to the paper by Deltcheff et al.,9 Moo3 crystallites are the main species on the catalyst surface at high Mo loading, and at low loading several kinds of molybdenum-oxo-species are formed, which are mainly responsible for catalytic reactions exhibited on MoO3/SiO2 catalysts. Banares et al.* further studied the effects of oxidants upon the selective oxidation of methane using 0 2 and N20. They concluded that pathways for the selective oxidation to HCHO and for the complete oxidation to C 0 2 were a redox cycle. The interaction between CH4 and lattice 02-ions in the Moo3/ Si02 catalyst is as shown in Scheme 1. 0 2 was concluded to be more effective lattice 02CH4(9)

-1

HCHO

adsorbed 0--

CO

lattice 02-

co2

-1

Scheme 1

gas phase oxidation

O2 or shift reaction

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

k,

3

0 2

5

O0 A

-

187

b

0

A

pQ,

NurnberofMoatomhm'

Number of Mo atoms/nm'

Figure 6.1 Conversion of methane and yield of formaldehyde on MoO&3iOz catalysts with (a) 86 m2g-' S O 2 support, and (b) 200 m2g-' Si02 support; 0means methane conversion and A means yield of formaldehyde

than N20 in reoxidizing the catalyst; whereas adsorbed 0- reacts nonselectively to produce CO. Aoki et al. prepared MoOJSi02 catalysts by two different methods: by a conventional impregnation with low Moo3 dispersions and by a soYgel method with high They tested these catalysts in the direct conversion of methane at 873 K in an excess amount of water vapor, and reported relatively high yields of oxygenates (ca. 4% yield of methanol and formaldehyde) on the highly dispersed catalysts, as given in Table 6.2. They concluded that the highly dispersed Moo3 over Si02 support might generate some complex active species, such as silicomolybdic acid, for the selective oxidation of methane in an excess amount of water vapor. Spencer et al.I2T1 studied the relationship between methane conversion and selectivity to formaldehyde during the partial oxidation of methane over Si02supported Moo3 and V2O5 catalysts at various temperatures. Figure 6.2 shows that the dependence of conversiordselectivity on temperature is weak on the MoO3/SiO2 catalyst, but strong on V205/Si02. This suggests that the mechanism of methane oxidation on Mo03/Si02is somewhat different from that on V205/Si02. Faraldos el al. l4 also compared the catalytic activities of Moo3/ Si02 with those of V205/Si02for selective partial oxidation of methane. The methane conversion on V205/Si02 catalyst was much higher than that on Mo03/Si02, but the selectivity to formaldehyde was higher on the MoO3/SiO2 catalysts at any methane conversion, as exhibited in Figure 6.3. The complementary selectivity trends between HCHO and CO in Figure 6.3 strongly suggest that HCHO was further oxidized into CO on both catalysts. The selectivity to C02 shows, however, a different trend; on the Moo3 catalyst C 0 2 was formed even at the very low conversion of methane, but no C02 was detected on V205 at the very low conversion and a further increase in the methane conversion resulted in an increase in C 0 2 formation. This indicates that HCHO and C02 are the primary products on the Moo3 catalyst, whereas on V2O5 catalyst C 0 2 is formed by further oxidation of CO. Most of the recent

1.o

2.9 8.2 12.0

0.58 1.2 4.0 6.6

conv. (010)

(K)

773 823 873 923 773 823 873 923

Methane

Temp.

I,

73 52 35 11

1

32 13 12

4

0 13 8 11

1

2 1

CHjOH 25 35 33 18 0 10 17 23

co

Selectivity (YO)

62

30 37

14

41 51 54 81

co2

0.20 0.17 0.52 0.07 0.86 1.7 3.8 1.8

HCHO -I-CHjOH

Yield (YO)

respectively, and the amount of catalyst employed was 1.5 g in all experimental runs.

HCHO

Flow rates of CH4.02 and water vapor were 1.8,0.2 and 2.0 L h-

Sol/gel

Impreg.

Catalyst

Table 6.2 Activity and selectivity of Mo031Si02 catalysts prepared by impregnation and sollgel methods for partial oxidation of methane in an excess amount of water vapor

1

873k 840k

921k

3

4

CH, conversion (%)

2

5

6

7

Q

i

i

3

4

s

CH, conversion (96)

6

7

9

5b h

g

Figure 6.2 Selectivitylconversion relationship for formation of HCHO during oxidation of methane with oxygen at different temperatures; (a) 3 Mo03ISi02 catalyst, and (b) V2O,lSiO, catalyst

0

0

A

8

* 896k

Catalysis

190 100

b

A

80 -

80 0

n

E

8 60 P :s

-

c

*L 1

5 6'0

h

0

I

$40

0

20 0

0

0

1

2

3

CH,conversion (%)

4

0

1

2

3

4

5

G

CH, conversion (%)

Figure 6.3 Selectivity to HCHO (O), CO (A) and CO, (0) vs. CH4 conversionfor (a) 0.8M0 and (b) 0.08 V: ( x M o and y V mean the number of Mo and V atoms on 1 nm2 of SiO,, surface)

work concerning the methane partial oxidation over Mo03/Si02 catalysts has accepted that HCHO and C 0 2 are the primary products of the reaction. Banares et a l l 5 studied the effects of additional alkali metal cations upon the structure and catalytic activity of molybdenum oxides in a 2.4% Moo3/ Si02 catalyst. New bands were observed in Raman spectra at 890 and 830 cm-' when alkali metal cations such as Na, K and Cs were doped into the catalyst. Since these alkali metal cations have higher affinity for surface molybdenum oxide species, it is concluded that the alkali-molybdates such as Na2M0207, K2M02O7 and Cs2M0207 were formed on silica surface. These alkali-molibdates were found to be thermally stable and not affected by hydratioddehydration treatments. The formation of alkali-molybdates on Si02 decreased the number of molybdenum oxide species (probably, isolated molybdate), hence the catalytic activity of Mo03/SiO2 for partial oxidation of methane decreased as the amount of doped Na ions increased, as shown in Figure 6.4(a). The number of remaining isolated molybdenum oxide species was estimated by monitoring changes in the intensity of Raman band at 986 cm- l, assigned to Mo=O band of the isolated molybdenum oxide species, with the amount of doped Na ions. The intensity was normalized to the intensity of Raman band of the S O 2 support at ca. 490 cm-', which could be considered as constant. Thus, the results given in Figure 6.4(b) show the turnover frequency (TOF) on the isolated molybdenum oxide species for HCHO production on Na-doped Mo03/Si02 catalysts. The TOF in Figure 6.4(b) is fairly constant, except for the sample with higher Na doping. This means that the poisoning mechanism involves the interaction of each Na ion with Mo=O bond in the isolated molybdenum oxide species, but not with Mo-0-Mo lattice bond in polymolybdate crystallites which was proposed by Spencer et a l l 6 Erdohelyi et a1.17 investigated the partial oxidation of methane on supported potassium molybdate catalysts in a fixed-bed continuous-flow reactor at ca. 860-923 K. The structures of the potassium molybdates deposited by impregnation were found to depend strongly upon the pH value

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

191

30

1

b A

h

r

cy

‘cn20

30 Y P

<0 10

0 E

$ 1

0.5 X

1

4

$10

I

k

U

Y

Y

8

0

0

400

800

1200

1600

wppm Na

Figure 6.4 (a) TOF of HCHO and the amount of remained surface Mo=O, and (b) TOF of HCHO referred to the remained surface Mo=O, (both are expressed as a function of the amount of Na added)

of the slurry containing support powder. The highest activity was observed on the magnesia-supported catalyst, but the complete oxidation of methane to carbon dioxide was dominant. On the silica-supported catalysts containing a large amount of KzM0207, a considerable amount of formaldehyde was produced when the reactant gas was composed of 10% 0 2 and 90% CH4 at 923 K. However, no formaldehyde was observed in the absence of oxygen in the feed gas, and products were only COZ and H20. On the basis of these observations, they emphasized the important role of adsorbed oxygen (0-)on the catalyst surface, i.e., lattice oxygen (02-) is involved in the complete oxidation of CH4 to C02, and the adsorbed oxygen (0-) enhanced the selective oxidation to HCHO as in Scheme 2. CH4(g) + O-(S)

-

CH~(S)+ O-(S) 2M06++ CH3-O(s) + OH-

CH~(S)+ OH-

-

CH~-O(S)

HCHO + 2M05++ H 2 0

Scheme 2

This mechanism starts with the elimination of a H atom from CH4 by adsorbed 0- ions to form methyl radicals, which was shown by Sun et aZ.18 who employed a double-layered catalyst bed of Sr/La203 (which predominantly produces CH3 radicals) followed by MoOJSi02. The formaldehyde yield obtained by this double-layered system was significantly higher than that obtained by a single bed system containing merely Mo03/Si02 catalyst. However, the lattice oxygen ( 0 2 -in ) the supported Moo3 catalysts has been generally accepted to play an important role in the selective oxidation of methane, as will be discussed later.

Catalysis

192

3

Active Sites on MoOJSiOz and Reaction Mechanism of Selective Oxidation of Methane

Liu et aZ.19 reported that the direct conversion of methane with N 2 0 was initiated by the formation of 0- ions at Mo(V1) sites, which were responsible for H abstraction from CH4 to form methyl radicals. Methyl radicals can react with the Moo3 surface to form methoxide complexes, leading to the formation of formaldehyde or methanol. Barbaux et aZ.20found, however, that 0- ions from N 2 0 migrated into Moo3 crystallites to form lattice 02-ions, which reacted with methane to produce formaldehyde selectively. There are two kinds of lattice 02-ions: one is the oxygen ions in Mo=O bonding, and the other is in Mo-0-Mo bonding of Moo3 crystallites. Smith and Ozkan2' suggest that Mo=O at the side plane of Moo3 crystallites is responsible for the production of formaldehyde, and Mo-0-Mo located at the basal plane promotes the complete oxidation into C02. They prepared unsupported Moo3 catalysts with different ratios of basal and side plane area, measured by computer-assisted two-dimensional SEM images. They also confirmed by TPR and FTIR measurements using l g 0 isotopes that oxygen ions in Mo-0-Mo bonding were more easily reduced than those in Mo=O bonding. Moo3 catalysts with large basal plane areas, named Moo3-R, were found to be more easily reduced than those with large side plane areas, named Moo3-C, suggesting that Mo-0-Mo bond were predominant on the basal planes and the Mo=O bond was mainly on the side planes of Moo3 crystallites. Figure 6.5 shows that the higher selectivity to formaldehyde was obtained on Moo3-C, whereas the higher selectivity to C02 was observed on the Moo3-R catalyst. This led to the conclusion that lattice 02-ions in the Mo=O bond at the side planes were responsible for selective production of formaldehyde, and those in the Mo-0-Mo bond at the basal planes promoted the complete oxidation of methane into C02. Similar results were reported by Hernandes and Ozkan22.23 for the selective oxidation of but-1-ene on unsupported Moo3 catalysts. They evidenced by Raman spectroscopy that more Mo-0-Mo sites were on the basal (010) planes and more Mo=O sites on the side (100) planes of Moo3 crystallites. They conclude that Mo=O sites are responsible for the selective oxidation, and that the complete oxidation was catalysed by the Mo-0-Mo sites. Banares and F i e r r 0 ~ ~have 9 ~ ~direct evidence of the incorporation of lattice 0 2 -ions of Moo3 crystallites into CH4 molecules to produce formaldehyde. First, they investigated the reaction using CH4 mixed with 1802 gas at 873 K over Mo03/Si02catalyst, and confirmed that no isotope-labeled products such as HCH180, C'802 and C160180were detected except a negligibly small amount of Cl80. This implies that oxygen either in the gas phase or on the catalyst surface is not involved in the reaction with methane. Then, they prepared the 180-labeled Mo03/Si02catalyst by H2-reduction of M o o 3 at 773 K to M002, followed by oxidation with 1 8 0 2 gas at the same temperature. The 180-labeled catalyst was used to carry out methane oxidation at 873 K using 1 6 0 2 gas. The results obtained are given in Figure 6.6, showing the production

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

HCHO

CO,

co

HCHO

193

co

CO,

Figure 6.5 Comparison of selectivities over Moo3-C and MoO3-R at equal conversion levels (2"/0)in methane oxidation. Moo3-C has a large side plane area and Mo03-R has a large basalplane area. P(a.u.)

Pfa

P( a.u .) 10.04

P (a.u .) I

0.

.02

0.

C1802

Reaction time (min)

0

10

20

30

40

Reaction time (rnin)

Figure 6.6 Changes in (HCH'601HCH'80) and (C'60~/C'60'801C'aO~) compositions following introduction of CH4/'602 into '"0-replaced Mo031Si02 catalyst at 873 K

of 180-labeled HCHO. From these observations, they emphasized the role of lattice 02- ions for the selective oxidation of methane on Moo3 through redox steps, i. e., a Mars-van Krevelen mechanism.26 Spencer et al. studied the effect of sodium ions on the partial oxidation of methane over a Mo03/Si02 catalyst, and found that sodium poisons the direct oxidation of methane into HCHO and C 0 2 , but promotes the oxidation of HCHO and CO into C02.16 In order to elucidate both the poisoning and promoting effects of sodium ions, they proposed a model for methane partial oxidation, as illustrated in Figure 6.7. The model starts with a thermal decomposition of the Mo-0-Mo bond to generate Mo-0- species on the surface of M o o 3 crystallites, followed by an attack of the methane molecule to form a methoxy intermediate, 1 [Figure 6.7(a)]. There are three routes from this methoxy intermediate to further products: first is the production of methanol by reacting with a hydrogen from an adjacent hydroxyl group on

194

Catalysis

1 Figure 6.7(a) Reaction pathway for initial attack of methane on Moo3

Figure 6.7(b) Reaction pathway for methane oxidation on MoO3ISiO2 catalyst

Moo3 crystallites [Figure 6.7(b)]. The second route is the production of formaldehyde by reaction of a hydrogen atom from the methoxy intermediate with the Mo(VI)=O bond. (If the catalyst is contaminated with sodium ions, this reaction route will be strongly inhibited because of the formation of Mo0-Na.) The last route is the direct production of C02, passing through the formation of a bridged species, 2, and the subsequent H-elimination by adjacent two Mo(VI)=O species. Further reaction of HCHO with the surface Mo(VI)=O species leads to the formation of formate species, 3. The replacement of one of the Mo=O bonds by Mo-O-Na will increase the electron density of the adjacent oxygen and facilitate the nucleophilic attack on HCHO. The formate intermediate, 3, decomposes to CO, assisted by the hydroxyl group on the surface of M o o 3 crystallites. The further reaction of CO with the surface Mo(VI)=O species forms the Mo-0-C - -0 intermediate, 4, which ultimately decomposes to produce C02. Spencer and Pereira' measured the relationship between the product selectivity and the methane conversion during partial oxidation of methane on a MoO3/SiO2 catalyst at temperatures of 843, 873, 898 and 923 K. They estimated the selectivity of HCHO at zero conversion of methane to be 89%,

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

195

indicating that HCHO is the primary product of partial oxidation of methane. The selectivity to CO at zero conversion of methane was 0% and increased with methane conversion, suggesting that CO is the secondary product. The selectivity to C02 did not change at any methane conversion, suggesting that C02 is also the primary product. On the basis of these observations, they proposed the following kinetic model for the partial oxidation of methane on the MoO3/SiO2 catalyst: CH4 + 0 2 -+ HCHO + H20 HCHO + 11202 + CO + H20 CH4 + 202 -+ C02 + 2H20

(rate constant: kl) (rate constant: k2) (rate constant: k3)

(6.1) (6.2) (6.3)

Reaction (6.2) (oxidation of HCHO) was studied independently and was found to be first-order in HCHO concentration. Assuming first-order in CH4 concentration and zero-order in 0 2 concentration, the rate constant at every reaction step above was estimated at the temperatures employed, and the activation energy at every step was also calculated (see Table 6.3). The relationship between product selectivity and methane conversion measured was well reproduced by the relationship calculated using the rate constants in Table 6.3(a). Amiridis et al.27also studied the kinetics of selective oxidation of methane on MoO$3iOz and V205/Si02 catalysts assuming the microkinetic models as given in Figure 6.8. The microkinetic models were based on the facts that HCHO and C 0 2 are the primary products of the methane oxidation on MoO3/SiO2, but only HCHO is the primary product on V205/Si02 catalyst, and that CH30H formed was promptly reoxidized into HCHO on both catalysts, since CH30H was not detected during the reaction. In order to optimize the rate constant of each reaction step in the microkinetic models, the pre-exponential factor and the activation energy of each step were estimated from the literature. For example, the activation energy was estimated according to the Polanyi expression, EA = Eo + aAH, where a was assigned a value of 0.5. The value of AH in the equation was first estimated from the heat of formation of the gaseous species involved in each elementary reaction, and then converted into the surface heat of formation by using the strength of its bond with the surface from literature data. The pre-exponential factors for adsorption of gaseous species on the catalyst surface were estimated to be approximately 1.2 x lo3 Pa-' s-l, assuming that the surface species were immobile. For reactions involving only adsorbed species and for desorption processes, the pre-exponential factors were estimated to be 1013 s-l, but were modified depending upon the extent of surface mobility of the intermediates, as well as their rotational freedom. The optimum rate constants obtained are shown in Table 6.3(b) and are compared with those obtained by the macrokinetic model in Scheme 3. CH4

c co2

R3

-

CH30H

1

R

CHsOCH,

Scheme 3

HCHO

CO

695 34 300 86

CH4 + 0 2 + HCHO + H 2 0 HCHO + 112 0 2 + CO + H20 CH4 + 202 + C 0 2 + 2H20

I)

189 134 189

R1 R2 R3

4.8~ 3.5 x 10' 5.8 x

1 . 9 lo-' ~ 9 . 5 10' ~ 2.4 x

Macrokinetic model

r2 r5 r8

MoO3lSiO2 catalyst

4.5~ 3 . 4 10' ~ 8.1 x

1 . 9 lo-' ~ 6 . 7 10' ~ 2.4 x

Microkinetic model

R6

& R5

7.1 x 1 . 4 10' ~ 5.7 x lo-'

5 . 6 lo-' ~ 3 . 6 10' ~ 1.3 x 10'

Macrokinetic model r2 r5 rll

V20sISi02 catalyst

1 . 4 lo-' ~ 1 . 4 10' ~ 5.3 x lo-'

I)

6 . 8 lo-' ~ 1.9 x 10' 1.6 x 10'

Microkinetic model

Activation energy (kJ mol-

Table 6.3(b) Turnoverfrequencies of CH4partial oxidation on Mo03/Si02 and V2051Si02 catalysts

Rate constant (h-

Elementary step

Table 6.3(a) Rate constants and activation energiesfor elementary steps

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

197

U

Figure 6.8(a) Active steps in the microkinetic model on the MoOjlSiO2 catalyst

Figure 6.8(b) Active steps in the microkinetics on the V2051Si02catalyst

The rate constants in the macrokinetic models were obtained experimentally, and they were in close agreement with those in microkinetic models obtained by kinetic simulations. The trend and the magnitudes of the rate constants on Mo03/Si02 catalyst given in Table 6.3(b) agree well with those given in Table 6.3(a).

4

Characterization of Surface Species Generated on MoO$302 by Heat Treatments

Spectroscopic studies to identify the surface species generated on Mo03/Si02 catalysts during preparation and subsequent heat treatments have been made recently. Raman spectroscopy is a useful tool to study the vibrational modes of metal-oxygen bonding in metal oxides. Most of the Raman work on supported molybdenum oxides is reviewed by Mestl and Srinivasa.28 Stencel et al.29730 studied the structures of Moo3 in silica-supported catalysts and concluded that there are three kinds of species present on the silica surface: crystalline Moo3; highly dispersed surface molybdates; and silicomolybdic acids (SMA; H4SiMo12040). SMA, which appears upon exposure to water vapor, was characterized by the strong Raman peaks at 980 and 962 cm-'. Crystalline Moo3 showed the peaks at 995 and 820 cm-', whereas the surface molybdates showed a peak at 955 cm- l . By monitoring changes in these peak intensities, they reported that SMA decomposed by heat treatment at 773 K, but was restored when the sample was contacted with water vapor at room tempera' studied the structures of surface molybdates on ture. Deskin et ~ 1 . ~also Mo03/Si02 catalysts. First, they observed that only Moo3 located in the

Catalysis

198

d

i

1100

l

l

900

,

,

700

I

I

I

500

l

300

l

30

Raman Shift (cm-')

Figure6.9 In situ Raman spectra of 5 wt% SMAISi02 catalyst in flowing oxygen saturated with water vapour at different increasing temperatures; (a) room temperature, (b) 443 K, (c) 503 K, (d) 573 K and (e) 773 K

surface layer of M o o 3 crystallites was reduced to MOO* at 623 K in H2/He flow, and at temperatures above 623 K bulk M o o 3 began to be reduced. Oxygen chemisorption measurements were carried out using Mo03/Si02 catalysts reduced at 623 K in order to estimate the dispersion of Mo on the catalysts. The degree of dispersion was found to depend upon the concentration of M o o 3 in the catalyst, and for the high concentration catalysts Mo ions were mostly dispersed in the form of small M o o 3 crystallites. Whereas for the catalysts with M o o 3 loading less than 1 wt%, Mo ions were well dispersed on the silica surface and Raman bands were detected at 955 and 677 cm-I; no bands at 980 and 962 cm-I suggesting that M o o 3 crystallites were not formed at these low surface concentrations. By analogy to Raman spectra of various molybdenum compounds, bands at 955 and 677 cm-' were assigned to M=O and Mo-0-Mo bonds, respectively, of distorted monomeric M o o 4 tetrahedral species, although recently the band at 955 cm-' has been accepted to be the band due to Mo=O of SMA, as mentioned below. Banares et aZ.32 have reported the formation of SMA on Mo03/Si02 catalysts during the selective oxidation of methane in the presence of water vapor, evidenced by in-situ Raman measurements. Raman signals assigned to SMA were observed at 975, 955, 615 and 240 cm-', when the catalysts were exposed overnight to a flow of oxygen with the saturated water vapor at room temperature [see Figure 6.9(a)].33 Figure 6.9 shows the thermal stability of

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

199

Table 6.4 Difference in the catalytic activity and selectivity of MoO3ISiO2 and SMAISiO2 catalystsfor selective oxidation of methane at 863 K % ! selectivit v

94 conversion

Catalyst 5% SMNSi02 5% Mo03/Si02

0.7 0.8

5.3 6.8

Reaction conditions: CH4 :0 2 = 7, W/F = 2 g h mol-

72 69

21 26

5 4

l.

SMA formed on the catalyst surface; the band at 240 cm-' remained constant when the catalysts were calcined at 573 K [see Figure 6.9(d)]. The bands observed at ca. 818,602 and 490 cm-' are due to the lattice vibrations of Si-0 bonds in the Si02 support. At temperatures higher than 573 K, the band at 240 cm- disappeared, and new bands appeared at 980 and 360 cm- characteristic of an isolated surface molybdenum oxide.34 Accordingly, silicomlybdic acids formed on Si02 were proved to decompose into Moo3 and Si02 above 573 K. Hence, no difference in the catalytic activity for selective oxidation of methane at 873 K was expected between MoO3ISi02 and SMNSi02 catalysts (see Table 6.4). They also reported the formation of P-Mo03 (Mo03-H20; molybdic acid) during methanol oxidation at 573 K because of the appearance of new Raman peaks at 838 and 487 cm- characteristics of P-Mo03. reported the formation of silicomolybdic acids on Si02 Deltcheff et supports, when Mo03ISi02 catalysts, prepared by calcination of a SiO2supported hexamolybdate sample at 773 K, were exposed to water vapor at room temperature overnight. SMA formation was evidenced by IR measurements; absorption peaks in the range around 955-957 and 902-914 cm-' are characteristic of Mo=O and Mo-O-Mo bonds of silicomolybdic acids, respectively. They studied the thermal stability of SMA, formed on the Si02 surface, by monitoring changes in the intensities of these IR peaks during heat treatments, and concluded that SMA on Si02 decomposed into Moo3 and Si02 above 563 K because of the appearance of a new peak around 1100 cm- assigned to the Mo=O bond in Moo3 crystallites. A different conclusion was, however, reached for low concentrations of SMA on Si02 catalysts. Moo3 crystallites were never observed at temperatures up to 773 K, although the SMA began to decompose at 563 K. They concluded that, for low concentration SMA/Si02 catalysts, Keggin structures in the SMA might decompose into polyoxomonomolybdate spread on the Si02 surface above 563 K. The thermal stability of SMA was not improved for SMNSiO2 catalysts, since the decomposition of pure SMA crystallites was observed at 573 K. This is controversial with the papers by Moffat et a1.,36937reporting that SMA stability was much improved up to 800 K when supported on Si02. Oxidation and dehydration of methanol are useful reactions to characterize the acidic and redox activities of catalysts employed. Dimethyl ether is produced on the acidic sites through the dehydration, and formaldehyde on the redox sites through oxidation of methanol. Thus, changes in the selectivity

',

',

~

',

1

.

~

9

~

~

Catalysis

200

0

K K

20

I

250

1

300

I

350

1

COD

Reaction temperature (C) Figure 6.10 Selectivity for methanol oxidation at 523 K on SMA generated over A40031 SiOr versus thermal pretreatment temperature: (0) HCHO, ( 0 ) HCOOCH3, ( 0 )(CH30)2cHZ,).( (CH3)Zo

of these products indicates the properties of active sites on the catalyst. Deltcheff et al.35investigated the acidic properties of the active sites on SMAI Si02 catalysts, prepared by exposing MoOJSi02 to water vapor overnight. As shown in Figure 6.10, the selectivity to dimethyl ether significantly decreased as the temperature of thermal treatment of the catalysts is increased, and stabilized at a low level (less than 5%) when the thermal treatment was carried out at temperatures higher than 583 K. The selectivity to formaldehyde (and methyl formate) increased with an increase in the thermal treatment temperature, and leveled off at temperatures higher than 583 K. This means that the SMA formed on the silica surface completely decomposed into Moo3 or surface molybdate species at 583 K, and that the SMA is not an active redox catalyst but has acid site needed to produce dimethyl ether from methanol. XANES (X-ray absorption near edge structure) is a useful tool to study the coordination states of metal atoms in metal compounds, and was successfully applied to determine the local site symmetry of supported molybdenum oxide phases in MoOJMgO catalysts.38 Hu et aZ.39studied extensively the structure of supported Moo3 catalysts by means of Raman and XANES measurements using various metal oxides as catalyst supports. Catalysts were prepared by incipient-wetness impregnation using an aqueous solution of ammonium heptamolybdate [ ( N H ~ ) ~ M o ~ O ~ ~followed - ~ H ~ Oby] , drying at 293 K overnight in air. These catalysts were called hydrated catalysts. They were calcined at 773 K in dry air and hence called dehydrated catalysts. A summary of the structures of molybdenum oxide species both in the hydrated and dehydrated conditions is given in Table 6.5(a). For hydrated catalysts, most of the molybdenum oxide species deposited over the supports were polymolybdate (M07024~- and M o ~ -)O ions ~ ~despite ~ the surface coverage, except over A1203 supports where MOO^^ - tetrahedral species were dominant at low

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

20 1

coverage. This is ascribed to the difference in the point of zero charge (PZC) of the supports used, since the structure of molybdenum oxide species in aqueous solution strongly depends upon the pH value of the solution. For a pH value of ca. 1.7-2.2 the main species in aqueous solution is and for the pH value of ca. 4.8-6.8 M07024~- is dominant. Above a pH value of 8.0, isolated and tetrahedrally coordinated MOO^^ - is the major species.4o It has been proposed that the final pH value of the solution in the filled pores of the support is close to the PZC of the support because of the fairly large capacity of the support at low Mo loadings.41 At high Mo loadings, the surface molybdenum oxide species drastically depresses the PZC of the support surface, hence the polymolybdate ions were deposited on A1203 at high Mo loadings, as given in Table 6.5. Whereas for the dehydrated samples, it was difficult to identify the molecular structures of polymolybdates generated on the different supports because of the dumped Raman spectra. It was, however, concluded that molybdenum oxide species on the dehydrated samples were highly distorted because of the appearance of a sharp band due to Mo=O at higher frequencies (ca. 1000 cm-'); for the hydrated sample the band due to Mo=O was observed in the region of 960-930 cm-l. It was also concluded from Raman measurement that isolated molybdate species were formed on the dehydrated Si02 supported samples because of the absence of the Raman band due to Mo-0-Mo bending around 220 cm- Coordination symmetry of the surface molybdenum oxide phase was judged from the splitting of the first fluorescent peak of XANES, which originates from the electron transition from the 2p to the 3d level in Mo ions, For tetrahedrally coordinated Mo ions the magnitude of splitting of the d-orbital is less than that of Mo in an octahedral coordination field, which is reflected to the spectral splitting in XANES measurements. The results given in Table 6.5(b) are, however, not clear enough to identify the symmetry ( T d or Oh),because of fewer reference data for polymolybdate samples so far. Hu et aZ.39942also determined the number of Mo atoms dispersed on unit surface area (1 nm2) of various metal oxide supports at monolayer coverage (see Table 6.6). The monolayer coverage of molybdenum oxide species was determined by the appearance of a Raman peak at 986 cm-I, assigned to the Mo=O bond of Moo3 crystallites. Table 6.6 shows a significantly small number of Mo atoms on Si02, compared with those on the other supports. This means a weak interaction of molybdenum oxide species with the Si02 support and a high mobility of molybdenum oxide species over the Si02 surface, which resulted in the aggregation to form Moo3 crystallites even at lower coverage of molybdenum oxide species on the S O 2 support.

'.

5

Catalytic Activity of SMA Generated On or Impregnated on SiO2

The possible presence of silicomolybdate on Mo03/Si02 catalysts was discussed in some early paper^,^^,^ and Castellan et ~ 1inferred . ~ the ~ presence of silicomolybdic acid (SMA) on the surface of Mo03/Si02 catalysts from the

Oxide support

XANES

* Major species, ** from Raman spectra.

Nb205

A 203 Ti02 Zr02 Si02

8.9 6.0-6.4 5.9-6.1 3.7-4.3 4.0

~~

PZC of support

Oh

Td

XANES

Structure at low loading Raman

Isolated

Isolated

Structure at low loading Raman

(b) In dehydrated catalysts

* Major species.

b205

Ti02 Zr02 Si02

A1203

Oxide support

(a) In hydrated catalysts

XANES

Structure in high loading Raman

Polymolybdate Polymolybdate Polymolybdate Isolated Polymolybdate

Structure at high loading Raman

XANES

Table 6.5 Summary of XANES and Raman structures of surface molybdenum oxide speices

h)

%C'

Q3

h)

0

180 55 39 55 380 5

20 6 4 6

* Theoretical monolayer coverage, ** based on monolayer surface area.

Si02

Nb205

ZrOz

A 203 Ti02

M003

(WYo M003)

Monolayer loading

175 53 39 47 275

8.0* 4.6 4.6 4.3 4.6 0.8**

Surface density Mo atoms nm-'

b

E

9

s

??

z!$

2

2

9 3

Surface area (m2 g - 9

Monolayer surface area (m2 g- I )

Support

$i?

Table 6.6 Surface density of Mo ions on the supported Moo3 catalyst at monolayer coverage

%

s

2.

sg

Catalysis

204

pH =7 a

pH=2

i

'

'

'

t

"

'

I

.

*

cm-1 Figure 6.1 1 IR spectra of the acetonitrile solution obtained after washing MoO3lSiO2 catalysts, prepared through the impregnation at (a) pH = 2, (b) pH = 7 and (c) pH = 1 1, and (e)IR bandposition of bulk SMA (H4SiM0,204,) 1200

1000

800

acidmetric titration of SMA in an aqueous leaching solution. Significant attention has been paid to the catalytic activity of SMA either generated on or impregnated on a Si02 support for the partial oxidation of methane. However, this reaction has been generally carried out at temperatures as high as 873 K, and the thermal stability of SMA has been explored. As was mentioned above, Stencel et al.29730were of the opinion that SMA generated on Si02 was stable at calcination temperatures up to 773 K, but Banares et aZ.32reported that SMA generated on Si02 decomposed upon calcination at 573 K. Deltcheff et al.35 also reported that the thermal stability of SMA in silica-supported catalysts was reduced by 20°C, comparing them to the thermal stability of pure SMA crystallites. Kasztelan et emphasized that the thermal stability of SMA was significantly improved when supported on SiO2 because they observed the that Keggin structure in the silica-supported SMA was largely intact even after calcination at 773 K. The existence of SMA in the calcined catalysts was demonstrated by IR and Raman analyses of the filtered solution obtained by washing the calcined catalysts with acetonitrile in order to extract the supported (or generated) SMA species selectively. Figure 6.1 1 shows IR spectra of the acetonitrile solution obtained by washing the calcined catalysts. They prepared four kinds of catalysts; three of them were prepared aZ.36737

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

205

by an impregnation of Si02 with an aqueous solution of ammonium molybdate at different pH values, and the rest were prepared by impregnation with an aqueous solution of SMA (12-molybdosilicic acid) at pH 2.5. These catalysts were calcined at 773 K, followed by acetonitrile extraction of the SMA species that were generated and remained on Si02 during preparation and calcination. All spectra shown in Figure 6.1 1 exhibited absorption peaks at 950, 920 and 810 cm-', which were characteristic of SMA.46Judging from changes in the peak intensities, it was possible to say that more SMA was generated on the catalyst prepared by impregnation at the lower pH value. This seems to be due to the fact that Si02 powder is easily dissolved in the acidic solution, and the dissolved Si02 will react with polymolybdate anions in the solution to generate SMA. They tested these catalysts for the methane oxidation reaction at 773 K using N20 as an oxidant. The results are given in Table 6.7, indicating the higher conversion of methane on the catalyst prepared in the lower pH solution. Consequently, they reported that the methane conversion increased as the amount of SMA generated on the catalyst increased, though the selectivity to formaldehyde decreased. Barbaux et ~ 1 also found that the selectivity and yield of HCHO increased with the amount of SMA generated in the MoO3/SiO2 catalysts. In another paper,48 Kasztelan and Moffat concluded that the proton in the SMA species was crucial in the oxidation process. According to their conclusion, protons are required to create an oxygen vacancy in the Keggin unit of the supported heteropolyoxometalate such as SMA, and that oxidants such as N20 in the gas phase adsorb on to the vacancy to form 02-species as follows,

.

2KUOH + KUO + K U D + H20 K U U + N20 + KUO + N2 where KU means Keggin unit, and KUC] represents the oxygen vacancy created. The 02- ions thus created promote the activation of CH4 to yield methyl radicals. They reported that the methane oxidation reaction on SMAI Si02 catalysts was poisoned by addition of Cs ions into the catalysts. Protons in the Keggin unit might be replaced with the Cs ions added to form a KUOCs structure, hence the Cs ions block the creation of oxygen vacancies in the Keggin unit through dehydration. Accordingly, the role of protons in SMA was concluded to create the oxygen vacancies, where methane oxidation was enhanced. They briefly mention the possibility of protonation of methane, H+ + CH4 + CH5+, to activate CH4. It has been reported that protons in the solid superacid, consisting of Cr3+-Zr4+-S04/Si02, attacked methane to produce CH5+as an intermediate to oxygenates such as methanol and f ~ r m a l d e h y d e . ~ ~ Olah et ~ 1 were . the ~ first ~ to propose the generation of penta-coodinated carbocations (CH5+) during the ionization of methane in the presence of a superacid (SbF5/HS03F) at 353-423 K. Their work was extended to the direct production of methanol from CH4 in the presence of the superacid, but hydrogen peroxide was necessary for the production of methan01.~' SMA formation by the solid state reaction between Moo3 and Si02 in the presence of water vapor can be expressed as follows;32

~

3.0 3.2 3.2 5.7

-

Moo3 laoding (wt%)

0.8 1.1 2.6 3.3 5.0

3 .O 4.0 8.4 10.0 19.6

Conversion (%)a CH4 N20 9.8 19.0 22.6 25.2

-

N20 TO@ 75 72 67 61 46

co

25 19 24 37 54

t t t -

t

-

CHjOH 9 9 2

-

Selectivities (O/o)a HCHO

co2

a

Reaction conditions are as follows: T = 773 K, mass of catalyst = 2 g, flow rate = 15 ml min-I, composition of CH4 :N20 = 33 :67. Apparent turnover number expressed in terms of molec atom-’ of Mo s - ’ . SiOz is from Davison-Grace and Grade-400, surface area is 750 m2 g.

Si02 support Mo03/Si02 (pH 11) MoO3/SiO2 (pH 7) Mo03/Si02 (pH 2) H4SiMo ,204dSi02

Catalysts

Table 6.7 Activities and characteristics of the catalyst prepared at different p H values

o\

N 0

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

100

I

207

50

20

B

I

0 0

I

2

L

3

4

Flow rate of water vaporovb) Figure6.12 Changes in the catalytic activity of 27 wt% SMAISiO2 catalyst for the selective oxidation of methane at 873 K with the fraction of water vapor in the feed gas, consisting of CH4 (1.2 L h- l ) and 0 2 (0.8 L h-I): (0) oxygenates ( 0 )CO, (A) C 0 2selectivity and (a)CH4 conversion

Si02 + 12Mo03 + 2H20 -, H4SiMo120a This reaction takes place even at room temperature, but the reaction rate is too low to detect the SMA formed over a short period. The rate of SMA formation will be faster at elevated temperatures, but the SMA species formed on Si02 are readily decomposed into Moo3 and Si02 at elevated temperatures because of the poor thermal stability of SMA. Sugino et al.52 thought that the rate of SMA formation will depend upon the fraction of water vapor in the feed gas, and that the rate of SMA formation rate will be faster compared to the decomposition rate when an excess amount of water vapor is added to the feed gas. Then, they studied the effects of water vapor upon the thermal stability of SMA and measured the catalytic activity for the selective oxidation of methane at 873 K using a CH4 : O2 mixture (molar ratio = 3 :2) on 27 wt% SMA /SO2 catalysts. The results are shown in Figure 6.12, which indicates a high yield of HCHO (ca. 20%) when the water vapor is higher than 50% of the feed gas. The selectivity to HCHO was around go%, and to CO and C02 about 5% each, at a methane conversion of ca. 20%. They concluded that the role of the excess amount of water vapor is to restore the SMA species rapidly on the Si02 surface from Si02 and Moo3, since IR spectrum of the catalyst after reaction showed the absorption peaks at 957 and 910 cm-l, assigned to the vibration of the Mo=O and Mo-O-Mo bonds of SMA, respectively. Further studies were made on the effects of SMA loading upon the catalytic activity for the selective oxidation of methane at 873 K, and revealed that the yield of HCHO increased when the SMA loading increased, and leveled off at a certain loading where a

Catalysis

208

100

80

h

8 6 0 i=,

.">

-4

Y

0

340 m

20

0

0

50

too

150 200 TrmcF)

250

300

350

Figure 6.13 Durability test of 27 wt% SMAISi02 catalyst for selective oxidation of methane at 873 K using CH4: 0 2 ( 3 :2 molar ratio) and water vapor (60 vol% in the feed): (0) oxygenates, (0) CO, (A) C02 selectivity and (m) CH4 conversion. Carbon balance (?A)measured by chemical analysis is given at the top of thejigure

monolayer dispersion of SMA on the Si02 surface was established. As can be seen in Figure 6.13, the catalytic activity did not decrease during a 350 h durability test at 873 K. The reactant gas fed to the catalyst (1.5 g) was a mixture of CH4 and O2 with a 3 : 2 molar ratio where the flow rates were 1.2 1 h-' for CH4 and 0.8 1 h-' for 0 2 , respectively. The flow rate of water vapor added was 3 1 h-I, corresponding to 60 vol% of water vapor in the feed gas. The amount of HCHO formed was measured by the chemical analysis (titration) as well by gas chromatographic analysis, and the carbon balance was 90- 110% (see the top of Figure 6.13). 6

Other Topics Concerning Heteropoly Acid Catalysts or Partial Oxidation of Methane

A number of papers concerning heteropoly compound catalysts were introduced by Okuhara, Mizuno and M i s ~ n o , but ~ ~ the . ~ ~papers reporting the methane selective oxidation on SMA at the higher temperatures are few because of lower thermal stability of SMA. Most of the papers were concerned with molybdophosphoric or tungstophosphoric acids for both homogeneous and

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

209

heterogeneous reactions. The papers reporting the synthesis of fine chemicals with heteropoly acids are summarized by K o z h e ~ n i k o v abut , ~ ~ most of the papers were concerned with homogeneous reactions at lower temperatures. Parmaliana et al.56 studied the partial oxidation of methane on various commercial, bare Si02 samples, and reported that the highest HCHO productivity (STYHCHO; space time yield of HCHO) was obtained with ‘precipitated’ SO2; ‘fumed’ Si02 resulted in the least reactive silica (see Table 6.8). They further reported that incorporation of Moo3 depressed the STYHCHo for the ‘precipitated’ SO2, but enhanced the STYHCHo for the ‘fumed’ SiO2. In contrast, addition of V2O5 to either ‘precipitated’ or ‘fumed’ Si02 led to a higher STYHCHO. They conclude that the reaction occurred on the reduced sites of Si02, where 0 2 was adsorbed and activated to react with methane. Moo3 and V2O5 may modify the catalytic properties of SiOz by affecting the process of oxygen activation on the catalyst surfaces. Kasztelan and M ~ f f a t ~ ~ also mentioned the catalytic activity of bare Si02 for the selective oxidation of methane. They said that HCHO was the main product when 0 2 was used as an oxidant at 773-873 K, but that C02 was dominant when N20 was employed. V205/Si02 is also a selective catalyst for the production of HCHO from methane, although the oxidation mechanism is somewhat different from Mo03/Si02 catalysts. Sun et aZ.58have studied the activity and selectivity of V205/Si02catalysts for the partial oxidation of methane, and reported that the catalytic performance was strongly dependent upon V2O5 coverage. At very low conversion of CH4, the selectivity to HCHO increased linearly with V2O5 coverage, indicating that the isolated V5+ species were responsible for the active sites. The selectivity exhibited a maximum at 1 wt% V205content, and that the lower selectivities at higher loading appeared to be due to the increase in the Lewis acidity of the catalysts. McCormic et al.59investigated the effects of Fe and Cr added to VOP04 catalysts on the selective oxidation of methane at 573-698 K using 0 2 as an oxidant. They found that both Fe- and Crpromoted VOP04 catalysts exhibited an enhanced selectivity to HCHO at very low levels of methane conversion. The activation energy was near 102 kJ mol-’ regardless of the presence of Fe and Cr ions in the catalysts. Since the active sites for HCHO production were V5+ (based on the results of 31PNMR measurements), the role of doped Fe and Cr ions was either to increase the rate of V5+ formation during activation or to stabilize the domains containing V5+ under the highly reducing methane oxidation conditions. Wang and Otsukam reported the significant effect of adding H2 to a CH4/02 mixture on the production of methanol from methane on iron phosphate catalysts (FeP04). Methanol was produced in the presence of H2 in the feed gas, but no methanol was observed when H2 was absent from the feed. Only a small amount of formaldehyde was detected in the absence of HZ.They studied the electronic states of Fe in the catalysts using XPS in order to reveal the role of co-fed H2 on the direct production of methanol from methane, and concluded that methanol production takes place through the interconversion of Fe(II1) to/from Fe(I1) in the catalysts. The reductive activation of oxygen by hydrogen on these iron atoms was proposed to generate the active and selective

Preparation method

Precipitation

Precipitation

SoUgel

Pyrolysis

Sample

Si 4-5P

D11-11

250 MP

M5

200

180

137

38 1

(m2 g-

9

s.A .BET ~

600 650

600 650

600 650

600 650

~~

TR (“C)

18.58 60.1 1 3.28 20.76

7 11 10

20 19 11 13

73 70 79 75

12

54.43 159.17 13 23

0.40 2.53

8.93 31.35

(g : ;&k

h-l)

STYHCHO 115.85 303.84

‘

7 14

~~

C02

18 23 31 30

CO

Selectivity (%) HCHO

2.33 7.99

s-])

75 63 56 47

14.36 44.29

Reaction rate molcH,&;;

Table 6.8 Formaldehyde productivity of commercial S O 2 samples in the partial oxidation of methane

6: Partial Oxidation of Methane Over Silicomolybdic Acid Catalysts

21 1

oxygen species for the conversion of methane into methanol. Cheng6' studied the oxidation of methanol and formaldehyde on Mo03/SiOz catalysts using IR spectroscopy, and reported that CH30H, HCHO and water vapor, one of the products of the oxidation reaction, adsorbed competitively on the same sites of the catalyst surface. Hence, the water vapor depressed the adsorption and successive oxidation of HCHO on the catalyst into CO and CO2 during partial oxidation of methanol. The water vapor that was co-fed improved the selectivity to HCHO in the selective oxidation of methanol. 7

Summary of the Selective Oxidation of Methane on Mo03/Si02 Catalysts

About 60% of the energy required for the methanol production process is consumed in the steam reforming process for the production of syngas (CO and H2). Accordingly, the development of the direct conversion of methane into methanol (and formaldehyde) can save a significant amount of energy, and thus depress the amount of C02 emitted into the atmosphere. The key technologies for a new and effective process are in the design of the reactor to suppress the sequential oxidation of methanol and/or formaldehyde, and the catalyst design to improve the methane conversion and the selectivity to oxygenates. To clarify the active species on silica-supported Moo3 and V205 catalysts, in-situ Raman spectroscopy has been used to observe the chemical species generated on the catalyst surface. Although the in-situ Raman spectroscopy has revealed the structures of molybdenum oxide species on the Moo3/ Si02 catalysts on the working catalyst, the conclusions obtained so far are still complicated and controversial. This is probably because of the presence of a variety of polymolybdates, depending upon the preparation and heat treatment conditions of the catalysts. The structural changes in molybdenum oxide species during heat treatment of Mo03/Si02 catalysts are summarized in Scheme 4. (low Moo3 loading) Mo03/Si02 catalyst<

(high Moo3 loading)

-

-

isolated Mood2-

polymolybdates

water vapour

* SMA formation

-

heat above 573 K

heat above 573 K

MOOScrystallites

Scheme 4

It is still ambiguous whether the SMA generated on the catalyst surface is the actual active species for selective oxidation of methane into oxygenates. In order to answer this question, and to design a new type of catalyst, the role of protons (H+) in SMA should be well understood. Another problem is concerned with the thermal stability of the SMA in the working state when used for selective methane oxidation, because many controversial results have been published. Although an extremely high yield of oxygenates has been obtained in our l a b ~ r a t o r ycareful , ~ ~ procedures to heat up the catalyst to the reaction temperature (873 K) will be required to reproduce our results.

212

Catalysis

Otherwise, the SMA was easily decomposed into Moo3 and SiOz, and the results similar to those on the conventional MoO3/SiOz catalysts are obtained.

References 1 2 3 4 5 6 7 8 9

10 11 12 13 14 15 16 17 18

19 20 21 22 23 24

25 26

27

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