Catalytic Partial Oxidation of Renewable Feedstocks

Transcription

1 Catalytic Partial Oxidation of Renewable Feedstocks A dissertation submitted to the faculty of the Graduate School of the University of Minnesota by Reetam Chakrabarti In partial fulfillment of the requirements for the degree of Doctor of Philosophy. Advisor: Lanny D. Schmidt July, 2012

2 c Reetam Chakrabarti 2012 All Rights Reserved

3 ACKNOWLEDGMENTS I owe a lot of gratitude to my advisor, Professor Lanny Schmidt for giving me the opportunity to join his group, and the freedom to pursue my ideas. His breadth of knowledge and unending curiosity about various things in the world have always amazed me. I thank him for always having an open door to discuss problems. I would like to thank Professor Aditya Bhan for his help, both academic and non-academic over the past four years, ever since I applied to the University of Minnesota. I would also like to thank Professor Alon McCormick and Professor Ulrike Tschirner for agreeing to be a part of my defense committee and reviewing my thesis. Thanks to the staff at the Characterization Facility at the University of Minnesota for answering questions regarding characterization techniques presented in this thesis. Thank you to all the teachers through school and college in India. I have enjoyed working with the graduate students and undergraduate students in the lab, all of whom have helped me in my experiments. In particular, I would like to thank Dr. Dave Rennard, Dr, Brian Michael, Dr. Josh Colby, Dr. Jake Kruger, and Sam Blass for their help during various stages of my research. I would like to thank Dr. Dave Rennard, Dr. Brian Michael and Dr. Josh Colby for helping me get started in my first summer with experiments starting with an empty hood. I thank Dr. Joshua Colby and Dr. Brian Michael for the great conversations in the lab, on our way down to the coffee room on the first floor, and in their house, the Green Monster. Both of them have been an inspiration with their engineering creativity and approach to doing experiments. Outside of his immense help in my research, I am grateful to Dr. Jake Kruger for a number of things which include the great conference trips, organizing a trip from the sky to the ground, and the trip to Sheboygan, WI for his wedding. I would like to thank Sam Blass for his comments on writing, and the research and non-research related conversations in the ping-pong room next to the lab, and for tennis and ice skating. I would like to thank Dr. Christine Balonek for her wonderful desserts and Jeremy W. Bedard for tickets for baseball games and i

4 ii rugby. Thanks to Richard Walter Hermann for help with lot of the experiments in this thesis. Thanks to Michael Skinner for introducing me to number of things while initially in the US, including darts and Easter gifts. Thank you to Hui Sun, Alex Marvin, David Nare for their help in research and insightful comments during group meetings. Thanks to the undergraduate students, Nils Persson, for his great commedy shows and piano skills, Tyler Josephson, Eric Hansen, Ed Michor, and Sheila Hunter for their help. I would like to thank my friends, Shameek Bose, Sujit Jogwar, Srinivas Rangarajan, Romain Le Picard, Aloysius Gunawan, Andrew Yongky, and my roommate Parthiv Daggolu for great moments in Minneapolis. In addition, thanks to my friends from undergrad, Umang Desai, Karan Kadakia, and Mansi Seth for great conversations that persisted through the last four years. I would like to thank my family in India, especially my parents and my sister Ishita for their support throughout my life.

5 ABSTRACT The current world energy and economic infrastructure is heavily reliant on fossil fuels such as coal, oil, and natural gas. The limited availability of fossil fuels along with environmental effects and economic uncertainties associated with their use has motivated the need to explore and develop other alternative sources of energy. Lignocellulosic biomass like fossil fuels is carbon-based and has the potential to partly supplant the energy supplied by fossil fuels. Lignocellulosic biomass is a complex mixture of polymers such as cellulose, hemicellulose, lignin along with small concentrations of inorganics and extractives. Recent research has shown that lignocellulosic biomass and biomass model compounds can be processed autothermally by catalytic partial oxidation in millisecond residence times over noble metal catalysts at high temperatures ( C) to syngas (a mixture of carbon monoxide and hydrogen). 1 3 The syngas stream can then be upgraded to fuels and chemicals. In Chapter 2, spatially resolved concentration and temperature profiles of methane and dimethyl ether, a model compound for biomass, are compared. Dimethyl ether can be produced renewably through syngas upgrading. Maximum temperature and concentration gradients were found within the oxidation zone. Most of the oxygen ( 95 %) was converted within the first 2.2 mm and syngas formation was observed despite the presence of oxygen. The catalytic partial oxidation process has been demonstrated using compounds which, unlike most biomass sources, contain negligible quantities of inorganics. Some of these inorganics have catalytic properties themselves and some may act as poisons for the Rh-based catalyst. The effects of common biomass-inorganics (silicon, calcium, magnesium, sodium, potassium, phosphorus, sulfur) on rhodium-based catalysts in autothermal reactors have been studied. To understand the effects of biomass inorganics on Rh catalysts, two sets of exiii

6 iv periments surveying common inorganics were performed - in the first set, inorganics were directly deposited on the rhodium catalyst and tested using steam methane reforming as a model reaction (Chapter 3); whereas in the second set, inorganics were introduced to a clean catalyst in an ethanol feed to simulate actual inorganiccontaining biomass (Chapter 4). In both sets of experiments, performance testing, catalyst characterization and regeneration were carried out to probe the mechanism of inorganic interaction with the rhodium-based catalyst. Large decreases in reforming activity were observed on phosphorus- and sulfur-doped catalysts. Deactivation due to calcium and magnesium was primarily due to blocking of active sites. Potassium and silicon were volatile at the high temperatures within the reactor. Potassium introduced alkaline chemistry promoting acetaldehyde formation from ethanol while phosphorus introduced acid chemistry promoting formation of ethylene from ethanol. The effects of potassium and phosphorus on catalytic partial oxidation of methane and ethanol at different concentrations and temperatures have been studied in Chapter 5. The synergistic effects of potassium and phosphorus were studied by distributing the inorganics together on the catalyst as monobasic potassium phosphate. The effects of both potassium and phosphorus were observed in the catalytic partial oxidation of methane on a potassium phosphate-doped catalyst at low temperatures. At high temperatures, only effects due to phosphorus were observed because of potassium volatilization. The results show that biomass-sources containing low concentrations of inorganics can be processed autothermally to a high selectivity syngas stream. The distribution and interactions of the inorganics within the catalyst can be used to design better pretreatment, processing, and regeneration strategies to minimize catalyst deactivation during biomass processing. Alcohols represent an important intermediate in different biomass upgrading routes. Chapter 6 discusses the behavior of butanol isomers, 1-butanol, isobutanol, 2-butanol, and tert-butanol over four different catalysts; Rh, Pt, RhCe, and PtCe at different fuel to oxygen (C/O) ratios. At low C/O ratios, equilibrium species such as CO, CO 2, H 2 and H 2 O were obtained while non-equilibrium species such as carbonyls and olefins were dominant at high C/O ratios. Low reforming activity was observed on Pt and PtCe catalysts. All isomers decompose primarily by dehydrogenation through a carbonyl intermediate except tert-butanol which decomposes by dehydration to isobutene; however, the reactivity of tert-butanol was unaffected.

7 v In Chapter 7, isobutanol autothermal reforming is integrated with a water gas shift stage downstream to produce hydrogen containing low concentrations of carbon monoxide for portable fuel cell applications. A RhCe-based catalyst was selected to carry out autothermal reforming of isobutanol while a PtCe catalyst was selected for the water gas shift stage. This staged reactor produced high yields of hydrogen (> 120 % selectivity) containing low concentrations of CO (< 2 mol %) in less than 100 ms making the effluent ideal for portable high temperature PEM fuel cell applications. The water gas shift stage also reduced the concentration of non-equilibrium products formed in the autothermal reforming stage by over 50 %. Thermodynamic analysis of the system showed that staged autothermal reforming of isobutanol integrated with a fuel cell can potentially lead to 2.5 times more efficient energy usage when compared to burning isobutanol in a conventional combustion engine. The results in this thesis give an insight into the mechanisms and processing challenges involved in converting renewable feedstocks to syngas by catalytic partial oxidation. Further experiments based on the conclusions of this thesis are discussed in Chapter 8. Spatial profile experiments to determine roles of mass transfer, steam reforming, and dry reforming during catalytic partial oxidation of oxygenates are proposed. Spatial profile studies for catalytic partial oxidation over inorganic-doped catalysts and feed are proposed to determine their concentrations and nature on the catalyst surface during reactor operation.

8 CONTENTS Acknowledgments Abstract Table of Contents List of Tables List of Figures i iii vi x xi 1 Introduction Current World Energy Scenario Limited Fossil Fuel Reserves Developing countries Climate Impact Energy Security Biomass as an Energy Source Structure of Biomass Converting Biomass to Fuels and Chemicals Upgrading Syngas for Energy Use Fischer-Tropsch Process Methanol Route Hydrogen Generation for Fuel Cells Catalytic Partial Oxidation Evolution of Research Upgrading Products from Catalytic Partial Oxidation Inorganics and Biomass Processing Poisoning Fouling Sintering Attrition/Mechanical Failure vi

9 CONTENTS vii Solid State Transformation Summary Mechanism of Catalytic Partial Oxidation Introduction Experimental Spatial Profiles Nature of Rhodium in Oxidation Zone Results Spatial profiles State of Rhodium in Oxidation Zone Conclusion Acknowledgements Effects of Biomass Inorganics on Rhodium Catalysts: I. Steam Methane Reforming Introduction Experimental Experimental Setup Product Analysis Catalyst Preparation Experimental Procedure Catalyst Characterization Results Performance Testing Catalyst Characterization Equilibrium Calculations Discussion Sulfur Phosphorus Silicon Sodium and Potassium Calcium and Magnesium Conclusions Acknowledgements Effects of Biomass Inorganics on Rhodium Catalysts: II. Autothermal Reforming of Ethanol Introduction Experimental Results and Discussion Silicon Sulfur

10 CONTENTS viii Phosphorus Potassium Sodium Calcium Magnesium Comparison of Inorganics Conclusions Acknowledgements Effects of Potassium and Phosphorus on Rhodium Catalysts for Catalytic Partial Oxidation Introduction Experimental Results Doping at Different Concentrations with Methane CPO Doping at Different Concentrations with Ethanol CPO Transient Studies with Methane CPO Transient Studies with Ethanol CPO Discussion Effect of Potassium Effect of Phosphorus Conclusions Acknowledgements Autothermal Partial Oxidation of Butanol Isomers Introduction Experimental Results Conversion and Temperature Syngas and Combustion Products C 4 Intermediates Other Intermediates Discussion Chemistry of the Isomers tert-butanol Isobutanol Surface Chemistry Effect of Catalyst Conclusion

11 CONTENTS ix 7 Autothermal Reforming of Isobutanol Introduction Experimental Catalyst Preparation Product Analysis Results CPO of Isobutanol Addition of WGS Stage Discussion CPO stage CPO + WGS Conclusions Acknowledgements Summary and Future Work Spatial Profiles Examination of Mechanism of DME CPO Inorganics Extending Spatial Profiles to Inorganics Techniques to Minimize Catalyst Deactivation Bibliography 123

12 LIST OF TABLES 2.1 Conversions and product selectivities (both in %) at 2.2 mm ( 95 % oxygen conversion) and 10.2 mm (end of catalyst) for (a) methane and (b) dimethyl ether Reactant conversions and product selectivities (both in %) during methane catalytic partial oxidation at C/O = 1 and total flow rate 5 SLPM Changes in product distributions upon doping with different inorganics during steam methane reforming Inorganic Content in Weight Percentage Dry Basis from Different Biomass Sources Summary of Effects Observed from Inorganics x

13 LIST OF FIGURES 2.1 Experimental setup for measurement of concentration and temperature profiles during methane and dimethyl ether catalytic partial oxidation Experimental setup for methane catalytic partial oxidation to determine oxidation state of the rhodium catalyst Methane, oxygen and temperature profile during methane catalytic partial oxidation Product concentration profiles during methane catalytic partial oxidation Dimethyl ether, oxygen and temperature profile during dimethyl ether catalytic partial oxidation Product concentration profiles during dimethyl ether catalytic partial oxidation Product flow rates at the end of 2.2 mm and 10.2 mm during (a) methane and (b) dimethyl ether catalytic partial oxidation XRD patterns for alumina, fresh (calcined rhodium) and used rhodium catalysts Methane conversion for catalysts doped with sulfur (A) and phosphorus (B) respectively. Hydrogen, carbon monoxide selectivities for catalysts doped with sulfur (C) and phosphorus (D) respectively. (A) also shows a schematic of the reactor setup Methane conversion for catalysts doped with potassium (A) and sodium (B) respectively. Hydrogen, carbon monoxide selectivities for catalysts doped with potassium (C) and sodium (D) respectively SEM images of (A) fresh 2.5 wt% Rh on α-al 2 O 3 catalyst (B) carbon filaments on catalyst doped with phosphorus (C) high resolution image of carbon filaments in catalyst doped with phosphorus showing rhodium particles at tip and (D) carbon structures on catalyst doped with potassium xi

14 LIST OF FIGURES xii 4.1 Doping and regeneration temperature profiles for Si (A). Doping and regeneration temperature profiles for S (B) Ethanol conversion, temperature profiles (A) and product selectivities (B) for doping with phosphorus Ethanol conversion, temperature profiles (A) and product selectivities (B) for doping with potassium Ethanol conversion, temperature profiles (A) and product selectivities (B) for doping with sodium Ethanol conversion, temperature profiles (A) and product selectivities (B) for doping with calcium Ethanol conversion and temperature profiles during doping and regeneration for magnesium Carbon monoxide selectivities during methane catalytic partial oxidation at 1 % and 10 % loading of potassium Methane and hydrogen selectivities during methane catalytic partial oxidation at 1, 10 and 100 % loading of phosphorus Methane conversion, carbon monoxide and hydrogen selectivities during methane catalytic partial oxidation at 1 and 10 % loading of monobasic potassium phosphate Carbon monoxide, hydrogen and ethylene selectivities during ethanol catalytic partial oxidation at 10 and 100 % loading of phosphorus Carbon monoxide, hydrogen and ethylene selectivities during ethanol catalytic partial oxidation at 10 and 100 % loading of monobasic potassium phosphate Methane conversions and temperatures during methane catalytic partial oxidation for 6 h at C/O ratios of 0.75 and 1.5 at 10 % loading of potassium (A) Methane conversions and temperatures during methane catalytic partial oxidation for 6 h at C/O ratios of 0.75 and 1.5 at 10 % loading of phosphorus. (B) Hydrogen selectivities over 6 h at C/O of 0.75 and (A) Ethanol conversions and temperatures during catalytic partial oxidation with ethanol containing 0.05 mol % potassium. (B) Carbon monoxide and acetaldehyde selectivities during the doping period (A) Ethanol conversions and temperatures during catalytic partial oxidation with ethanol containing 0.05 mol % phosphorus. (B) Carbon dioxide, hydrogen and ethylene selectivities during the doping period Reactor configuration for the autothermal CPO of the butanol isomers Conversion and catalyst backface temperature of the four butanol isomers Selectivities to H 2 and H 2 O from each butanol isomer

15 LIST OF FIGURES xiii 6.4 Selectivities to CO and CO 2 from each butanol isomer Selectivities to major carbonyl and C 4 olefins from each butanol isomer Selectivities to ethylene and propylene from each butanol isomer Proposed reaction mechanisms for each butanol isomer in the autothermal reactor Reactor configuration for autothermal catalytic partial oxidation of isobutanol Product distributions and temperatures from catalytic partial oxidation of isobutanol Product distributions and temperatures after integration of water gas shift with catalytic partial oxidation of isobutanol CO conversion in the water gas shift stage Thermodynamic analysis of the current system including photosynthesis, fermentation, catalytic partial oxidation-water gas shift, hydrogen use in fuel cell and combustion

16 CHAPTER ONE INTRODUCTION In this chapter, an overview is presented of the present global energy shortage and challenges associated with using biomass as an alternative to fossil fuels. Gasification, which is the principal topic of biomass upgrading in this thesis, involves producing synthesis gas (also called syngas), a mixture of carbon monoxide and hydrogen which can be upgraded through different routes to fuels and chemicals. Catalytic partial oxidation can convert different feedstocks to a high selectivity syngas stream autothermally over noble metal catalysts without any external heat input. Inorganics are an important constituent of lignocellulosic biomass and their mechanisms of altering catalyst activity are presented. 1.1 Current World Energy Scenario Energy is essential to sustain the growth and development of human life. The quality of life and standard of living in any civilization is greatly influenced by its energy consumption. 4 For example, there exists over an order of magnitude difference between the per capita energy consumption in North America and Africa. 5 The current world energy consumption is 14 TWH and approximately 80 % of it is supplied from fossil fuels such as coal, oil, and natural gas. 6,7 The world energy consumption is predicted to increase by approximately 35 % by the year The following sections describe the reasons for the present energy crisis and the why the current energy consumption model is unsustainable: 1

17 1.1 Current World Energy Scenario Limited Fossil Fuel Reserves Coal, oil, and natural gas have been the principal sources of energy for use in transportation, electricity generation, residential, and industrial sectors, supplying almost 80 % of the world s energy. 6,7 However, current estimates indicate that existing oil and coal reserves will be depleted to the point of being prohibitively expensive for exploration and recovery within approximately the next 50 and 120 years respectively. 6,9 Peak production of both is expected to occur in approximately the next 20 years. 10 Recently, unconventional sources of oil such as tar sands in Alberta, Canada have received attention. However, there exist concerns over the the high amount of energy required and environmental impact associated with extracting oil from such unconventional sources. 11 Current estimates for natural gas consumption indicate that reserves will last for approximately 60 years. 6 Natural gas use is predicted to increase with the discovery of large deposits of natural gas in unconventional sources. Over the past decade, large deposits of natural gas within shale formations in the US have been found resulting in gas production levels similar to the peak levels that existed during the 1970 s. 12 From 2000 to 2010, the percentage of natural gas from unconventional sources has increased from 1 % to 20 %. Natural gas represents a more clean energy source than coal and can be upgraded to hydrogen or liquid fuels. However, the environmental effects associated with use of chemicals for extracting natural gas from these unconventional sources such as shale formations by hydraulic fracturing are widely debated Developing countries The rapid increase in energy consumption by developing countries in Asia and Africa is another reason for the current energy crisis. Studies predict that over the next 25 years, the energy consumption of OECD nations (Organization of Economic Co-operation and Development countries, comprised mostly of North American and European Nations) will account for about 7 % of the global energy consumption increase while non-oecd countries (82 % of the current world population and 57 % of the current energy consumption 19 ) account for the remaining 93 % of the world energy consumption increase. 8 Of this increase, China and India alone account for more than 50 % of the predicted increase in global energy consumption.

18 1.2 Biomass as an Energy Source Climate Impact Perhaps the biggest driving force for reducing fossil fuel consumption and developing an energy infrastructure based on alternative energy sources is the issue of climate change associated with use of fossil fuels and subsequent global warming. Approximately 30 Gigatonnes of CO 2 are estimated to be released into the atmosphere every year from fossil fuel combustion. 20,21 As a consequence of rapid use of fossil fuels, CO 2 levels in the atmosphere have increased from 280 ppm before the industrial revolution to 390 ppm at present. 20 Increasing CO 2 levels may result in severe ecological and environmental impacts such as increased flooding, unpredictable weather patterns, and decreased agricultural productivity. 22,23 Hence, for a sustainable ecosystem it is necessary to reduce current CO 2 emissions and explore CO 2 neutral or non-carbon based sources of energy Energy Security Existing fossil fuel reserves are concentrated in few geographical locations around the world. In addition, political issues in those regions can disrupt prices of fossil fuels in the global market leading to economic uncertainty and energy insecurity. Developing local alternative sources of energy can reduce energy dependence on potentially unstable foreign nations. 1.2 Biomass as an Energy Source Sources of energy that are being considered as potential alternatives to fossil fuelbased energy sources are biomass, photovoltaic, hydroelectric, nuclear, electrochemical (batteries). The issues of safety associated with use of nuclear energy as evidenced in March 2011 by the reactor meltdown in Japan has led many countries to develop plans to shut down existing nuclear power plants. Each of the above sources of energy has associated drawbacks, which combined with favorable economics of fossil fuels, has prevented widespread application of these alternative energy technologies. It is therefore likely that a combination of alternative energy sources rather than a single one may be the most feasible means of replacing fossil fuels. Lignocellulosic biomass has potential as an alternative source of energy, as it is abundant and represents the only sustainable source of renewable carbon

19 1.2 Biomass as an Energy Source 4 Furthermore, CO 2 produced during the combustion of biofuels is consumed for the growth of biomass, preventing an increase in greenhouse gas emissions and making the overall process carbon neutral. A recent study conducted by the U.S. Department of Energy (DOE) and the U.S. Department of Agriculture showed that the U.S. has enough lignocellulosic biomass to sustainably produce biofuels replacing more than 30 % of its current petroleum consumption by 2030 if efficient biomass processing technologies exist Structure of Biomass Biomass has a complex structure and the composition of various components in biomass changes depending upon the biomass source. A major difference between composition of biomass and fossil fuels is the presence of high amounts of oxygen in biomass as compared with fossil fuels (O/C 1 for carbohydrate rich biomass, O/C 0 for crude oil). 29,30 As a result, the energy density of lignocellulosic biomass ( 8 MJ/kg) is much lower that that of gasoline ( 40 MJ/kg). 31,32 In general, any biomass source is comprised of the following components: Cellulose Cellulose (C 6 H 10 O 5 ) n (n ,000) is a polymer of glucose consisting of β-1,4 glycosidic linkages. 29,33,34 Cellulose is present in the form of microfibres held together by hydrogen bonding and Van Der Waals forces. Cellulose comprises of approximately wt % of dry weight of lignocellulosic biomass. 30,33 The structural arrangement and intermolecular hydrogen bonding contribute towards making the cellulose structure highly resistant to both acid and enzymatic hydrolysis. 29,35 Cellulose can exist in biomass in either crystalline or amorphous form, the crystalline form being more stable and resistant to hydrolysis. Hemicellulose Hemicellulose is a heteropolymer of different C 5 (xylose, arabinose) and C 6 sugars (glucose, galactose, mannose). 29,33,34 Hemicellulose comprises approximately wt % of the dry weight of woody lignocellulosic biomass and consists of fewer repeating units than cellulose (n 150). 30,33 Compared to celluose, hemicellulose can be more

20 1.2 Biomass as an Energy Source 5 easily hydrolyzed by acids or enzymes. Unlike cellulose, hemicellulose exists in the amorphous form due to its branched structure. 36 Lignin Lignin is a complex 3-D network of phenolic derivatives and comprises of approximately wt % of the dry weight of woody lignocellulosic biomass. 33,34 Lignin binds cellulose fibres together and helps in protecting plants from microbial attack. Extractives Extractives consists of components like fats, waxes, alkaloids, proteins, gums and resins. 34,36 Extractives can generally be removed from woody biomass using polar solvents such as methanol and water. Ash Ash or inorganics in biomass consists of various elements such as Si, Ca, K, P, Mg, Na, etc. present in different forms such as carbonates, chlorides, sulfates. 37,38 Their overall and individual concentration changes depending upon the biomass source Converting Biomass to Fuels and Chemicals The exact composition of individual components in lignocellulosic biomass can influence the type of processing technique being used or pretreatment technique being utilized. Currently there are three technologies being investigated for upgrading biomass to fuels and chemicals. Gasification Gasification involves heating biomass in the presence of steam or oxygen at high temperatures ( C) to produce syngas, a mixture of carbon monoxide and hydrogen. 36,39 Depending on the H 2 /CO ratio, different fuels and chemicals can be produced from syngas. Syngas is the building block for over 50 % of the petrochemicals synthesized worldwide. 40 Gasification can take place with or without a catalyst and can be carried out in a fluidized bed reactor or updraft/downdraft gasifiers. 36

21 1.2 Biomass as an Energy Source 6 Gasification generally results in the formation of tars and chars which can reduce performance efficiencies of downstream processes. 36,39 Concentrations of tars are reduced by employing downstream techniques such as filtration or scrubbing. 41 Pyrolysis Pyrolysis involves heating lignocellulosic biomass in the absence of oxygen in an inert atmosphere at temperatures between 375 and 550 C, followed by rapid cooling resulting in formation of a liquid product referred to as bio-oil. 36,39,42 Depolymerization of the biomass structure takes place during heating and along with bio-oil, solid char and non-condensable gases are also produced. Fast pyrolysis processes require residence times less than 1 s and bio-oil yields of up to 80 % (on a dry basis) are obtained. 42 Bio-oil is a mixture of over 400 compounds containing different functional groups such as aldehydes, ketones, acids, alcohols, polyols and aromatic phenols. 36,39,42 Bio-oil can be upgraded using catalysts by carrying out deoxygenation reactions to synthesize fuels and chemicals or to syngas by gasification. 36,43 Pretreatment and Hydrolysis Pretreatment followed by hydrolysis involves depolymerizing the biomass structure to produce monomer sugars from cellulose and hemicellulose. 36,39 Pretreatment results in removal of the lignin fraction and decreases crystallinity of cellulose. The sugar monomers obtained after hydrolysis can be upgraded by either fermentation (biological technique using enzymes) or using catalysts to produce fuels and chemicals. Each of these techniques has certain associated drawbacks. Gasification generally results in formation of tars and chars, which can lead to additional expensive cleanup costs. 36,39 During pyrolysis, bio-oil formed is acidic (ph 2-4), corrosive, toxic, and its composition changes with time. 32,36,39 Pretreatment steps associated with hydrolysis are expensive and if fermentation is the next upgrading step, it involves long residence times (24-48 h). 36 In case of ethanol production from corn starch, energy intensive distillation is required.

22 1.3 Upgrading Syngas for Energy Use Upgrading Syngas for Energy Use Syngas is an important intermediate for synthesis of chemicals (methanol, ammonia) and fuels. The following describe the various routes to upgrade syngas for producing energy Fischer-Tropsch Process The Fischer-Tropsch process involves upgrading syngas to hydrocarbon based fuels. The process typically takes place over iron or cobalt based catalysts at temperatures of C for low temperature Fischer-Tropsch (LT-FT) and C for high temperature Fischer-Tropsch (HT-FT). 44,45 Typical pressures are approximately bar for LT-FT and 25 bar for HT-FT. Several companies have either Fischer- Tropsch plants in operation or are in the process of being setup. 46 These technologies are mostly based on converting natural gas to syngas followed by upgrading through the Fischer-Tropsch process. The Fischer-Tropsch reaction may be used to synthesize higher hydrocarbons (Eq. 1.1) or mixed alcohols (Eq. 1.2) and may be represented as: 45 (2n + 1)H 2 + nco C n H 2n+2 + nh 2 O (1.1) 2 nh 2 + nco C n H 2n+1 OH + (n 1)H 2 O (1.2) Methanol Route Another route for upgrading syngas is through a methanol intermediate. Methanol, one of the top manufactured chemicals in the world, is synthesized from syngas over copper-zinc oxide based catalysts at temperatures of C and pressures of approximately bar. 47,48 Methanol synthesis is equilibrium limited and the reactions may be represented as: CO + 2 H 2 CH 3 OH H r = -91 kj/mol (1.3) CO H 2 CH 3 OH + H 2 O H r = -49 kj/mol (1.4)

23 1.4 Catalytic Partial Oxidation 8 Nobel Laureate George Olah advocated the existence of a future methanol-based economy as it has the potential to be a valuable intermediate for production of fuels and chemicals. 49 Methanol can be either converted to dimethyl ether (DME) or olefins/gasoline by the MTO/MTG (methanol to olefins/methanol to gasoline process) Both DME synthesis and MTO/MTG take place on acidic catalysts such as zeolites. DME is a clean burning fuel and can be used as a diesel substitute. 54, Hydrogen Generation for Fuel Cells Several studies have focused on the advantages of a hydrogen-based economy since hydrogen has a high energy density on a mass basis and can be used as an energy carrier By use of hydrogen in fuel cells, chemical energy can be converted into electrical energy for use in residential, industrial and transportation sectors producing only heat and water. The CO tolerances are different for various types of fuel cells. 59 PEM (proton exchange membrane) fuel cells are an attractive option for transporation, residence and commercial applications due to their high power densities. However, the catalyst in a PEM fuel cell is sensitive to CO poisoning, thereby requiring CO concentrations in the ppm range. Higher temperatures result in improved CO tolerances in PEM fuel cells. 60 Initial CO removal from syngas takes place through the water gas shift reaction (Eq. 1.5) shown below: CO + H 2 O CO 2 + H 2 H r = -41 kj/mol (1.5) To reduce the CO concentration to ppm levels, CO methanation (Eq. preferential CO oxidation (Eq. 1.7) are employed. 1.6) or CO + 3 H 2 CH 4 + H 2 O H r = -206 kj/mol (1.6) CO O 2 CO 2 H r = kj/mol (1.7) 1.4 Catalytic Partial Oxidation Schmidt and coworkers have used catalytic partial oxidation (CPO) has been used to convert different feedstocks to syngas. 1 3,61 72 The advantage of catalytic gasification

24 1.4 Catalytic Partial Oxidation 9 over the non-catalytic route is that higher feed conversions and selectivities to desired syngas products are attainable. CPO takes place at high temperatures ( C) over noble metal catalysts (Rh, Pt, Ru) with no tar or char formation, thereby reducing or eliminating cleanup costs. The residence times within the catalyst bed are on the order of milliseconds, thereby making it possible to have smaller reactor sizes as well as lower catalyst weights. Unlike fossil fuels, since biomass is a distributed resource, the cost of transporting biomass from its source to the processing location is an important economic consideration in implementing biomass as an alternative energy source. Due to the short residence times in CPO, it may be economically feasible to have small reactors at the biomass source and transport syngas produced to a central processing location Evolution of Research Hickman and Schmidt in 1993 first reported experiments on CPO. 61 Methane CPO was carried out autothermally over rhodium- and platinum-based catalysts. The overall partial oxidation reaction with methane is shown below (Eq. 1.8) CH O 2 2 CO + 2 H 2 H r = -36 kj/mol (1.8) Research over the next ten years involved CPO of higher alkanes (C 2 C 16 ) Processing higher alkanes which exist in the liquid state at room temperature (eg. decane C 10 H 22, hexadecane C 16 H 34 ) required a preheating section before the catalyst to vaporize the feed prior to contacting the catalyst In 2004, Deluga et al. showed that prevaporized ethanol could be converted to hydrogen with selectivities greater than 100 % over rhodium-based catalysts by steam addition. 68 Further research over the next 3 years focused on CPO of oxygenated hydrocarbons with different functional groups such as alcohols, acids, and esters In 2006, Salge et al. demonstrated the CPO of nonvolatile liquid feeds to syngas with high selectivities by Reactive Flash Volatilization (RFV). 1 Dauenhauer (2007) and Colby (2008) used RFV to convert solid feedstocks such as aspen, polyethylene, and cellulose to high selectivity synthesis gas stream autothermally without any external heat input. 2,3 Thus, CPO is a highly versatile and robust process capable of producing high yields of syngas from wide range of feedstocks, ranging from small molecules like methane to actual biomass, a complex network of polymers.

25 1.5 Inorganics and Biomass Processing Upgrading Products from Catalytic Partial Oxidation CPO is limited by thermodynamics over rhodium-based catalysts at high temperatures. 3,70,72,73 CPO products over rhodium-based catalysts typically consist of thermodynamic equilibrium products - CO, CO 2, H 2 and H 2 O. Also, no tar or char formation is observed because thermodynamics does not predict char formation at the high temperatures involved with the feedstocks being processed. 2,3 Since the contact times are in the order of milliseconds, these reactors run close to adiabatic conditions. As the product stream is at high temperatures, it can be upgraded to fuels (Section 1.3) and chemicals downstream using different chemistries without any additional heat input. 1.5 Inorganics and Biomass Processing The research on CPO to date has been performed using feedstocks that contain little to no inorganics with no observable catalyst deactivation. However, actual biomass, depending on the source can contain from less than 1 % to 25 % by weight inorganics. 38 Even if deactivation is taking place, it may not be observed since CPO reactors at high temperatures operate close to thermodynamic equilibrium. 2,3,70,73 Hence, due to thermodynamic equilibrium limitations, deactivation may be taking place without being observed on the time scale of experiments (tens of hours). CPO reactions over rhodium catalysts have typically been run for hours with pure feedstocks without any observable catalyst deactivation. It is possible that the catalyst will deactivate over longer time scales with even small concentrations of impurities in the feed. For example ppm levels of impurities in glycerol feed resulted in significant loss of catalyst activity over 400 h. 74 Inorganics in biomass present an important obstacle in any catalytic biomass upgrading process by changing the activity of the catalyst. Even after pyrolysis, bio-oil still contains up to 0.3 % inorganics which can influence subsequent catalytic upgrading processes. 32,33 Since noble metal catalysts are expensive, efficient use of the catalyst is essential. Hence, regeneration of the catalyst is critical to maintain economic feasibility of the process. The reactor configuration needs to be altered based on the time-scales of deactivation and regeneration. For example, a catalyst which deactivates within a few seconds, as is the case with an FCC zeolite catalyst, an entrained-type reactor

26 1.5 Inorganics and Biomass Processing 11 with riser is used. 75,76 For deactivation in the time scale of seconds or minutes, a swing type reactor may be used. For deactivation which takes place on the time scale of years, a fixed bed reactor is used. In general, catalyst deactivation can be attributed to any of the five following mechanisms: Poisoning Poisoning refers to deactivation of the catalyst due to strong chemisorption of impurity species on the catalyst surface. 75,77,78 Poisoning can be either selective or nonselective depending on whether the impurity binds specifically to certain catalyst sites. Through electronic interactions with catalyst active sites, poisons prevent adsorption of other species or can restructure the catalyst surface thereby altering catalyst activity. Sulfur is a common industrial catalyst poison for many processes (steam reforming, water gas shift, ammonia synthesis); hence sulfur removal is necessary to minimize catalyst deactivation Fouling Fouling refers to physical blocking of the active site by the impurity. 75,77 This is usually the deactivation mechanism observed with coke formation. Due to physical blocking of active sites, reaction species are unable to access them leading to changes in catalyst activity. In addition, large amounts of fouling may also affect heat transfer between catalyst particles which can alter reaction rates and product distributions Sintering Sintering involves deactivation of the catalyst due to aggregation of metal clusters at high temperatures to form larger aggregates. 75,77,78 Sintering is generally temperature dependent. By aggregating to form larger clusters, the surface area and hence surface energy is reduced thereby leading to a more energetically favored state. The metal support interaction and associated environment around the catalyst can also influence sintering. For example, in case of strong metal support interaction (SMSI), sintering is reduced. 79

27 1.6 Summary Attrition/Mechanical Failure This mechanism involves structural disintegration of the catalyst due to mechanical or thermal stresses. 75,77,78 Mechanical stresses may arise due to collision with other particles or at high flow velocities leading to erosion. This is common in fluidized bed reactors. Thermal stresses are encountered during rapid heating/cooling which may take place during startup/shutdown leading to structural disintegration Solid State Transformation This mode of deactivation can refer to the active catalyst material reacting with the support to form a new phase leading to loss of catalytic activity or the active catalyst reacting with species in the gas or vapor phase to form volatile compounds leading to loss of catalyst. 77,78 The nature of the environment around the catalyst and the temperature influence the transformation and change in catalytic activity. 1.6 Summary This thesis presents an insight into the mechanisms, process design, and challenges associated with CPO of various renewable feedstocks. Upgrading lignocellulosic biomass to syngas by CPO represents an attractive technology to partly supplant the energy supplied by fossil fuels. Chapter 2 investigates the mechanism of CPO by spatial profiles and characterization techniques such as X-ray diffraction (XRD) and X-ray photoelectron spectroscopy (XPS). Biomass inorganics can affect the activity of rhodium catalysts by a single or a combination of the any of the above five modes of deactivation. Chapters 3 and 4 survey the effects of common biomass inorganics. Chapter 5 studies the effect of potassium and phosphorus in detail at different temperatures and concentrations. To understand the mechanism by which individual inorganics affect the rhodium catalyst, along with performance testing, different characterization techniques such as Scanning Electron Microscopy (SEM), XPS, XRD, and chemisorption are used. To amplify catalyst deactivation so that activity changes could be observed on time scales feasible in a laboratory, experiments were performed at lower temperatures 700 C, such that the product distributions were away from thermodynamic equilibrium. At higher temperatures, the catalyst would be closer to thermodynamic equilibrium and the volatility of inorganics would increase thereby minimizing the ef-

28 1.6 Summary 13 fects of catalyst deactivation. Alcohols represent important intermediates in biomass upgrading processes such as pyrolysis and hydrolysis. 36,39 Chapter 6 investigates the CPO of four butanol isomers on Rh, RhCe, Pt, and PtCe catalysts. For ultilizing hydrogen in syngas for fuel cells, carbon monoxide removal is necessary. In Chapter 7, integration of CPO and water gas shift for production of high-purity hydrogen for fuel cells is examined.

29 CHAPTER TWO MECHANISM OF CATALYTIC PARTIAL OXIDATION Catalytic partial oxidation (CPO) reactors have complex concentration and temperature gradients which makes it difficult to understand the mechanism of autothermal operation. Previous research in the Schmidt group using the capillary sampling technique during methane CPO has shown the existence of two zones - oxidation and reforming Approximately all of the oxygen in the feed is converted within the first 2 mm of the catalyst bed. In this chapter, temperature and concentration profiles during the CPO of methane and dimethyl ether CH 3 OCH 3, an oxygenated hydrocarbon are presented. Results with the two feeds were compared and the nature of the profiles were similar with sharp gradients existing in the oxidation zone with slower chemistries in the reforming section. The oxidation state of the rhodium catalyst in the oxidation zone during methane CPO was also analyzed using XPS and XRD. 2.1 Introduction Catalytic partial oxidation (CPO) has been shown to convert a wide range of feedstocks ranging from small molecules such as methane to actual biomass autothermally to a high selectivity thermodynamic equilibrium synthesis gas stream. 1 3,61,63,84,85 The process takes place over noble metal based catalysts such as rhodium and platinum and typical temperatures in a CPO reactor are in the order of C with residence times approximately 50 ms. CPO involves a combination of oxidation and reforming reactions. CPO was first demonstrated by Hickman and Schmidt in

30 2.1 Introduction 15 and comprises of the following reactions involving combustion (Eq. 2.1), partial oxidation (Eq. 2.2), steam reforming (Eq. 2.3), CO 2 reforming (also called dry reforming, Eq. 2.4) and water gas shift (Eq. 2.5). 82 CH O 2 CO H 2 O H r = -803 kj/mol (2.1) CH O 2 CO + 2 H 2 H r = -36 kj/mol (2.2) CH H 2 O CO + 3 H 2 H r = 206 kj/mol (2.3) CH 4 + CO 2 2 CO + 2 H 2 H r = 247 kj/mol (2.4) CO + H 2 O CO 2 + H 2 H r = -41 kj/mol (2.5) Syngas stream can be upgraded to fuels or chemicals by the Fischer-Tropsch process, methanol, or hydrogen. Due to the short contact times (order of milliseconds), CPO is suitable for portable or small-scale applications like fuel cells unlike conventional steam reformers which have residence times in the order of seconds. 80 The capillary sampling technique has been demonstrated with simple hydrocarbons such as methane and ethane. 86 The noble metal catalyst during CPO can be divided into two sections - an oxidation section where oxygen is present and a reforming section downstream without any gas phase oxygen. Using the temperature and concentration profiles within the reactor, the mechanism of CPO on different catalysts can be determined. With oxygenated hydrocarbons, a problem with the technique is homogeneous chemistry within the reactor and sampling system which can introduce artifacts in the analysis. Kruger et al. showed that significant conversion of oxygenated hydrocarbons occur at temperatures in excess of 500 C which becomes more prominent in the presence of oxygen. 85,87 Dimethyl ether, CH 3 OCH 3 (DME) was shown to have negligible homogeneous chemistry even at temperatures in excess of 600 C indicating its high stability in the gas phase. Therefore, all chemistry can be attributed to that taking place on the catalyst surface. DME concentration and temperature profiles will help in better understanding the mechanism of CPO of oxygenates and hence of lignocellulosic biomass. As in methane CPO, complete

31 2.1 Introduction 16 oxidation (Eq. 2.6), partial oxidation (Eq. 2.7), CO 2 reforming (Eq. 2.8), steam reforming (Eq. 2.9) and water gas shift reactions (Eq. 2.5) take place during DME CPO. DME is considered as a diesel substitute and has lower NO x and particulate emissions than diesel fuel. 54,55 DME steam reforming as a source of hydrogen for fuel cells has been reported in the literature CPO of DME to produce high selectivity hydrogen represents an attractive option for small-scale or portable applications. CH 3 OCH O 2 2 CO H 2 O H r = kj/mol (2.6) CH 3 OCH O 2 2 CO + 3 H 2 H r = -38 kj/mol (2.7) CH 3 OCH 3 + CO 2 3 CO + 3 H 2 H r = 246 kj/mol (2.8) CH 3 OCH 3 + H 2 O 2 CO + 4 H 2 H r = 204 kj/mol (2.9) Initially, the interplay between the various reactions was unclear with both a direct and indirect mechanism of CPO being proposed The direct mechanism suggests that syngas products, H 2 and CO are formed directly from partial oxidation (Eq. 2.2) while the indirect mechanism suggests the existence of two distinct reaction zones - an oxidation zone producing combustion products (Eq. 2.1) and a reforming zone producing syngas (Eq. 2.3). Using the spatially resolved measurements by the capillary sampling technique, Schmidt and coworkers showed that all the oxygen is consumed within approximately 2 mm over Pt and Rh based catalysts and syngas production was observed even within the oxidation zone Reaction rates, concentration and temperature gradients were highest in the oxidation zone where oxygen was present in the gas phase ( 2 mm). However, as reported by Horn et al., the formation of syngas in the oxidation zone does not confirm the prevalance of a direct mechanism of CPO formation due to the occurrence of multiple surface reactions on the catalyst surface. 81 Detailed numerical simulation studies involving surface reactions taking place during methane CPO suggest that in short contact time reactors, at high temperatures, syngas formation in the oxidation zone takes place mostly by the direct mechanism along with syngas combustion, steam reforming and water gas

32 2.2 Experimental 17 shift reactions After all the oxygen is consumed, syngas is produced by steam reforming. During CPO, the net environment is highly reducing due to high amounts of hydrogen produced, hence rhodium exists in a reduced state in the reforming section. However, within the oxidation zone, where oxygen is present, the state of the rhodium catalyst is subject of debate in literature. Grunwaldt et al. showed using in situ X-ray absorption spectroscopy studies at lower temperatures ( 400 C) the presence of an indirect mechanism and in the oxidation zone, rhodium was present in the oxidized state. 98,99 Numerical simulations of short contact time reactors at temperatures similar to CPO ( C) suggest that rhodium is most likely to be in the reduced state even in the presence of oxygen Since CPO reactors operate at high temperatures, it is possible that the state of the catalyst might change during shutting down the reactor. Typical shut down of a CPO reactor involves setting oxygen followed by fuel flow rates to zero and subsequent cooling of the reactor with an inert gas. It is therefore possible that the catalyst in the oxidation zone might be oxidized during reaction, however the catalyst might get reduced due to pyrolysis of the fuel at high temperatures after shutting off the oxygen flow. To avoid any change to the catalyst during shutdown, both oxygen and methane flows were shut off simultaneously using an on-off valve. Here, XPS and XRD are used to experimentally analyze the oxidation state of the rhodium catalyst in the oxidation zone. 2.2 Experimental Spatial Profiles Methane and DME CPO was carried out in a 19 mm ID quartz reactor. The catalyst consisted of a 80 ppi (pores per linear inch) α-alumina foam monolith (17 mm diameter, 10 mm long) coated with 5 wt% rhodium. Blank 80 ppi α-alumina foams were used as front and back heat shields. The three monoliths (front heat shield, rhodium catalyst, and back heat shield) were wrapped in aluminosilicate cloth to prevent bypassing of gases. A 700 µm hole was drilled along the axis of all three monoliths with a diamond drill bit. Flow rates of gases to the reactor were controlled using mass flow controllers accurate to within ± 2%.

33 2.2 Experimental 18 Micrometer Stainless steel union To QMS/ Thermocouple Stainless steel guide Argon + Methane + Oxygen Quartz reactor tube Front heat shield Back heat shield Fused silica capillary 80 ppi 5 wt % Rh/α-Al 2 O 3 catalyst Exhaust Figure 2.1: Experimental setup for measurement of concentration and temperature profiles during methane and dimethyl ether catalytic partial oxidation.

34 2.2 Experimental 19 Rhodium nitrate was used as the rhodium precursor and catalysts were prepared by the incipient wetness technique as described previously, 100 then dried in a vacuum oven and calcined in a furnace in air at 800 C for 6h. Products were analyzed at the exit of the reactor for integral performance data and within front heat shield to test for homogeneous chemistry with a HP 5890 gas chromatograph with a 60 m Plot-Q column. For spatial profile studies, a fused silica capillary (650 µm OD) was connected to a 1/16 inch stainless steel union. The other end of the union was connected to a quadrupole mass spectrometer (QMS). The stainless steel union was mounted to a micrometer screw. The fused silica capillary moved along the axis of the reactor through the drilled holes within the 3 monoliths by turning the micrometer screw. Temperatures were measured by passing a thermocouple (Omega, K-type, 270 µm O.D.) through the union and the capillary, until the end of the thermocouple was slightly below the end of the capillary, similar to that shown by Horn et al. 80 The experimental setup is shown in Fig After turning the micrometer screw, the system was allowed to equilibrate for approximately 5 min. The positions shown are accurate within (0.3 mm). The profiles along the axis of the catalyst were analyzed with the QMS which has much faster analysis times than a gas chromatograph. Errors in carbon balances were typically within 10 %. The water flow rate was quantified by closing hydrogen and oxygen balances on it and averaging the two values. Selectivity to a particular species was defined on a carbon (for all carbonaceous species) or hydrogen (for H 2 and H 2 O) basis as number of atoms of carbon or hydrogen in a particular species to total number of carbon or hydrogen atoms in the products, not taking into account unconverted fuel. The carbon flow rate used in the experiments was approximately 25 mmol/min. The C/O ratio is defined as the ratio of carbon atoms in fuel to oxygen atoms from air. For spatial profile measurements of methane and DME, experiments were performed at a C/O ratio of 1.2 and at a total gas flow rate of 2.4 SLPM. This corresponds to a residence time of approximately 10 ms at reactor temperatures. The flow rate through the QMS sample line was approximately 5 ml/min which is negligible compared to the total flow rate through the system thereby causing negligible changes to flow behavior within the system.

35 2.2 Experimental 20 Argon On-off valve Methane + Oxygen Quartz reactor tube 80 ppi α-alumina monoliths 0.25 g 2.5 wt % Rh/α-Al 2 O 3 catalyst To GC and Incinerator Figure 2.2: Experimental setup for methane catalytic partial oxidation to determine oxidation state of the rhodium catalyst Nature of Rhodium in Oxidation Zone Experiments were carried out in a 19 mm ID quartz tube. Methane CPO was carried out at a total flow rate of 5 SLPM and a C/O ratio of 1. The methane flow rate was 1 SLPM. A 2.5 wt% rhodium catalyst was supported on 1.3 mm α-alumina spheres g of the catalyst was placed between two 80 ppi α-alumina foams. Another 80 ppi α-alumina foam was placed below the back heat shield and an Omega type-k thermocouple was placed between the two foam monoliths. The experimental setup is shown in Fig The monoliths were wrapped in aluminosilicate cloth to prevent bypassing of gases. Products were analyzed with a HP 5890 gas chromatograph as in the spatial profile experiments. CH 4 was quantified on the flame ionization detector while H 2, O 2, CO, and CO 2 were quantified using the thermal conductivity detector. Errors in carbon balances were typically within 10 %. Water flow rate was quantified by closing hydrogen and oxygen balances on it and averaging the two values. Experiments were run for one hour before methane and oxygen flows were shut off simultaneously using an on-off valve.

36 2.3 Results 21 Catalysts were analyzed by XRD and XPS. XRD was performed in Bruker D5005 with a 2.2 kw Copper Source. The spheres were crushed into a powder and then placed in the XRD. A step size of 0.02 and a dwell time of 1 s was used. XPS was performed in SSX-100 with 1.38 A.U. aluminium K-α radiation. The spot size used was 800 µm and the Carbon 1s peak at ev was used as a reference. 2.3 Results Spatial profiles Spatially resolved temperature and concentration profiles of methane and DME were measured during CPO using the capillary sampling technique. Up to the beginning of the catalyst (end of front face), no fuel or oxygen conversion was observed within the front heat shield on the gas chromatograph indicating absence of homogeneous chemistry with both methane and DME. Spatial profiles of methane showed that most of the methane conversion and almost all of the oxygen ( 95 %) consumption takes place within the first 2.2 mm of the catalyst as shown in Fig The temperature profile initially increased due to exothermic oxidation chemistry and then decreased due to endothermic reforming. The nature of the trends are consistent with previously observed results with spatial profiles of methane Syngas formation was observed from the very beginning of the catalyst (Fig. 2.4a and Fig. 2.4b). Sharp gradients in all product distributions were observed in the oxidation zone. The water concentration profile (Fig. 2.4a) showed a peak at the end of the oxidation zone and then decreased due to steam reforming reaction (Eq. 2.3). The hydrogen profile increased throughout the length of the catalyst (Fig. 2.4a) while the CO concentration profile was almost constant after about 3.5 mm (Fig. 2.4b), just after all of the oxygen was consumed. The CO 2 concentration profile was almost constant after 2mm indicating absence of CO 2 reforming (Eq. 2.4) over Rh catalysts consistent with previous results (Fig. 2.4b) Similar trends were observed during spatial profiles of DME. DME being more reactive than methane, higher overall conversion of DME was observed and 95 % of the oxygen was consumed within the first 2.2 mm (Fig. 2.5). Similar to methane, syngas formation for DME was observed within the oxidation zone (Fig. 2.6a and

37 2.3 Results 22 F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld 5 w t % R h c a ta ly s t C H 4 B a c k h e a t s h ie ld O 2 A x ia l C o -o rd in a te (m m ) T T e m p e ra tu re ( o C ) Figure 2.3: Methane ( ), oxygen ( )and temperature profile ( ) during methane CPO. Fig. 2.6b). For both methane and DME, 99.9 % oxygen consumption was observed within 3.5 mm. CO production like during methane CPO was observed primarily in the oxidation zone with its profile almost constant after 3.5 mm as shown in Fig. 2.6b. Hydrogen concentration profile increased throughout the length of the catalyst (Fig. 2.6a) similar to methane CPO. Comparison of the H 2 and CO concentration profiles during methane and DME CPO revealed that the profiles were steeper during DME CPO than with methane, which may be attributed to the higher overall reactivity of DME. Comparing the water evolution profiles, the initial profile was much steeper with a sharper peak for methane compared with DME. Examination of the CO 2 concentration profile showed that the rate of CO 2 increase was slower with DME (Fig. 2.6b) than with methane (Fig. 2.4b). In case of methane, the CO 2 profile was constant after approximately 2 mm. With DME, the rate of CO 2 increase was much slower, gradually increasing up to 3.5 mm and was almost constant thereafter. Thus, compared to methane, in case of DME, CO 2 evolution takes place almost to the point of 99.9 % oxygen conversion (3.5 mm). This may be due to CO 2 reforming of DME, which caused the CO 2 evolution to be less steep as CO 2 was consumed by

38 2.3 Results w t % R h c a ta ly s t 1 4 F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld H 2 H 2 O B a c k h e a t s h ie ld A x ia l C o -o rd in a te (m m ) (a) 5 w t % R h c a ta ly s t 1 2 F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld C O C O 2 B a c k h e a t s h ie ld A x ia l C o -O rd in a te (m m ) (b) Figure 2.4: Product concentration profiles during methane CPO. (a) shows hydrogen( ) and water( ) concentration profiles. (b) shows carbon monoxide ( ) and carbon dioxide( ) concentration profiles.

39 2.3 Results w t % R h c a ta ly s t T F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld C H 3 O C H 3 B a c k h e a t s h ie ld O 2 A x ia l C o -o rd in a te (m m ) T e m p e ra tu re ( o C ) Figure 2.5: DME ( ), oxygen ( ) and temperature profile ( ) during DME CPO. DME dry reforming. Due to competing pathways of CO 2 and steam reforming of DME, CO 2 evolution was spread out compared to methane (2 mm vs 3.5 mm) and water concentration profile showed a less steep rise and less sharp peak compared to methane. After 3.5 mm, during DME CPO, due to reduced temperatures, CO 2 reforming does not appear to play a role. Ma et al. reported 100 % conversion of DME and 80 % conversion of CO 2 during dry reforming of DME over Ni-based catalysts at temperatures similar to that involved in the current system at lower space velocities. 101 Further experiments are necessary to gain a detailed understanding of the role of dry reforming over Rh-based catalysts during DME CPO. Temperature profile trends were nearly identical in nature for DME and methane showing an initial steep rise due to highly exothermic oxidation chemistry and subsequent decrease due to endothermic reforming chemistry after all the oxygen was consumed (Fig. 2.3 and Fig. 2.5). Though most of the oxygen was consumed within 2 mm, the temperature profiles showed a peak at about 3.2 mm from the front face of the catalyst due to the gas and catalyst surface not being in thermal equilibrium. 80,82 Temperatures during DME CPO were higher than methane CPO throughout the

40 2.3 Results w t % R h c a ta ly s t H F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld H 2 O B a c k h e a t s h ie ld A x ia l C o -o rd in a te (m m ) (a) 5 w t % R h c a ta ly s t C O F lo w ra te (m m o l/m in ) F ro n t h e a t s h ie ld B a c k h e a t s h ie ld 0-2 C O A x ia l C o -o rd in a te (m m ) (b) Figure 2.6: Product concentration profiles during DME CPO. (a) shows hydrogen( ) and water( ) concentration profiles. (b) shows carbon monoxide ( ) and carbon dioxide( ) concentration profiles.

41 2.3 Results m m 2.2 m m m m m m 2.2 m m m m F lo w ra te (m m o l/m in ) F lo w ra te (m m o l/m in ) H 2 C O H 2 O C O C H H 2 C O H 2 O C O 2 D M E (a) CH 4 (b) DME Figure 2.7: Product flow rates at the end of 2.2 mm and 10.2 mm during (a) methane and (b) dimethyl ether catalytic partial oxidation. Species 2.2 mm 10.2 mm S H S CO S CO S H2 O X CH (a) Species 2.2 mm 10.2 mm S H S CO S CO S H2 O X CH3 OCH Table 2.1: Conversions and product selectivities (both in %) at 2.2 mm ( 95 % oxygen conversion) and 10.2 mm (end of catalyst) for (a) methane and (b) dimethyl ether. (b) catalyst by 50 C due to the higher reactivity of DME. The flow rates of syngas, combustion products and feed (DME or methane) during methane and DME CPO at various stages along the catalyst are shown in Fig mm corresponds to the point where for both methane and DME CPO, 95 % oxygen conversion was observed. The product selectivities and conversions for methane and DME CPO are shown in Tables 2.1a and 2.1b: Based on the data in Table (2.1a) and (2.1b), the overall equations at 2.2 mm (Eq. 2.10) and 10.2 mm (Eq. 2.11) for methane CPO may be written as:

42 2.3 Results 27 CH O CO CO H H 2 O H r = -445 kj/mol (2.10) CH O CO CO H H 2 O H r = -325 kj/mol (2.11) Similarly, for DME CPO (neglecting CH 4 and C 2 H 6 as their selectivities are small, 1%), at 2.2 mm (Eq. 2.12) and 10.2 mm (Eq. 2.13): CH 3 OCH O 2 1.8CO + 0.2CO H H 2 O H r = -529 kj/mol (2.12) CH 3 OCH O CO CO H H 2 O H r = -457 kj/mol (2.13) The above equations highlight the differences in CO 2 concentrations at 2.2 mm and 10.2 mm during methane and DME CPO. The less exothermic heat of reaction at 10.2 mm compared to 2.2 mm is due to the presence of endothermic reforming chemistry between 2.2 mm and 10.2 mm State of Rhodium in Oxidation Zone The product distribution obtained during methane CPO at a total flow rate of 5 SLPM with C/O = 1 and Ar/O 2 ratio corresponding to air stoichiometry (79/21) is shown in Table 2.2. The temperature below the back heat shield was 635 C and oxygen breakthrough was observed ( 62 % oxygen conversion). XRD and XPS measurements were used to verify the oxidation state of the rhodium catalyst in the oxidation zone. With the used rhodium catalyst, XPS measurements showed the presence of rhodium in the reduced (metallic state) with 3d 5/2 and 3d 3/2 binding energies of and ev respectively. The 3d 5/2 and 3d 3/2 binding energies of the fresh (after calcination) 2.5 wt % rhodium catalyst were 308.3

43 2.3 Results 28 Species Exit value S H S CO 60.4 S CO S H2 O 58.6 X CH X O Table 2.2: Reactant conversions and product selectivities (both in %) during methane catalytic partial oxidation at C/O = 1 and total flow rate 5 SLPM. and ev respectively indicating presence of rhodium in the oxidized state. Values reported represent the average of XPS scans over two spheres. The XRD pattern of blank alumina spheres, fresh (after calcination) 2.5 wt % Rh spheres and used 2.5 wt % Rh spheres are shown in Fig With the used rhodium catalyst, XRD measurements showed the absence of rhodium oxide and the presence of metallic rhodium peaks at 2 theta values of 41, 48, and 70 degrees. Fresh 2.5 wt % Rh spheres showed the presence of rhodium oxide during XRD. These experiments show that despite the presence of high temperatures and oxygen within the oxidation zone, rhodium is present in the reduced state. This contradicts the results of Grunwaldt et al. who reported through in situ X-ray absorption studies that rhodium is present as rhodium oxide in the oxidation zone. 98,99 However, those experiments were performed at much lower temperatures ( 400 C). At the high surface temperatures within the oxidation zone in a typical CPO reactor ( C), 82,86 the kinetics of different surface reactions are much faster. Previous research by Horn et al. has shown that the CPO process is mass transfer limited. 81 This suggests that the kinetics of reactions involving oxygen containing species are very fast compared to reaction of oxygen with rhodium to give rhodium oxide. As a result, the surface concentration of oxygen containing species is very low. This has been shown to be true in simulation studies of methane CPO at comparable temperatures which have shown that OH is the main oxidizer species in the oxidation zone and its coverage is in the order of 0.01 in the oxidation zone. 95,96 This shows that within the oxidation zone, the reaction takes place by Regime II as described by Maestri et al. 95 which involves primarily direct mechanism of syngas formation along with syngas combustion and water gas shift, and concentration of oxygen on the catalyst surface is almost zero. Spatially resolved concentration and temperature profiles were similar for methane

44 2.3 Results 29 R h p e a k R h a g e d R h 2 O 3 p e a k In te n s ity (a.u.) R h fre s h b la n k α-a l 2 O θ (d e g re e s ) Figure 2.8: XRD patterns for alumina, fresh (calcined rhodium) and used rhodium catalysts. Peaks are normalized to the largest α-alumina peak at 43 degrees.

45 2.4 Conclusion 30 and DME implying similar mechanism for CPO for different molecules. In the absence of surface oxygen, as is likely the case in the current experiments, it has been shown that after adsorption of DME, dehydrogenation steps take place to form CH x OCH x which ultimately form syngas, CO and H 2 and desorb. 85,102 It may therefore be expected that similar to methane, CPO of DME also proceeds primarily through a direct mechanism involving DME dehydrogenation (pyrolysis), syngas combustion, and water gas shift reactions within the oxidation zone. 2.4 Conclusion Spatially resolved concentration and temperature profiles for methane and DME CPO were obtained using the capillary sampling technique. With both feedstocks, 95 % oxygen was consumed within approximately 2.2 mm of the beginning of the catalyst. Syngas formation was observed even within the oxidation zone. Largest temperature and concentration gradients were observed within the oxidation zone. Dry reforming appeared to play a role during DME CPO while no evidence of it was observed during methane CPO. Further experiments are necessary to gain a better understanding of the role of dry reforming during DME CPO. XPS and XRD measurements showed that even in the presence of oxygen, within the oxidation zone, rhodium was present in the reduced state as metallic rhodium indicating within the oxidation zone, syngas formation takes place mostly by the direct mechanism. 2.5 Acknowledgements The author would like to acknowledge the Minnesota Corn Growers Association for financial Support. The author thanks Dr. Jacob Kruger and Samuel Blass for valuable discussions regarding experiments performed in this chapter.

46 CHAPTER THREE EFFECTS OF BIOMASS INORGANICS ON RHODIUM CATALYSTS: I. STEAM METHANE REFORMING 1 In Chapter 2, the complex concentration and temperature gradients existing during catalytic partial oxidation were presented. In this chapter, the effects of inorganics on rhodium catalysts were studied under controlled isothermal conditions, by depositing the inorganics directly on the catalyst and using steam methane reforming at 700 C as a model system. Inorganic elements in biomass present a major challenge for large-scale application of catalytic processing. Rhodium-based catalysts have been shown to gasify lignocellulosic biomass to syngas with high selectivities by reactive flash volatilization. In this research, the effect of inorganics commonly found in biomass, Na, K, Ca, Mg, Si, P, and S, on rhodium catalysts was investigated using steam methane reforming (SMR) as a model system. SMR was carried out at 700 C on a heated fixed bed of 2.5 wt % Rh/α-Al 2 O 3 catalysts. The inorganics were uniformly deposited on the catalyst (1 inorganic atom/5 rhodium atoms), followed by performance testing and characterization through H 2 chemisorption, SEM, XPS and XRD. Phosphorus, potassium and sulfur decreased the methane conversion the most (> 15 %) among the inorganics studied. No significant deactivation was observed upon doping with calcium, magnesium and silicon. Phosphorus and sulfur strongly reduced the dispersion 1 Parts of this chapter appear in Reetam Chakrabarti, Joshua L. Colby, Lanny D. Schmidt, Effects of Biomass Inorganics on Rhodium Catalysts: I. Steam Methane Reforming, Applied Catalysis B : Environmental 107 (2011) c 2011 Elsevier B.V. Reproduced with permission from Elsevier. 31

47 3.1 Introduction 32 of the rhodium catalyst. Addition of phosphorus and potassium caused formation of carbon-based structures, and phosphorus also increased the rhodium binding energies by 0.6 ev in the XPS spectrum, indicating rhodium oxidation. 3.1 Introduction Fossil fuels such as coal, oil and natural gas have been a main energy source for more than a century. These fossil fuels are carbon-based and have a high energy density. 103 Lignocellulosic biomass has potential as an alternative source of energy and represents a sustainable source of renewable carbon to produce fuels and chemicals Gasification is a common technique for processing biomass to fuels and chemicals. It produces syngas, a mixture of hydrogen and carbon monoxide that can be upgraded to diesel fuel using the Fischer-Tropsch process or to methanol. Schmidt et al. recently established the reactive flash volatilization technique by which lignocellulosic biomass can be gasified autothermally to a high-selectivity syngas product stream free from tars and chars in millisecond contact times. 1 3 This process takes place in an oxygen-deficient environment at high temperatures ( C) over rhodium-based catalysts. One of the major challenges in catalytic gasification is the presence of inorganics in biomass which may alter the catalyst activity and process chemistry. The effect of these inorganics on rhodium catalysts in reactive flash volatilization has not yet been investigated in detail. Inorganic elements present in biomass are Na, K, Ca, Mg, Mn, P, Si, Cl, S, Fe, Al and small concentrations of heavy metals. 37,104 Their quantities and concentrations vary depending on the type of biomass. For example, hardwoods and softwoods contain approximately 1 wt % inorganics whereas most annuals contain approximately 3-10 % and rice hulls contain up to 25 %. 37 Understanding the nature of the interactions of inorganics with rhodium is necessary for efficient processing of biomass by catalytic gasification. During reactive flash volatilization of lignocellulosic biomass, inorganics can alter the activity of the catalyst by any combination of three different phenomena: (a) accumulation of fly-ash near the front face of the catalyst which affects mass transfer of reactants and inorganics as well as heat transfer in the system, (b) poisoning due to strong chemisorption on active sites and (c) fouling or physical blocking of the active sites. Mechanisms (a) and (c) may be expected to be more predominant with less volatile inorganics such

48 3.1 Introduction 33 as Ca and Mg. The objective is to understand the interactions of the inorganics with the active sites due to mechanisms (b) and (c). Direct introduction of inorganics in the feed makes it difficult to decouple mechanism (a) from mechanisms (b) and (c). Therefore, the inorganics were deposited uniformly throughout the catalyst prior to reaction. This system is not representative of how actual biomass-derived inorganics would ordinarily reach the catalyst. However, by eliminating the effect of the inorganics being filtered out near the front face, it is possible to attribute the deactivation to mechanisms (b) and (c). Steam methane reforming (SMR) was chosen as a model system to study the effect of inorganics on the rhodium-based catalyst. The primary reactions in SMR are: CH 4 + H 2 O CO + 3 H 2 (3.1) CO + H 2 O H 2 + CO 2 (3.2) SMR represents a controlled and well-defined system since it is kinetically limited 105,106 and the conversion of methane can be related to the activity of the rhodium catalyst. Also, methane exhibits very little homogeneous chemistry at the temperatures involved. 107 Therefore, the change in conversion of methane can be directly related to the change in activity of the rhodium catalyst introduced upon addition of inorganics. By performing SMR over undoped and inorganic-doped rhodium catalysts, the effect of inorganics on the rhodium sites involved in reforming to produce syngas can be studied. It has been shown that the catalyst in reactive flash volatilization consists of three sections- volatilization, oxidation and reforming. 3 Reforming, the downstream section of the catalyst in reactive flash volatilization, involves water gas shift and steam reforming of hydrocarbons to give a product stream with high selectivities to syngas. SMR simulates the reforming zone in reactive flash volatilization as it involves similar types of reactions. Also, by using methane as a probe, homogeneous chemistry in the system can be neglected. 107 The effect of common biomass inorganics was examined by adding them through aqueous precursors to the rhodium-based catalyst at a concentration of 1 atom of inorganic for every 5 atoms of rhodium. The effect of these inorganics was studied by performance testing for both undoped and doped catalysts for SMR at 700 C on 2.5 wt % Rh/α-Al 2 O 3 catalysts. Catalysts were characterized by H 2 pulsed chemisorp-

49 3.2 Experimental 34 tion, SEM, XRD and XPS before and after doping to study dispersion, microstructure, crystallite formation and electronic interactions respectively upon introduction of inorganics. In Chapter 4, inorganics were introduced in the feed for reactive flash volatilization to study their effects on rhodium catalysts. 108 In this study, we directly deposit them on the catalyst; which eliminates any deactivation that may occur by accumulation of inorganics near the front face of the catalyst when introducing them in the feed. Both techniques are necessary to gain a detailed understanding of the interaction of inorganics on rhodium catalysts for biomass processing applications. 3.2 Experimental Experimental Setup SMR was carried out in a quartz reactor (20 mm I.D.) over 3 g of 2.5 wt % Rh supported on 1.3 mm diameter α-al 2 O 3 spheres. A blank 80 ppi α-al 2 O 3 monolith (17 mm diameter, 10 mm long) was used to support the catalyst bed (Fig. 3.1A inset). Two blank 45 ppi α-al 2 O 3 monoliths were placed about 1cm upstream of the 80 ppi monolith which acted as static mixers. The monoliths were held against the reactor walls by wrapping them with aluminosilicate cloth. Mass flow controllers regulated the flow rate of gases (N 2, H 2, CH 4 ) to the reactor, accurate to within ± 2 %. Water was fed using a liquid handling pump through a heated coil maintained at approximately 200 C to generate steam. A type-k thermocouple was placed between the two 45 ppi monoliths and a benchtop temperature controller (Omega CSC 32) controlled the temperature between the monoliths to within ± 1 C. The entire reactor tube was enclosed within a tube furnace. Reforming being an endothermic process, this arrangement helped to insulate the reactor from the heat effects of the reaction Product Analysis Products were analyzed with an Agilent 6890 Gas Chromatograph equipped with a 30 m Plot-Q column and both thermal conductivity and flame ionization detectors. Water was quantified by closing the hydrogen and oxygen balances and averaging the two results. During operation, the carbon, hydrogen and oxygen balances generally closed to within ± 5 %.

50 3.2 Experimental 35 Selectivities to CO and CO 2 were calculated on a carbon basis and to H 2 and H 2 O were on a hydrogen basis. Selectivity to a product is defined as (atoms in product species)/(atoms in converted fuel). Steam fed to the reactor was not considered fuel. The sum of selectivities to products on carbon or hydrogen basis was 100 %, within the limits of experimental error Catalyst Preparation Catalysts used in all the experiments were 2.5 wt % Rh on α-al 2 O 3 spheres (1.3 mm diameter, Saint Gobain Norpro). Catalysts were prepared by the incipient wetness technique [8]. Rhodium nitrate solution (Sigma-Aldrich) was used as the rhodium precursor. Catalysts were dried, then calcined in a furnace at 800 C for 6 h Experimental Procedure Each catalyst was initially aged at 850 C for 3 h in a mixture of 20 % steam, 10 % methane and 70 % nitrogen at a total flow rate of 2 SLPM. Subsequently, the temperature was reduced to 700 C and the performance of the catalyst for SMR was measured. The residence time within the catalyst bed was approximately 20 ms at 700 C. The inorganic species of interest was added to the catalyst (one inorganic atom for every five rhodium atoms) through its precursor by the incipient wetness technique and its performance was compared with the undoped catalyst at 700 C Catalyst Characterization SEM images were taken on JEOL 6700 equipped with a secondary electron detector. Most samples were coated with 100 A.U. of carbon to reduce charging effects. XRD was carried out on a Bruker D 5005 diffractometer equipped with a 2.2 kw sealed Cu source. Spheres were crushed to a fine powder increasing the peak intensities during analysis. XPS studies were carried out on Surface Science SSX 100 with a monochromatic Al K-α source ( ev). The spot size used in the measurements corresponded to 800 m. All peaks were calibrated with respect to the C 1s peak at 285 ev. H 2 pulsed chemisorption was carried out by injecting pulses of hydrogen from a 5 % hydrogen in argon mixture into a quartz tube containing 0.2 g of the catalyst

51 3.3 Results 36 sample. 15 pulses were injected into the catalyst at 1 minute intervals at 25 C with a dwell time of 15 s. The outlet was connected to a thermal conductivity detector. 3.3 Results The effect of inorganics on rhodium was studied by measuring the reactor performance for 5 h and through characterization by H 2 chemisorption, SEM, XPS and XRD before and after doping with inorganics Performance Testing All of the undoped catalysts showed similar baseline performance with about 67 % methane conversion. Hydrogen and carbon monoxide selectivities were about 68 % and 49 % respectively. These parameters were used as an indicator of reactor performance and plotted with time for each inorganic during the 5 h operation period. The changes in conversions and selectivities are expressed in terms of the absolute difference between the average doped (over 5 h) and undoped values. For example, if the conversion decreased from 60 % to 30 %, this is equivalent to (60-30) = 30 % and not 50 % decrease in conversion. Sulfur Dimethyl sulfoxide was used to add sulfur to the catalyst. Sulfur decreased the methane conversion the most among all the inorganics studied. However, partial regeneration of the catalytic activity was observed during the 5h test period. Methane conversion (Fig. 3.1A) and hydrogen selectivity increased (Fig. 3.1B) with time from 19 % and 32 % respectively at the beginning to 31 % and 42 % at the end of 5 h. The CO selectivity increased by about 10 % and was almost constant throughout the 5h duration (Fig. 3.1B). Phosphorus Ammonium phosphate was used as a precursor for introducing phosphorus to the catalyst. Phosphorus decreased the methane conversion by about 17 % (Fig. 3.1C). Also, CO selectivity increased by 22 % whereas H 2 selectivity decreased by 13 % (Fig.

52 3.3 Results D). After the 5 h performance testing period, a significant fraction of the spheres appeared black due to coke formation. Silicon Tetraethylorthosilicate (TEOS) was added to the aged rhodium catalyst to introduce silicon. The changes in methane conversion and product selectivities (not shown) were within the limits of experimental error. Sodium and Potassium Acetate precursors were used to add sodium and potassium to the catalyst. Sodium and potassium decreased methane conversions by 9 % and 16 %, respectively (Fig. 3.2, A and C). Potassium decreased the H 2 and CO selectivities by 5 % and 7 % respectively, whereas the decrease with sodium was within experimental error as shown in Fig. 3.2, B and D. Calcium and Magnesium Calcium acetate and magnesium acetate tetrahydrate were used as calcium and magnesium precursors. Conversions and syngas selectivities were unchanged within limits of experimental error for both calcium and magnesium (not shown). The results of performance testing are summarized in Table Catalyst Characterization H 2 chemisorption The effect of the inorganics on the dispersion of rhodium on the α-al 2 O 3 support was analyzed by H 2 pulsed chemisorption. All the aged rhodium catalysts had a rhodium dispersion of approximately 10 %. Phosphorus and sulfur decreased the dispersion of the catalyst (10.7 % to 3.6 % for phosphorus, 9.5 % to 4.4 % for sulfur). For the other inorganics, the change in dispersion was within the limits of experimental error. SEM Images were taken for each of the catalysts studied. Fresh, aged and Na, Ca, Mg, Si, and S doped catalysts appeared similar in terms of morphology, each showing

53 3.3 Results A S u lfu r C P h o s p h o r u s M e th a n e c o n v e rs io n (% ) T h e rm o c o u p le P ro d u c ts to G C a n d in c in e ra to r R h /α A l O c a ta ly s t p p i fo a m 4 5 p p i fo a m s N itro g e n, M e th a n e, S te a m R h d o p e d w ith S U n d o p e d R h T im e (m in ) B H R h d o p e d w ith P U n d o p e d R h T im e (m in ) D C O S e le c tiv ity (% ) H 2 C O C O s e le c tiv ity (S d o p e d ) C O s e le c tiv ity (u n d o p e d ) H 2 s e le c tiv ity (S d o p e d ) H 2 s e le c tiv ity (u n d o p e d ) T im e (m in ) H T im e (m in ) C O C O s e le c tiv ity (P d o p e d ) C O s e le c tiv ity (u n d o p e d ) H 2 s e le c tiv ity (P d o p e d ) H 2 s e le c tiv ity (u n d o p e d ) Figure 3.1: Methane conversion for catalysts doped with sulfur (A) and phosphorus (B) respectively. Hydrogen, carbon monoxide selectivities for catalysts doped with sulfur (C) and phosphorus (D) respectively. (A) also shows a schematic of the reactor setup.

54 3.3 Results A P o ta s s iu m 7 0 C S o d iu m M e th a n e c o n v e rs io n (% ) R h d o p e d w ith K U n d o p e d R h R h d o p e d w ith N a U n d o p e d R h S e le c tiv ity (% ) T im e (m in ) C O B H 2 C O s e le c tiv ity (K d o p e d ) C O s e le c tiv ity (u n d o p e d ) H 2 s e le c tiv ity (K d o p e d ) H 2 s e le c tiv ity (u n d o p e d ) T im e (m in ) T im e (m in ) C O D T im e (m in ) H 2 C O s e le c tiv ity (N a d o p e d ) C O s e le c tiv ity (u n d o p e d ) H 2 s e le c tiv ity (N a d o p e d ) H 2 s e le c tiv ity (u n d o p e d ) Figure 3.2: Methane conversion for catalysts doped with potassium (A) and sodium (B) respectively. Hydrogen, carbon monoxide selectivities for catalysts doped with potassium (C) and sodium (D) respectively.

55 3.3 Results 40 Inorganic - X CH4 (%) S CO (%) S H2 (%) S(t=0) S(t=5h) P Si Na K Mg Ca Table 3.1: Changes in methane conversion ( X CH4 ), carbon monoxide ( S CO ) and hydrogen selectivities ( S H2 ) at 700 C upon doping with different inorganics (1 atom inorganic/ 5 atoms of rhodium) as compared to aged undoped catalyst (2.5 wt % Rh/α-Al 2 O 3 ). Average values of methane conversion, carbon monoxide and hydrogen selectivities were 67 %, 49 % and 68 % whereas equilibrium values at 700 C were 99 %, 73 % and 81 % respectively. 109 rhodium particles around 50 nm in diameter (Fig. 3.3A). Carbon formation was observed on the surface of catalysts doped with phosphorus and potassium. In the case of phosphorus-doped catalyst, the carbon formed was filamentous (Fig. 3.3B) with rhodium particles at the tips (Fig. 3.3C), whereas the carbon formed with potassium-doped catalyst showed the presence of needle-like structures (Fig. 3.3D). After placing the phosphorus-doped catalyst in a furnace at 500 C for 30 min, the filamentous carbon on the catalyst was almost completely eliminated. The carbon on the potassium-doped catalyst remained even on heating at 750 C for 30 min. XPS The electronic interactions of the inorganics with rhodium were analyzed by XPS. The aging process decreased the rhodium binding energies of 3d 5/2 and 3d 3/2 peaks by about 1 ev showing conversion of rhodium oxide (Rh 2 O 3, Rh +3 ) to rhodium metal (Rh 0 ). Sodium, calcium, magnesium and silicon lowered the binding energy of rhodium by about 0.3 ev, whereas potassium decreased it by about 0.6 ev compared to the aged rhodium catalyst. Sulfur addition caused no change whereas phosphorus increased the rhodium binding energies compared to the aged rhodium catalyst by approximately 0.6 ev.

56 3.3 Results 41 Figure 3.3: SEM images of (A) fresh 2.5 wt% Rh on α-al 2 O 3 catalyst (B) carbon filaments on catalyst doped with phosphorus (C) high resolution image of carbon filaments in catalyst doped with phosphorus showing rhodium particles at tip and (D) carbon structures on catalyst doped with potassium.