SrCeO3-BASED PROTONIC CONDUCTORS FOR HYDROGEN...

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SrCeO3-BASED PROTONIC CONDUCTORS FOR HYDROGEN PRODUCTION AND SEPARATION BY WATER GAS SHIFT, STEAM REFORMING, AND CARBON DIOXIDE

REFORMING REACTIONS

By

JIANLIN LI

A DISSERTATION PRESENTED TO THE GRADUATE SCHOOL OF THE UNIVERSITY OF FLORIDA IN PARTIAL FULFILLMENT

OF THE REQUIREMENTS FOR THE DEGREE OF DOCTOR OF PHILOSOPHY

UNIVERSITY OF FLORIDA

2009

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© 2009 Jianlin Li

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To my grandmother, parents and sisters and friends who encouraged and supported me in good times and bad

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ACKNOWLEDGMENTS

The completion of the research work for this dissertation would be impossible without the

assistance of so many people. First and foremost I would like to thank my advisor, Dr. Eric D.

Wachsman, who guided me through my entire course of study and encouraged me to reach a

higher level of success. I am truly grateful for his generosity in sharing his knowledge and

insights with me.

I would like to thank Florida Institute for Sustainable Energy and the National Aeronautics

and Space Administration for funding NASA3-2930. I would also like to thank my committee

members Dr.Mark Orazem, Dr. David Norton, Dr. Scott Perry and Dr. Ying Shirley Meng

(replacement for Dr. Wolfgang Sigmund) for their valuable time and contributions to my

dissertation.

In addition, I would like to thank Dr. Heesung Yoon and Dr. Takkeun Oh for their valuable

discussion and assistance. Thanks to Dr. Sean Bishop and Dr. Martin Van Assche for their

valuable comments, editing and friendship. Furthermore, I want to thank the following people for

their expert advice, assistance and valuable friendship: Dr. Keith Duncan, Dr. Xin Guo, Dr.

Guojing Zhang, Dr. Yanli Wang, Dr. Aijie Chen, Dr. Cynthia Kan, Dr. Jeremiah Smith, Mr. Eric

Macam, Mr. Dohwon Jung, Mr. Bryan Blackburn, Mr. Danijel Gostovic, Mr. Nicholas Vito, Dr.

Matthew Camaratta, Dr. Shobit Omar, Dr. Jinsoo Ahn, Mr. Eric Armstrong, Mr. Dongjo Oh, Mr.

Byungwook Lee, Mr. Kangtaek Lee, Dr. Briggs White, and our secretary Mrs. Jennifer Tucker.

Most of all, many thanks to my parents and other family members for their unselfish

support and love throughout my life. Their trust in me has made me a strong and confident

person. I also would like to thank my friends Dr. Qi Wei, Mrs. Rongrong Liu, Dr. Xiaomin Lv,

Mr. Tianyuan Deng, Mr. Ting Zhu, Mr. Zhiliang Kong, Mrs. Xing Zhang, Mr. Hanneng Li, and

others which are far too many to mention for their priceless friendship.

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TABLE OF CONTENTS

page

ACKNOWLEDGMENTS.................................................................................................................... 4

LIST OF TABLES................................................................................................................................ 8

LIST OF FIGURES .............................................................................................................................. 9

ABSTRACT ........................................................................................................................................ 12

CHAPTER

1 INTRODUCTION....................................................................................................................... 14

2 LITERATURE REVIEW ........................................................................................................... 19

2.1 Hydrogen Production Technologies................................................................................. 19 2.1.1 Thermochemical Processes ................................................................................... 19 2.1.2 Electrolytic Processes ............................................................................................ 20 2.1.3 Photolytic Processes .............................................................................................. 20

2.2 Hydrogen Separation Membranes .................................................................................... 21 2.3 Proton Conducting Materials ............................................................................................ 22 2.4 Structure of SrCeO3........................................................................................................... 23 2.5 Proton Transport in SrCeO3 .............................................................................................. 24 2.6 Hydrogen Permeation ....................................................................................................... 26 2.7 Hydrogen Membrane System Design .............................................................................. 27

3 FABRICATION OF SUPPORTED TUBULAR SrCe 0.9Eu 0.1O3-δ and SrCe0.7Zr0.2Eu0.1O3-δ THIN FILM MEMBRANES ................................................................... 38

3.1 Introduction ....................................................................................................................... 38 3.2 Fabrication of Supported Thin Film Membranes ............................................................ 39

3.2.1 Materials Synthesis ................................................................................................ 39 3.2.2 NiO-SCZ82 Slurry for Support ............................................................................ 39 3.2.3 SrCe0.7Zr0.2Eu0.1O3-δ Thin Film Membranes on NiO-SCZ82 Support ............... 40

4 HIGH TEMPERATURE SrCe0.9Eu0.1O3 -δ PROTON CONDUCTING MEMBRANE REACTOR FOR H2 PRODUCTION USING THE WATER GAS SHIFT REACTION ..... 50

4.1 Introduction ....................................................................................................................... 50 4.2 Experimental...................................................................................................................... 52 4.3 Results and Discussion ..................................................................................................... 53

4.3.1 Thermodynamic Calculation ................................................................................. 53 4.3.2 Experimental Conversion ...................................................................................... 54 4.3.3 H2 Production ......................................................................................................... 57

4.4 Conclusions ....................................................................................................................... 58

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5 STABILITY OF SrCe1-xZrxO3-δ UNDER WATER GAS SHIFT REACTION CONDITIONS............................................................................................................................. 67


5.3.1 Stability under Wet CO Conditions ...................................................................... 69 5.3.2 Decomposition Mechanism................................................................................... 71 5.3.3 Hydrogen Permeability ......................................................................................... 72

5.4 Conclusions ....................................................................................................................... 73

6 HYDROGEN PERMEATION OF THIN SUPPORTED SrCe0.7Zr0.2Eu0.1O3-δ MEMBRANES UNDER DIFFERENT OXYGEN PARTIAL PRESSURE .......................... 82

6.1 Introduction ....................................................................................................................... 82 6.2 Experimental...................................................................................................................... 82

6.2.1 Membrane Fabrication .......................................................................................... 82 6.2.2 Membrane Morphology......................................................................................... 83 6.2.3 Membrane Permeation .......................................................................................... 83

6.3 Result and Discussion ....................................................................................................... 83 6.3.1 Heat Treatment ...................................................................................................... 83 6.3.2 Flow Rate Effect on H2 Permeation ..................................................................... 84 6.3.3 H2 Permeation as a Function of Thickness .......................................................... 84 6.3.4 Effect of Temperature, H2 and H2O Partial Pressure in the Feed Side on H2

Permeation............................................................................................................. 85 6.3.5 Activation Energy .................................................................................................. 86 6.3.6 Long Term Stability .............................................................................................. 87

6.4 Conclusions ....................................................................................................................... 87

7 SrCe0.7Zr0.2Eu0.1O3-δ-BASED HYDROGEN TRANSPORT WATER GAS SHIFT REACTOR................................................................................................................................... 95


7.3.1 Heat Treatment of the Membranes ....................................................................... 96 7.3.2 H2O/CO Effect on CO Conversion....................................................................... 96 7.3.3 H2O/CO Effect on H2 Production ......................................................................... 97 7.3.4 H2O/CO Effect on H2 Production and H2/CO ..................................................... 98 7.3.5 Flow Rate Effect on WGS Reaction..................................................................... 99 7.3.6 CO Concentration Effect on WGS Reaction ....................................................... 99 7.3.7 Long Term Stability ............................................................................................ 100

7.4 Conclusions ..................................................................................................................... 100

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8 HIGH TEMPERATURE SrCe0.7Zr0.2Eu0.1O3-δ MEMBRANE REACTOR FOR H2 PRODUCTION AND SEPARATION USING THE STEAM REFORMING OF METHANE ................................................................................................................................ 112

8.1 Introduction ..................................................................................................................... 112 8.2 Experimental.................................................................................................................... 113 8.3 Results and Discussion ................................................................................................... 114

8.3.1 Thermodynamic Calculation Results ................................................................. 114 8.3.2 Experimental Results ........................................................................................... 116

8.3.2.1 Influence of CH4/H2O on the SRM .................................................... 116 8.3.2.2 Influence of CH4 concentration on the SRM ..................................... 117 8.3.2.3 Influence of total flow rate on the SRM ............................................. 118 8.3.2.4 Influence of the H2 membrane reactor on the SRM .......................... 118 8.3.2.5 Long term stability ............................................................................... 119

8.4 Conclusions ..................................................................................................................... 119

9 HIGH TEMPERATURE SrCe0.7Zr0.2Eu0.1O3-δ PROTON CONDUCTING MEMBRANE REACTOR FOR CARBON DIOXIDE REFORMING OF METHANE .... 134

9.1 Introduction ..................................................................................................................... 134 9.1.1 Carbon Dioxide Reforming of Methane (CDRM) ............................................ 134 9.1.2 Membrane Reactors for the CDRM ................................................................... 135 9.1.3 Reaction Mechanism and Kinetics ..................................................................... 136

9.2 Experimental.................................................................................................................... 138 9.3 Results and Discussion ................................................................................................... 139

9.3.1 CH4/CO2 Effect on Conversion, H2/CO and H2 Production ............................. 139 9.3.2 Flow Rate Effect on Conversion, H2/CO and H2 Production ........................... 142 9.3.3 CH4/CO2/H2O Effect on XCH4, XCO2, H2/CO and H2 Production ..................... 143

9.4 Conclusions ..................................................................................................................... 144

10 CONCLUSIONS AND FUTURE WORKS ........................................................................... 155

10.1 Conclusions ..................................................................................................................... 155 10.2 Future Work ..................................................................................................................... 157

LIST OF REFERENCES ................................................................................................................. 160

BIOGRAPHICAL SKETCH ........................................................................................................... 169

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LIST OF TABLES

Table page 2-1 Properties of relevant hydrogen selective membranes [14]................................................. 30

2-2 Conductivities of potential proton conducting membranes [19]. ........................................ 31

2-3 Structural parameters of SrCeO3. .......................................................................................... 32

5-1 Intensity ratios between the strongest peaks of CeO2 and SCZ82. ..................................... 81

5-2 Tolerance factors of SrCe1-xZrxO3-δ. ..................................................................................... 81

6-1 Activation energy as a function of H2 partial pressure under dry H2 and H2/3% H2O conditions balanced by Ar. .................................................................................................... 94

6-2 Activation energy as a function of H2O partial pressure with a constant H2 flow rate of 20 cm3/min. ........................................................................................................................ 94

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LIST OF FIGURES

Figure page 2-1 Proton conductivities of various oxides [37]. ....................................................................... 33

2-2 Structure of SrCeO3 A) (001) projection and B) AO12 (blue) and BO6 (green) ................ 34

2-3 XRD pattern of SrCeO3 [46] ................................................................................................. 35

2-4 Predominant proton transfer between oxygen sites (shown by arrows) in the CeO6 octahedra of orthorhombically distorted BaCeO3 and SrCeO3 ........................................... 36

2-5 Comparison four categories setups of cross-flow operation [14]. ...................................... 37

3-1 XRD patterns of as-calcined SCZ82 and SCZE721 samples at 1300 oC ........................... 42

3-2 DV-E Viscometer ................................................................................................................... 43

3-3 Viscosity of NiO-SCZ82 slurry as a function of shear rate. ................................................ 44

3-4 Schematic process flow chart for fabrication of SCZ721 thin film membranes on NiO-SCZ82 supports.............................................................................................................. 45

3-5 Tape caster for making ceramic green tapes ........................................................................ 46

3-6 Process sequence for fabricating one end closed green body supports .............................. 47

3-7 Pictures of tubular SCZE721 thin film membrane coated on the inner side of NiO (or Ni)-SCZ82 support at each processing step ......................................................................... 48

3-8 SEM images of the NiO-SCZ82 and SCZE72 ..................................................................... 49

4-1 Morphology of thin film membranes and experimental setup ............................................ 59

4-2 Thermodynamic equilibrium of WGS under A) H2O/CO =1/1 and B) H2O/CO =2/1 ...... 60

4-3 Blank reference effluent gas composition as a function of temperature under .................. 61

4-4 Catalytic effluent gas composition as a function of temperature under ............................. 62

4-5 Catalytic effluent gas composition with in situ H2 removal as a function of temperature for H2O/CO =2/1 feed gas ................................................................................ 63

4-6 Temperature dependence of XCO under 3% CO + 3 % H2O and 3% CO + 6% H2O ........ 64

4-7 H2 production under 3% CO + 6% H2O as a function of temperature for three reactor configurations. ........................................................................................................................ 65

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4-8 H2 yield and syngas H2/CO ratio as a function of temperature under 3% CO + 6% H2O and with in situ H2 removal ........................................................................................... 66

5-1 Temperature profile and photograph of the membranes ...................................................... 74

5-2 XRD pattern of SrCeO3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h .................. 75

5-3 XRD Pattern and their lattice parameters of SrCe1-xZrxO3-δ as-calcined at 1300 oC ......... 76

5-4 XRD Pattern of SrCe1-xZrxO3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h at 800 oC...................................................................................................................................... 77

5-5 XRD Pattern of SrCe0.8Zr0.2O3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h ....... 78

5-6 XRD Pattern of SrCe0.8Zr0.2O3-δ after stability experiment at different atmospheres at 800 oC...................................................................................................................................... 79

5-7 H2 permeation as a function of time under 5% CO and 3% H2O at 900 oC ....................... 80

6-1 SrCe0.7Zr0.2Eu0.1O3-δ membrane and experimental setup ..................................................... 88

6-2 H2 permeation and H2 recovery as a function of feed flow rates ........................................ 89

6-3 H2 permeation vs thickness at 900oC .................................................................................... 90

6-4 H2 permeation as a function of H2 partial pressure and temperature .................................. 91

6-5 H2 permeation as a function of feed steam concentration and temperature. ...................... 92

6-6 H2 permeation as a function of time...................................................................................... 93

7-1 Membrane morphology and experiment setup ................................................................... 101

7-2 Gas compositions of the reactor side effluent as a function of temperature..................... 102

7-3 XCO as a function of temperature ........................................................................................ 104

7-4 H2 production as a function of temperature ........................................................................ 105

7-5 H2 yield and H2/CO in the reactor side effluent as a function of temperature ................. 107

7-6 The XCO, H2 production and H2/CO in the reactor side effluent as a function of flow rates under 900 oC ................................................................................................................ 109

7-7 The XCO, H2 production and H2/CO in the reactor side effluent as a function of CO concentrations with H2O/CO=2/1 ....................................................................................... 110

7-8 The performance of the membrane reactor as a function of time under 900 oC .............. 111

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8-1 Membrane morphology and experimental setup ................................................................ 121

8-2 Influence of CH4/H2O on XCH4 under thermodynamic equilibrium ................................. 122

8-3 Influence of CH4/H2O concentrations on XCH4 under thermodynamic equilibrium with CH4/H2O=1/2 and Ar as the diluent ........................................................................... 123

8-4 Thermodynamic calculation of carbon formation as a function of temperature and CH4/H2O [130] ..................................................................................................................... 124

8-5 Influence of CH4/H2O on XCH4............................................................................................ 125

8-6 Influence of CH4/H2O on SCO, SCO2 and H2/CO in reactor side effluent .......................... 126

8-7 Influence of CH4/H2O on H2 production ............................................................................ 127

8-8 Influence of CH4 concentration on SRM ............................................................................ 128

8-9 Influence of total flow rate on SRM ................................................................................... 129

8-10 Influence of reactor configurations on SRM. ..................................................................... 132

8-11 The performance of the membrane reactor as a function of time under 850 oC .............. 133

9-1 Membrane morphology and experimental setup ................................................................ 145

9-2 XCH4 and XCO2 as a function of temperature and CH4/CO2 ............................................... 146

9-3 SH2 and SCO as a function of temperature and CH4/CO2 .................................................... 147

9-4 H2 production as a function of temperature and CH4/CO2. ............................................... 148

9-5 H2/CO in the reactor side effluent as a function of temperature and CH4/CO2................ 149

9-6 XCH4, XCO2, SH2 and SCO as a function of total flow rate. .................................................. 150

9-7 H2/CO in reactor side effluent as a function of total flow rate. ......................................... 151

9-8 H2 production as a function of total flow rate. ................................................................... 152

9-9 XCH4 and XCO2 as a function of temperature. ...................................................................... 153

9-10 H2 production and H2/CO as a function of temperature. ................................................... 154

10-1 A SrCe0.7Zr0.2Eu0.1O3-δ thin film membrane coated on graphite-SrCe0.8Zr0.2O3-δ substrate. ............................................................................................................................... 159

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Abstract of Dissertation Presented to the Graduate School of the University of Florida in Partial Fulfillment of the Requirements for the Degree of Doctor of Philosophy

SrCeO3-BASED PROTONIC CONDUCTORS FOR HYDROGEN PRODUCTION AND

SEPARATION BY WATER GAS SHIFT, STEAM REFORMING, AND CARBON DIOXIDE REFORMING REACTIONS

By

Jianlin Li

August 2009 Chair: Eric Wachsman Major: Materials Science and Engineering

Hydrogen has been considered as an ideal energy carrier for a clean and sustainable energy

future. New ceramic membranes have potential to reduce the syngas (a mixture of hydrogen and

carbon monoxide) cost by 30-50% and incorporate hydrogen production and separation into one

unit. SrCe1-x-yZryEuxO3-δ has been investigated to maximize hydrogen production and enhance

stability. 10 at% europium was used to fabricate tubular micro-cracking free membranes. 20 at%

zirconium was used to enhance the stability of SrCe0.9Eu0.1O3-δ.

Supported SrCe0.7Zr0.2Eu0.1O3-δ thin film membranes on NiO-SrCe0.8Zr0.2O3-δ substrates

were developed. Hydrogen permeation flux through these membranes was proportional to the

transmembrane Hydrogen partial pressure gradient with a 1/4 dependence and controlled by bulk

diffusion. A maximum Hydrogen permeation of 0.23 and 0.21 cm3/cm2 min was obtained for the

33 μm thick SrCe0.7Zr0.2Eu0.1O3-δ membrane at 900 oC with 100% H2 and 97% H2/3% H2O as the

feed gases, respectively. Hydrogen permeation was stable under wet H2, and conditions of WGS

reaction, steam reforming of methane (SRM), and carbon dioxide reforming of methane

(CDRM).

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Thermodynamic equilibrium calculations were carried out for WGS reaction and SRM.

Hydrogen production and separation through WGS reaction, SRM and CDRM with

SrCe0.7Zr0.2Eu0.1O3-δ membranes were investigated. In situ removal of hydrogen through

hydrogen membranes moves the reaction toward the products side resulting in higher conversion

and hydrogen yield. 77% and 44% increase in the CO conversion for the WGS reaction was

achieved compared to the thermodynamic calculation data under 900 oC with H2O/CO = 1/1 and

2/1, respectively. 73% and 42% enhancement in the hydrogen production was achieved

simultaneously. For the SRM, the hydrogen membrane increased both the CH4 conversion and

total hydrogen production by 15% at 900 oC compared to the conventional reactor with only Ni

catalyst.

Whereas the H2/CO in the syngas product from the SRM is too high to produce liquid fuels

through the Fischer-Tropsch process, it is too low from the CDRM. However, an appropriate

value can be obtained by combining the SRM and CDRM. The H2/CO between 700 oC to 900

oC, for instance, is between 1.9-1.7 and 2.5-2.0 for CH4/CO2/H2O = 2/1/1 and 2/1/1.5,

respectively.

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CHAPTER 1 INTRODUCTION

H2 is perceived as an ideal energy carrier for a clean and sustainable energy future. It is the

simplest common element consisting of only one proton and one electron. It is the most abundant

chemical-energy resource in the world. However, it is not a primary source of energy as it occurs

only in nature in combination with other elements, primarily with oxygen in water and with

carbon, nitrogen and oxygen in living materials and fossil fuels.

Although generation of H2 from renewable energy sources has the potential to provide a

sustainable energy cycle, fossil fuels would provide a short- to medium-term solution to generate

H2 without additional adverse environmental impacts [1, 2]. The major source of H2 is steam

reformation of natural gas. Therefore, improvements in the efficiency and cost of H2 production

from natural gas are necessary in the near term. Gas separation membranes and membrane

reactors based on ion conducting ceramics may provide the technological advance necessary to

increase the efficiency and reduce the cost of H2 production from natural gas. However, other

sources of H2 must be developed for the envisioned H2 economy, and coal provides the greatest

U.S. domestic resource-based option. The U.S. DOE is developing a FutureGen plant based on

coal gasification, solid oxide fuel cells (SOFCs), and ion conducting membranes that will

produce H2 and electricity with zero emissions and carbon sequestration; thereby, not

contributing to global warming. The use of coal will help ensure America's energy security by

developing technologies that utilize a plentiful domestic resource.

Membrane reactor technology holds the promise to circumvent thermodynamic

equilibrium limitations by in situ removal of product species, resulting in improved chemical

yields. Mixed-conducting oxide-membrane technology presents the possibility for a dramatic

reduction in the cost of converting petroleum and coal derived feed stocks to H2 and other value-

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added hydrocarbons. Some perovskite oxides such as SrCe1-xEuxO3-δ exhibit both ionic and

electronic (mixed) conductivity. Because of their significant electronic conductivity, these mixed

ionic-electronic conductors (MIECs) have an internal electrical short and the ionic species

selectively permeates through a dense film of the material under a differential partial pressure,

such as H2 permeation. The potential permeation rates of these materials are extremely high.

SrCeO3-δ is a protonic conductor with high protonic conductivity and relatively low

electronic conductivity. To be suitable for H2 separation, a membrane material must have

comparable protonic and electronic transference numbers. In addition, the proton transference

number must be much higher than the transference number for oxygen ion. To that end, my

colleague, Dr. Takkeun Oh, has investigated the effect of dopant concentration in SrCe1-xEuxO3-δ

(0.05≤x≤0.2) on ambipolar conductivity [3] and demonstrated that the maximum ambipolar

conductivity increases with temperature and Eu dopant concentration. However, it is difficult to

fabricate a tubular thin film membrane with En dopant concentration higher than 10 at% without

micro-cracking [4]. Therefore, 10 at% Eu dopant concentration was used in my work to maintain

mechanical stability.

My overall goal is to demonstrate the feasibility of producing H2 from hydrocarbon based

fuels using advanced proton conducting membranes. The objective of my research is to improve

the stability of SrCeO3-δ using Zr as a dopant; to fabricate tubular supported SrCe0.7Zr0.2Eu0.1O3-δ

thin film membranes; to measure the H2 permeation of the SrCe0.7Zr0.2Eu0.1O3-δ membrane; and

to incorporate H2 permeation and total production using this membrane through water gas shift

(WGS) reaction, steam reforming of methane (SRM) and carbon dioxide reforming of methane

(CDRM).

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Fabrication of supported tubular thin film membranes: According to the Wagner

equation, when transport is bulk diffusion limited permeation through a MIEC membrane is

inversely proportional to thickness [5]. Therefore, our research has focused on the development

of thin film mixed protonic-electronic conducting membranes using porous tubular supports for

increased hydrogen production [6, 7]. SrCe0.9Eu0.1O3-δ and SrCe0.7Zr0.2Eu0.1O3-δ thin film

membranes were fabricated by tape casting followed by a rolling process (chapter 3).

SrCe0.9Eu0.1O3-δ membrane reactors for H2 production through WGS reaction: WGS

reaction converts CO and H2O to CO2 and H2. It is used to shift the CO/H2 ratio in the syngas

prior to Fischer-Tropsch synthesis and/or increase H2 yield. The WGS reaction is exothermic and

limited by thermodynamic equilibrium. When a H2 membrane reactor couples the H2 production

and separation together, continuous removal of H2 decreases the H2 concentration in the reaction

system and moves the reaction forward. As a result, CO conversion and H2 yield can be

increased. The thermodynamic equilibrium of the WGS reaction was calculated. The WGS

reaction was investigated under three reactor configurations and as a function of temperature and

H2O/CO (chapter 4).

Stability improvement of SrCe0.9Eu0.1O3-δ: The H2 permeation of SrCe0.9Eu0.1O3-δ is

stable under wet H2 atmospheres but degrades under dry H2 conditions [7]. It is unstable under

the WGS reaction conditions as well. Therefore, the stability of this material needs to be

improved. Zr has been used to improve the chemical stability of BaCeO3-δ system [8-12] and

SrCe0.95Yb0.05O3-δ [13]. The stability of SrCe0.9Eu0.1O3-δ was improved with zirconium dopant.

The stability of SrCe0.8Zr0.2O3-δ was investigated under different atmospheres. CO2 was found to

cause the decomposition of SrCe0.8Zr0.2O3-δ (chapter 5).

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H2 permeation properties of the SrCe0.7Zr0.2Eu0.1O3-δ thin film membranes: The H2

permeation properties of the SrCe0.7Zr0.2Eu0.1O3-δ membranes are not known yet. It was

investigated as a function of H2 partial pressures, feed flow rates, steam partial pressures and

temperature. The activation energy of the permeation process was discussed. The long term

stability of H2 permeation under wet H2 and the conditions of WGS reaction and SRM was

investigated as well (chapter 6).

SrCe0.7Zr0.2Eu0.1O3-δ based H2 transport WGS reactor: Chapter 4 compares the CO

conversion and H2 yield under different reactor configurations and they are significantly

improved with the H2 membrane reactor. In chapter 7, the SrCe0.7Zr0.2Eu0.1O3-δ effect on the

WGS reaction was investigated in details as a function of temperature, H2O/CO, CO

concentration, and CO feed flow rates. A long term stability experiment was carried out as well.

SrCe0.7Zr0.2Eu0.1O3-δ membrane reactors for H2 production through SRM: Currently,

the major H2 is produced from SRM. The SRM reaction is limited by thermodynamic

equilibrium. It needs to be carried out at high temperature to achieve high CH4 conversion.

Therefore, the SRM reaction is highly capital intensive. Catalytic ceramic membranes supply an

option to incorporate H2 separation and SRM into one unit which can increase the CH4

conversion or decrease the operating temperature. In chapter 8, the thermodynamic equilibrium

of the SRM was calculated. The SrCe0.7Zr0.2Eu0.1O3-δ membrane effect on the SRM was

investigated by comparing the performance under three different reactor configurations. The

performance of the SRM with the SrCe0.7Zr0.2Eu0.1O3-δ membrane was investigated as a function

of temperature, CH4/H2O, CH4 concentration, and CH4 feed flow rates. The membrane stability

under the SRM conditions was studied as well.

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SrCe0.7Zr0.2Eu0.1O3-δ membrane reactor for H2 production through CDRM: While

SRM is the major process for H2 production, it produces a large amount of CO2 simultaneously.

0.3-0.4 million cubic meters of CO2 will be produced when one million cubic meters of H2 is

produced through a typical SRM H2 plant. Therefore, CO2 sequestration has drawn lots of

interest. The capture and disposal of CO2 costs a significant portion of the total cost of H2

production by the SRM process. The net cost of CO2 disposal, however, could be significantly

reduced if CO2 sequestration is accompanied by an enhanced product. CDRM provides one

solution to sequester CO2 and produce syngas simultaneously. In chapter 9, the performance of

the CDRM with the SrCe0.7Zr0.2Eu0.1O3-δ membrane was investigated as a function of

temperature, CH4/CO2, CH4 concentration, and CH4 feed flow rates. In addition, whereas the

H2/CO in the syngas from the SRM is too high to produce liquid fuels through the Fischer-

Tropsch process, it is too low in the syngas from the CDRM. Therefore, the SRM and CDRM

was combined to obtain appropriate H2/CO values.

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CHAPTER 2 LITERATURE REVIEW

2.1 Hydrogen Production Technologies

H2 production processes are based on the separation of H2 from H2-containing feedstocks.

It can be produced using a variety of domestic energy resources - fossil fuels, such as coal and

natural gas; renewables, such as biomass, and renewable energy technologies, including solar,

wind, geothermal, and hydropower; nuclear power. H2 production technologies fall into three

general categories: thermochemical processes, electrolytic processes and photolytic processes.

2.1.1 Thermochemical Processes

Steam methane reforming (SMR): SMR is the most efficient and widely used process for

the production of H2. About 95% of the H2 in the United States is produced using this process (3-

25 bar, 700 oC-1000 oC).

In this process, high-temperature steam is used to extract H2 from a methane source such as

natural gas. This process consists of three steps: 1) reformation of the methane with high

temperature steam to obtain a syngas; 2) using a WGS reaction to form H2 and CO2, and 3)

purification. The reactions are listed below:

Step 1: CH4 + H2OCO + 3H2 1 6.205 −+=∆ kJmolH o (2-1)

Step 2: CO + H2OCO2 + H2 1 6.40 −−=∆ kJmolH o (2-2)

After the first two steps, a membrane is required to extract high-purity H2 from the H2 and

CO2 stream.

Partial oxidation: In this process, a fuel and oxygen are combined in proportions such that

the fuel is converted into a mixture of H2 and CO. There are several modifications of this

process, depending on the composition of the process feed and type of the fossil fuel used. Partial

oxidation of methane can be described by the following equation:

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CH4+1/2O2CO+2H2 1 36 −−=∆ kJmolH o (2-3)

Coal gasification: Coal is converted into syngas by reacting coal with oxygen and steam

under high pressures and temperatures. Its gasification reaction may be represented by the

(unbalanced) reaction equation:

CH0.8+O2+H2OCO+CO2+H2+other species (2-4)

An advantage of this technology is that CO2 can be separated and captured more easily

from the syngas instead of being released into the atmosphere. If CO2 can be successfully

sequestered, H2 can be produced from coal gasification with near-zero greenhouse gas emission.

This technology is most appropriate for large-scale, centralized H2 production.

Other thermal processes: Other processes include (1) splitting water using heat from a

solar concentrator, and (2) gasifying or burning biomass (i.e., biological material, such as plants

or agricultural waste) to generate a bio-oil or gas, which is then reformed to produce H2.

2.1.2 Electrolytic Processes

Electrolysis: In electrolysis, electricity is used to split water (H2O) into H2 and oxygen.

The reaction takes place in a unit called an electrolyzer. There are three major electrolyzers:

polymer electrolyte membrane electrolyzer, alkaline electrolyzers, and solid oxide electrolyzers.

Current electrolysis systems are very energy intensive. The challenge is to develop low cost and

more energy efficient electrolysis technologies.

2.1.3 Photolytic Processes

Photolytic methods: In photolysis, sunlight is used to split water. Two photolytic

processes are being explored: (1) photobiological methods, in which microbes, when exposed to

sunlight, split water to produce H2, and (2) photoelectrolysis, in which semi-conductors, when

exposed to sunlight and submersed in water, generate enough electricity to produce H2 by

21

splitting the water. These processes offer long-term potential for sustainable H2 production with

low environmental impact.

2.2 Hydrogen Separation Membranes

H2 selective membranes can be classified into four categories: polymeric, metallic, carbon

and ceramic. Table 2-1 summarizes their properties [14].

Polymeric membranes are dense membranes, transporting species through the bulk of the

material. They cope well with high pressure-drops and are low cost. However, their operating

temperatures are limited to 90-100 oC. They are sensitive to certain chemicals and have weak

mechanical strength.

Palladium and palladium alloy membranes are the typical metallic membranes. They have

been studied and used as membrane reactors [15-17]. They have high H2 selectivity. However,

palladium-based membrane reactors have been operated at low temperatures of 300-500 oC [18].

They are highly sensitive to chemicals such as sulphur, chlorine and even CO. In addition,

palladium based membranes are expensive since palladium is a precious metal.

Carbon membranes separate H2 from other gases using small pores which only H2 can pass

through. They are usually used in non-oxidizing environments from 500 to 900℃. However,

they are difficult to fabricate and very brittle. Their selectivity is low, in the range of 4-20.

Ceramic membranes are a combination of a metal with a non-metal. They can be porous or

dense. Porous ceramic membranes generally are separation membranes on more porous ceramic

substrates. Their operating temperature is between 200 and 600 oC. One drawback for these

membranes is their poor hydrothermal stability. Dense ceramic membranes are also called proton

conducting membranes. The selectivity is very high since only H2 ions can transport through the

membranes. SrCeO3-δ and BaCeO3-δ are typical materials with an operating temperature from

22

600 to 900℃. Their chemical stability in the presence of certain species (e.g., CO2 and H2S) is a

major concern.

2.3 Proton Conducting Materials

There are numerous kinds of proton conducting membranes. Table 2-2 summarizes their

conductivities and operating temperatures [19]. Among those proton conducting materials,

perovskite (in form of ABO3, e.g., SrCeO3, BaCeO3) and related structures are of interest as

proton selective membrane materials [20-30]. In order to enhance the ionization of H2, cermet

membranes with a continuous metallic phase were also studied [31, 32]. In addition, complex

perovskites in the form of A2B1+x′B1-x″MxO6-δ (where A=Ba or Sr, B’=trivalent ion and

B”=pentavalent ion) or A3B1+x′B2-x″Mx2O9-δ (where A=Ba or Sr, B’=trivalent ion and

B”=pentavalent ion) [5, 33-36] have been developed to increase stability of the perovskite

oxides.

To be commercially useful, H2 separation membranes, the perovskite oxides should have

both high electronic and protonic conductivity, higher than 0.1 S cm-1 [19], and be stable in

operating conditions. While the electronic conductivities of SrCeO3 and BaCeO3 are relatively

low, they can be increased significantly by substituting Ce4+ with aliovalent ions (Y, Yb, Gd and

Eu) [3, 20]. The oxygen vacancies, typically created by acceptor doping to maintain

electroneutrality, play an important role for proton conduction.

Figure 2-1 shows the proton conductivities of several electronic-protonic conducting

ceramics [37]. Among those ceramics, BaZrO3-δ exhibits the highest proton conductivity.

However, its electronic conductivity is very low. To achieve high H2 permeability, the electronic

and protonic transference numbers should be comparable [27, 28]. Therefore, BaZrO3-δ is not a

good candidate for H2 separation membrane. BaCeO3-δ based oxides exhibit oxygen ion

23

conductivity comparable to their proton conductivity and thus are not proper for applications

where oxygen is present [38, 39]. Furthermore, BaCeO3 undergoes a complex sequence of phase

transitions [40], including a first order phase transition from orthorhombic to rhombohedral at

663-673 K where the two distinct oxygen sites become crystallographically equivalent. In

contrast, SrCeO3 undergoes no high temperature structural phase transitions up to 1273 K [41]. It

has high total conductivity and highest proton transference number which is due to a distorted

orthorhombic structure of SrCeO3-δ inhibiting oxygen ion conduction [42, 43]. Therefore,

SrCeO3-δ based oxides can be promising for selective H2 separation if their electronic

conductivity can be improved by a proper doping.

2.4 Structure of SrCeO3

SrCeO3 has an orthorhombic structure at room temperature. The lattice parameters of

SrCeO3 are a=6.126 Å, b=8.574 Å, c=6.000 Å. The theoretical density is 5.81 g/cm3 [44]. The

space group is Pnma (no. 62) [45], Z=4, with the A cation, Sr, situated on mirror planes (4c); the

B-site cation, Ce, situated on centers of inversion (4b); and two oxygen positions: O (1) on

mirror planes (4c), and O (2) in general positions (8d). The coordination numbers of Sr, Ce and

O are 12, 6 and 6, respectively. Table 2-3 shows the parameters of SrCeO3 [42] . The crystal

structure of SrCeO3 is shown in Figure 2-2. Its XRD pattern is shown in Figure 2-3 [46].

The conductivity of SrCe1-xAxO3 (A= aliovalent ions) has drawn great interest from

researchers. The typical conductivity is found to be between 10-2 to 10-3 S cm-1[23, 27, 28, 47-

50]. The thermal expansion coefficient and thermal conductivity of SrCeO3 are 1.11 × 10-5 K-1

from room temperature to 1273 K and 2.95 W m-1 K-1 at room temperature, respectively [51].

The heat capacity was determined in the temperature range of 373-1400 K as follows [44]:

CP (J mol-1 K-1)= 120.1+5.45× 10-3 T-1.26× 106 T-2 (2-5)

24

2.5 Proton Transport in SrCeO3

Vehicle mechanism and Grotthuss mechanism are the two mechanisms generally accepted

for proton diffusion in perovskite systems. Vehicle mechanism was first brought forward by

K.D.Kreuer et al. in 1982 [52]. According to this model, a proton does not migrate as H+ but as

+3OH , +

4NH , etc., bonded to a “Vehicle ” such as H2O, NH3 etc. The “unloaded” vehicles move

in the opposite direction. As for SrCeO3 based materials under wet H2 (H2 & H2O) atmosphere,

the vehicle is •OOH . The major proton incorporation reaction is shown below using Kröger-Vink

notation:

•×•• ↔++ OOO OHOVOH 22 (2-6)

With Kröger-Vink notation, ••OV represents oxygen vacancies, X

OO represents oxide ions

on an oxygen lattice site, and •OOH represents protons associated with oxide ions on an oxygen

lattice site. Then, •OOH is driven by H2 partial pressure gradient and transfers. The radii of O2-

and OH- are 1.32Å and 1.35Å, and they have to go over a saddle point in the diffusion process,

i.e., they need to overcome high energetic barriers. According to K.D. Kreuer et al., the

activation enthalpy for site exchange of an oxygen and an oxygen-ion vacancy is 0.8eV

(77.2kJ/mol) [53]. Hence, their diffusion coefficients are relatively slow and vehicle mechanism

usually takes place at high temperature.

Over the last decade, general agreement has been formed that protons transfer between

fixed oxygen sites via the Grötthuss mechanism at intermediate temperatures in ABO3

perovskites. Isotope effect (H+/D+) measurements of perovskite oxides have suggested that the

conduction mechanism is due to proton hopping between adjacent oxygen ions (Grötthuss

mechanism) rather than by hydroxyl ion migration (vehicle mechanism) [54-57].

25

The process consists of two steps, the proton hopping translation between O-O bonds and

reorientation of the hydroxide ion on the oxygen site. Outside the high H2 partial pressure

surface, a 1/2H2 would decompose into a proton, H+, and an electron, e-. In the material, H+ hops

from O2- to O2-and e- hops between cerium ions.

+−+ ↔+ 34 CeeCe (2-7)

Inside the low H2 partial pressure surface, the H+ combines with e- to form 1/2H2,

terminating the net H2 diffusion. The proton hopping translation from one oxygen site to another

oxygen site depends on the O-O bond length. A longer O-O bond favors formation of O-H bond,

but impedes the proton hopping; conversely, a shorter O-O bond favors the proton hopping, but

it is not favorable for hydroxide ion reorientation.

Molecular dynamic studies in both cubic and orthorhombic perovskites proton conducting

oxides support for the Grötthuss mechanism. Munch et al determined that the reorientation

process occurs relatively fast (10-12s) compared to the proton transfer process (10-9s) indicating

the proton transfer process is the rate limiting step [58, 59] . The activation energy for rotational

diffusion of protonic defects is small, 65 meV (0.63kJ/mol) and the activation enthalpy for

proton diffusion was 0.41 eV (39.6kJ/mol) [60].

In cubic perovskites, neighboring oxygen ions are treated as equivalent sites. In contrast,

oxygen ions must be treated differently in low symmetry orthorhombic perovskites (such as

SrCeO3). The most basic oxygen sites are O1 and O2 in SrCeO3 and BaCeO3, respectively [60].

Proton transfer between oxygen sites in BaCeO3 and SrCeO3 is shown in Figure 2-4 [61]. The

long-range transport between O2 sites in BaCeO3 should be easier than transport between O1 and

O2 sites in SrCeO3 because the O1 and O2 are chemically different in SrCeO3. This difference in

proton transport is a possible reason for lower conductivity in SrCeO3 than BaCeO3 [62].

26

2.6 Hydrogen Permeation

H2 permeates through H2 permeable membranes via ambipolar diffusion of protons and

electrons under a H2 chemical potential gradient [63]. The motion of electrons, the minority

carrier, gives rises to the H2 permeation by charge compensated transport of protons in the same

direction.

A few theoretical works on H2 permeation in a MIEC have been published [64-66]. In

MIECs, the flux of each charge carrier species, k, is driven by chemical and electrical forces.

Along the dimension of net transport (x= membrane thickness), assuming the bulk diffusion is

the rate limiting step, mass transfer rates per cross-sectional area for component k are give by

[67]:

)()()( 22 dx

dFzdx

dFzFzRT

CDJ kk

k

kk

k

kk

Kkk

φµσηση −−=∇−=∇−= (2-8)

The terms in the parenthesis represent the chemical and electrical potential gradient, and zke is

the charge of the species k. In open circuit conditions, the net current resulting from all fluxes is

zero:

∑ ∑∑ −−

=====

)(03

1 dxdez

dxd

ezeJzIi k

k

k

kkk

k

φµσ (2-9)

Equation (2-9) can be rearranged:

∑−=

ii

i

i

dxd

ezt

dxd µφ (2-10)

12 =++ −+ OHe ttt (2-11)

where ti is the transference number of species i.

We insert equation (2-10) into equation (2-8) and obtain:

27

)(22 ∑+−

=i

i

i

ik

k

k

kk dx

dztz

dxd

ezj µµσ (2-12)

Now, equilibria between neutral and charged species are introduced. For MIECs, it is natural to

consider transport by oxygen ions, protons and electrons.

−=+ 2'2 OeO eOO ddd µµµ 22 +=− (2-13)

'eHH += + eHH ddd µµµ −=+ (2-14)

We insert equations (2-13) and (2-14) into equation (2-12) for the flux of oxygen ions:

])(2)[(4 2

2

2 dxdt

dxdtt

dt

j HH

OHe

OtotO

µµσ++−

−

− ++−

= (2-15)

where σtot is the total conductivity. If local thermodynamic equilibrium is achieved,

)ln(21

21

222

0OOOO PRT+== µµµ

2

ln2 OO PdRTd =µ (2-16)

We insert equation (2-16) into equation (2-15) and obtain:

]ln4

ln)(8

[1J''2

'2

''2

'2

2222-2 22O ∫ ∫ −++− −+−= O

O

H

H

P

P

P

P HOHtotOeHOtot PdttF

RTPdtttF

RTL

σσ (2-17)

where L is sample thickness.

Following the same manner, we obtain the proton flux:

∫ ∫ ++−= −+−++

''2

'2

''2

'2

2'222 ln)(2

ln4

1J 22H

O

O

H

H

P

P

P

P HeOHtotOOHtot PdtttF

RTPdttF

RTL

σσ (2-18)

2.7 Hydrogen Membrane System Design

Currently there are two types of membrane configurations: flat and tubular [14]. The

building block of a membrane system is called the module. Module types based on flat

membranes are the plate-and-frame and spiral-wound modules. Tubular type membrane modules

28

are subdivided into tubular (diameters > 10 mm), capillary (0.5 mm < diameters < 10 mm) and

hollow fiber modules (diameters < 0.5 mm) [68]. Table 2-4 lists the packing densities of these

modules.

There are single stage and multistage membrane processes. A stage is formed by one or

more membrane modules assembled into an operating unit that provides a specific function

different from any other stages that may be utilized in the same process. Multistage membrane

systems are built to improve membrane system performance.

Single stage membrane process is the simplest membrane process. Membrane operations

can be subdivided into dead-end and cross-flow operations [14]. In dead-end operation, there is

no retentate stream. It is not preferred since non-permeating species in time become more

abundant on the feed side, resulting in concentration polarization and driving force. As a result,

the transport through the membrane decreases. Instead, flows run alongside the membrane in the

cross-flow operations. Deterioration of membrane flux in time is limited in this configuration.

The cross-flow operation can be distinguished into four categories: co-current, counter-current,

cross-flow with perfect permeate mixing, and perfect mixing [14].

In the co-current operation, feed and permeate flows run in the same direction, whereas,

they run in opposite direction in the counter-current operation. In the perfect permeate mixing,

the permeate is mixed to form one homogeneous permeate composition along the membrane

length coordinate. The perfect mixing setup results in homogeneous compositions in both the

feed side and permeate side. Figure 2-5 compares the configurations and driving forces of these

four categories of cross-flow operations. Generally, the membrane results are obtained in this

sequence due to the driving forces: counter-current > cross-flow with permeate mixing > co-

current > perfect mixing [69].

29

Membranes can selectively take away reaction products, thereby shifting the equilibrium to

the product side. If chemical reactions are carried out in a membrane module, the system is

called a membrane reactor [14]. Catalysts are usually necessary to carry out reactions and

accommodated inside the membrane reactors. Three types of arrangements are found to

accomplish this [69]: catalyst placed inside the feed stream, catalyst placed in a membrane top

layer, and catalyst placed inside the membrane itself. To place catalyst inside the feed stream is

easy to prepare and operate. In contrast, it is difficult to replace the catalyst placed in a

membrane top layer or inside the membrane itself since replacing the catalyst usually means

replacing the whole membrane.

30

Table 2-1. Properties of relevant hydrogen selective membranes [14].

Dense Polymer Micro porous

ceramic Dense metallic Porous carbon Dense ceramic

Temperature range

<100 oC 200-600 oC 300-600 oC 500-900 oC 600-900oC

H2 selectivity low 5-139 >1000 4-20 >1000 H2 flux ( 10-3 mol/m2s) at dP=1 bar

low 60-300 60-300 10-200 6-80

Stability issues Swelling, mechanical strength

Stability in H2O

Phase transition

Brittle, oxidizing

Stability in CO2

Poisoning issues

HCl, SOx, (CO2)

H2S, HCl, CO Strong adsorbing vapors, organics

H2S

Materials Polymers Silica, alumina, zirconia, titania, zeolites

Palladium alloy

Carbon Proton conducting ceramics

Transport Mechanism

Solution/ diffusion

Molecular sieving

Solution/ diffusion

Surface diffusion, molecular sieving

Solution/ diffusion (proton conduction)

Development status

Commercial by Air Products, Linde, BOC, Air Liquide

Prototype tubular silica membranes available up to 90 cm. Other materials only small samples

Commercial by Johnson Matthey; prototype membrane tubes available up to 60 cm

Small membrane modules commercial, mostly small samples (cm2) available for testing

Small samples available for testing

31

Table 2-2. Conductivities of potential proton conducting membranes [19].

32

Table 2-3. Structural parameters of SrCeO3. Atom X Y z Sr 0.2500 0.0116 -0.0447 Ce 0.0000 0.5000 0.0000 O1 0.2500 0.6059 0.0432 O2 -0.0558 0.7006 0.2988

Table 2-4. Packing densities of different hydrogen membrane modules [68]. Module Plate-and-frame Spiral-wound Tubular Capillary Hollow fiber

Packing Density (m2/m3)

100-400 300-1000 300 600-1200 30000

33

-6

-5

-4

-3

-2

-1

0 0.5 1 1.5 2 2.5

Logσ

(S/c

m)

1000/T (K-1)

BaZrO3

BaCeO3

SrTiO3

SrCeO3

CaZrO3

SrZrO3

PH2O

=30 hPa

Figure 2-1. Proton conductivities of various oxides [37].

34

A

B

Figure 2-2. Structure of SrCeO3 A) (001) projection and B) AO12 (blue) and BO6 (green).

35

20 30 40 50 60 70 80

Inte

nsity

(a.u

.)

(0 1

1)

(2 1

1)

(1 2

1)

(3 1

1)

(0 2

2)

(1 2

2)

(0 3

7)

(4 0

2)

(2 3

1)

(3 1

3)

(4 2

2)

(1 1

6)

(0 4

4)

(4 0

4)

2θ (o)

Figure 2-3. XRD pattern of SrCeO3 [46].

36

Figure 2-4. Predominant proton transfer between oxygen sites (shown by arrows) in the CeO6 octahedra of orthorhombically distorted BaCeO3 and SrCeO3. The degree of basicity is indicated by the color of the oxygen sites (purple = more basic) [61].

37

Figure 2-5. Comparison four categories setups of cross-flow operation [14].

38

CHAPTER 3 FABRICATION OF SUPPORTED TUBULAR SrCe 0.9Eu 0.1O3-δ AND SrCe0.7Zr0.2Eu0.1O3-δ

THIN FILM MEMBRANES

3.1 Introduction

Perovskite-type oxides such as SrCeO3-δ exhibit significant proton conductivities in H2-

containing atmospheres when oxygen vacancies and other charged defects are introduced by the

partial substitution of trivalent cations for Ce [70]. We previously investigated the effect of

dopant concentration in SrCe1-xEuxO3-δ (0.05≤x≤0.2) on ambipolar conductivity [3] and found

that the maximum ambipolar conductivity increases with temperature and Eu dopant

concentration. However, we also found that Eu dopant concentrations higher than 10 at% result

in mechanical instability. Therefore, 10 at% Eu dopant is used in my dissertation.

Compared to planar membranes, tubular membranes have much larger area and do not

require any high-temperature seals to isolate permeated gas from input gas. In addition, the

Wagner equation shows that when transport is bulk diffusion limited permeation through a

mixed ionic-electronic conducting (MIEC) membrane is inversely proportional to thickness [5].

Thus, many studies have focused on the fabrication of thin film membranes [5, 28, 64, 66, 71-

73]. Therefore, our research has focused on the development of thin film mixed protonic-

electronic conducting membranes using porous tubular supports for increased H2 production [6,

7]. SrCe0.9Eu0.1O3-δ and SrCe0.7Zr0.2Eu0.1O3-δ thin film membranes were investigated in my work.

The fabrication of SrCe0.9Eu0.1O3-δ membrane has been addressed in reference [74]. In this

chapter, I focus on the fabrication of SrCe0.7Zr0.2Eu0.1O3-δ membranes. Based on our previous

experience, NiO-SrCe0.8Zr0.2O3-δ was used to fabricate the support structure to maintain

mechanical integrity. Eu was eliminated from the support composition since electronic

39

conduction is not functionally necessary for the support. NiO was used to create porosity and to

serve as the catalyst, by reduction to Ni when the membrane was subsequently exposed to H2.

3.2 Fabrication of Supported Thin Film Membranes

3.2.1 Materials Synthesis

Polycrystalline SrCe0.8Zr0.2O3-δ and SrCe0.7Zr0.2Eu0.1O3-δ powders were prepared by

conventional solid-state reaction by mixing stoichiometric amounts of SrCO3 (99.9%, Alfa-

Aesar), CeO2 (99.9%, Alfa-Aesar) ZrO2 (99.9%, Alfa-Aesar) and Eu2O3 (99.9%, Alfa-Aesar)

powders, followed by ball milling and calcining at 1300 oC. Figure 3-1 shows their XRD

patterns. Both are orthorhombic structure.

3.2.2 NiO-SCZ82 Slurry for Support

Homogeneous and stable slurry is very necessary for tape casting. The NiO-SCZ82 slurry

was achieved by two stage ball milling. Firstly, 46.7 wt% NiO was mixed with SCZ82 powder

and dissolved in ethanol and toluene which served as solvents. To stabilize the slurry against

flocculation of the particles, a certain amount of solsperse (2400SC, Avecia) was used as a

dispersant and added into the solution. This solution was ball milled for 24hrs. Secondly, Binders

and plasticizers were added to the solution and the solution was ball milled for another 24 hrs.

PVB was chosen as the binder to provide plasticity of the solution. Plasticizer can soften the

binder and increase the flexibility of the green body. For tape casting process, the plasticizer

must be soluble in the same solution used to dissolve the binder. Specific combinations of binder

and plasticizer are used in tape casting process. Here, PEG and DBP were used as the

plasticizers.

Rheology plays an important part in the processing of ceramics from colloidal suspensions.

When the suspension is consolidated by casting methods, including slip and tape casting, the

suspensions are required to contain the highest possible fraction of particles to reduce the

40

shrinkage during drying the cast and to produce a consolidated powder form with high packing

density. In addition, the suspension should have a low enough viscosity to be poured. [75]

The rheological properties of the slurry can be characterized by viscosity, η, defined by

γτη /= (3-1)

where τ and γ are the shear stress and shear rate, respectively. There are a number of types of

rheological behavior of colloidal suspension:

(1) Newtonian: Viscosity is constant with change in shear rate. (2) Thixotropic: Viscosity decreases as shear rate increases and is also a time dependent. (3) Dilatant/shear thickening: Viscosity increases as shear rate increases. (4) Plastic: Viscosity decreases as shear rate increases after an initial threshold stress. (5) pseudo-plastic/shear thinning: Viscosity decreases as shear rate increases

The viscosity of the slurry is measured by DV-E Viscometer (Brookfield) (Figure 3-2).

The accuracy is guaranteed to be ± 1%. Figure 3-3 shows the viscosity as a function of shear rate

at 25 oC. LV3 spindle is used and % torque is between 10 and 100 for the whole measurement.

The viscosity of the NiO-SCZ82 slurry is plastic and decreases with increasing shear rate.

3.2.3 SrCe0.7Zr0.2Eu0.1O3-δ Thin Film Membranes on NiO-SCZ82 Support

Figure 3-4 shows the process flow design for the preparation of the SrCe0.7Zr0.2Eu0.1O3-δ

(SCZE721) thin film membranes on tubular NiO-SCZ82 supports. The tubular NiO-SCZ82

support was fabricated using tape-casting (Pro-Cast) followed by a rolling process. Figure 3-5

shows the tape caster. The process sequence for making a one end closed tubular-type green

body support is shown in Figure 3-6. After the green tubes were pre-sintered at 1100 oC,

SCZE721 was coated on the inner side of the supported by colloidal coating. Then the SCZE721

membranes were sintered at 1520 oC together with the NiO-SCZ82 supports to achieve dense

membranes. Figure 3-7 and Figure 3-8 show the photographs and morphology of the tubular

membranes at different processing steps. As shown in Figure 3-8, the pre-sintered structures (A

41

and B) are very porous. The sintered structure is pretty dense. There are still some pores on the

supported structure, but those pores are isolated and close pores. The thin film surface is crack-

free.

42

20 30 40 50 60 70 80

2θ (o)

Inte

nsity

(a.u

.)

(4 2

2)

(0 3

7)

(0 2

2)

(2 1

1)

(4 0

2)

(0 1

1)

(1 1

6)

(4 0

4)

(1 2

2)

(2 3

1)

(0 4

4)

SCZ82

SCZE721

(1 2

1)

(3 1

1)

(3 1

3)

Figure 3-1. XRD patterns of as-calcined SCZ82 and SCZE721 samples at 1300 oC.

43

Figure 3-2. DV-E Viscometer.

44

0

5 103

1 104

1.5 104

2 104

2.5 104

0 0.1 0.2 0.3 0.4 0.5 0.6

Visc

osity

(cP)

Shear Rate (s-1)

Figure 3-3. Viscosity of NiO-SCZ82 slurry as a function of shear rate.

45

Figure 3-4. Schematic process flow chart for fabrication of SCZ721 thin film membranes on NiO-SCZ82 supports.

Mix NiO, SrCe0.8Zr0.2O3-δ, and dispersant with ethanol

and toluene, ball mill

Add binder, plasticizers and ball mill

De-air

Tape casting

Rolling on rod

Pre-sinter

Coat SrCe0.7Zr0.2Eu0.1O3-δ on inner side of the pre-sintered

support

Final sinter

Mix SrCe0.7Zr0.2Eu0.1O3-δ, and dispersant in ethanol,

ball mill

46

Figure 3-5. Tape caster for making ceramic green tapes.

47

Figure 3-6. Process sequence for fabricating one end closed green body supports.

48

Figure 3-7. Pictures of tubular SCZE721 thin film membrane coated on the inner side of NiO (or Ni)-SCZ82 support at each processing step.

NiO-SCZ82 tubular green substrate

Pre-sintered NiO-SCZ82 substrate (SCZE721 was coated on the inner side of the substrate)

Sintered SCEZ721membrane on NiO-substrate

Reduced SCEZ721membrane on Ni-SCZ82 substrate

49

A B

C D

Figure 3-8. SEM images of the NiO-SCZ82 and SCZE721 A) surface of the pre-sintered NiO-SCZ82; B) surface of the as-coated SCZE721 thin film; C) cross section of the sintered membrane and D) thin film surface of the sintered membrane.

50

CHAPTER 4 HIGH TEMPERATURE SrCe0.9Eu0.1O3 -δ PROTON CONDUCTING MEMBRANE REACTOR

FOR H2 PRODUCTION USING THE WATER GAS SHIFT REACTION

4.1 Introduction

Syngas mixtures containing mostly H2 and CO are typically generated at elevated

temperatures via the conversion of natural gas, coal, biomass, petroleum and organic wastes [76].

The water-gas shift (WGS) reaction, equation (4-1), converts CO into CO2 and provides

additional H2.

H2O + CO H2 + CO2 6.40−=∆ oH kJmol-1 (4-1)

The WGS reaction is often used in conjunction with steam reforming of methane or other

hydrocarbons and is of central importance in the industrial production of H2, ammonia, and other

bulk chemical utilizing syngas [77]. It is an important method for further enhancing H2 yield

and/or to shift the H2/CO. This is especially important for synthesis gas derived from coal, which

tends to have a H2/CO of ~0.7 compared to the ideal of ~2 for the Fischer-Tropsch process.

The WGS reaction is an exothermic reaction. Thermodynamic equilibrium favors high

conversion of CO and steam to H2 and CO2 at low temperatures. Therefore, it is typically a two-

stage shift process, a high-temperature WGS and a low-temperature WGS, with each process

employing separate catalysts [78, 79]. In addition, a cooling step is necessary before the second

stage. U.S. Department of Energy for the production of H2 suggested an alternative concept by

carrying out the WGS reaction at high temperature in a H2-selective membrane reactor [80]. New

ceramic membranes have potential for cost reduction of syngas production by 30-50% [81] and

provide one solution to incorporate the WGS reaction and H2 separation into one unit.

Selectively continuous removal of H2 will drive the WGS reaction equilibrium forward. As a

result, the requirement to use a two-stage shift reaction and a cooling step can be eliminated and

51

the WGS reaction may be carried out at higher temperatures [19]. This would allow the WGS

reaction to be operated at low H2O/CO without the thermodynamic constraint [16, 82, 83]. In

addition, the reaction rate of the homogeneous WGS reaction at high temperature could be high

enough that permeation through the membrane could be the rate limiting step [84]. Therefore, the

need for the introduction of heterogeneous catalyst could be eliminated.

There is growing interest in the WGS reaction assisted by a catalytic membrane reactor.

The WGS reaction has been carried out under various operating conditions using porous Vycor

glass coated with ruthenium (III) chloride trihydrate. The highest CO conversion was 85% at 157

oC and at a permeate rate of 0.64 cm3/min. The CO conversion is lower than the equilibrium

valune (99.9%) at the same conditions [85]. Extensive research has been focused on Palladium

and palladium alloy membranes [15-17]. However, palladium-based membrane reactors have

been evaluated most extensively at low temperatures of 300-500 oC [18]. The Palladium

membranes are highly fragile due to thermal excursion in the presence of H2 which causes poor

durability. In addition, Palladium based membranes are usually expensive. An alternative for

membrane reactor is ceramic membranes. It has been pointed out that a minimum between the

efficiency penalty and system complexity in a conventional integrated gasification combined

cycle power plant is obtained when the H2 and CO2 is separated at high temperature using a

catalytic ceramic membrane reactor [85]. Most of the catalytic ceramic membrane reactors are

SrCeO3-δ and BaCeO3-δ based perovskite mixed protonic electronic conductors [26-28, 32].

SrCeO3-δ has high total conductivity and highest proton transference number compared to

BaCeO3-δ and SrZrO3-δ [49]. However, its electronic conductivity needs to be improved. We

previously successfully improved the electronic conductivity of SrCeO3-δ using Eu dopant

(0.05≤x≤0.2) [3] and selected SrCe0.9Eu0.1O3-δ to fabricate supported tubular thin membranes to

52

maintain mechanical stability [86, 87]. The tubular SrCe0.9Eu0.1O3-δ membrane coated on NiO-

SrCeO3-δ support was applied to incorporate the WGS reaction and H2 separation [87].

In this chapter, a WGS membrane reactor was developed using a SrCe0.9Eu0.1O3-δ tubular

membrane to incorporate the WGS reaction and H2 separation into one unit. Results from both

thermodynamic equilibrium calculation and experiment were compared to show the effect of the

ceramic membrane on the WGS reaction. The thermodynamic equilibrium of the WGS reaction

was calculated for H2O/CO = 1/1 and 2/1. The improved CO conversion and H2 yield using the

membrane reactor was further confirmed by carrying out the WGS reaction under three

situations: (1) blank reference, (2) with Ni catalyst, and (3) with Ni catalyst and in situ H2

removal. In addition, appropriate operating temperature region without carbon formation for

each H2O/CO was addressed since carbon formation is detrimental to the WGS reaction as it

causes catalyst deactivation.

4.2 Experimental

Polycrystalline SrCeO3-δ and SrCe0.9Eu0.1O3-δ powders were prepared by conventional

solid-state reaction by mixing stoichiometric amounts of SrCO3 (99.9%, Alfa-Aesar), CeO2

(99.9%, Alfa-Aesar) and Eu2O3 (99.9%, Alfa-Aesar) powders, followed by ball milling and

calcining at 1300 oC. A NiO-SrCeO3-δ tubular support was fabricated using tape-casting (Pro-

Cast) followed by a rolling process. The tubular support was sealed at one end and pre-sintered.

SrCe0.9Eu0.1O3-δ was coated on the inner side of the pre-sintered support. The tubular membranes

were finally sintered at 1450 oC. A detailed preparation process has been discussed in our

previous work [74].

The membrane tube is about 17 cm long and 0.48 cm in diameter (Figure 4-1 A). An SEM

image after experiment shows that the membrane is dense and ~23 µm thick on a porous support

53

(Figure 4-1 B). A thermal insulator was applied to the bottom of the membrane tube, forming an

insulating region to drop the temperature and allow O-ring sealing of the tube [6]. The area of the

membrane above the insulator zone is considered the active area and is about 12 cm2. The WGS

reaction was carried out from 600 oC to 900 oC under 3% CO + 3% H2O and 3% CO + 6% H2O

(total flow rate of 20 sccm balanced by Ar). Gas flow rates were controlled by mass flow

controllers.

Figure 4-1 C shows a photo of the experimental setup. A thermo couple was placed axially

at the middle of the membrane to control temperature. Argon was used as tracer to detect

leakage. The reactants, CO and H2O, were flowed into the quartz chamber and exposed to the Ni

catalyst on the outside of the membrane tube. The reactor side effluents were analyzed by gas

chromatography (GC) (Varian CP 4900). Helium was used as a sweep gas on the inner side of

the membrane. The permeated H2 together with Ar (leakage) were analyzed by a mass

spectrometer (Q100MS Dycor Quadlink).

4.3 Results and Discussion

4.3.1 Thermodynamic Calculation

The thermodynamic equilibrium conditions of the WGS reaction were calculated using

Thermocalc software [88] with a total pressure of 1 atm. Figure 4-2 (A) and Figure 4-2 (B) show

the temperature dependence on species mole fraction with feed H2O/CO = 1/1 and 2/1,

respectively. The mole fractions of the reactants, CO and H2O, increase with increasing

temperature, which is attributed to the exothermic nature of the WGS reaction. Thus

thermodynamic equilibrium moves to the reactant side at elevated temperature.

Achieving a carbon deposition free operating temperature region is very important for the

WGS reaction since carbon formation may block the pores of the porous support and lead to

catalyst deactivation as well as cracking of the membrane. It is shown from the thermodynamic

54

equilibrium results that carbon formation is in general favored at low temperature and low

H2O/CO. Carbon will not form at temperatures higher than 590 oC with H2O/CO = 1. This shifts

to 550 oC for H2O/CO = 2. This is in agreement with the results by Xue et al. [89]. They reported

the risk of carbon formation due to side reactions increased as the H2O/CO decreased. The

formation of carbon with H2O/CO = 1 was thermodynamically favorable over the entire

temperature range examined (up to 500 oC). However, a carbon-free operation condition was

achieved at temperatures higher than 230 oC with H2O/CO = 3.

Low temperatures also favor CH4 and H2O formation, which is clearly shown in Figure 4-2

A. If no side reactions are considered, the mole fractions of CO and H2O are equal to each other

with H2O/CO = 1/1. Similarly, the mole fractions of H2 and CO2 are the same. However, the

mole fraction of CO is lower than that of H2O at 710 oC and the mole fraction of H2 is less than

that of CO2. This is attributed to the consumption of H2 to form CH4 and H2O. Higher H2O/CO

can extend the operating temperature of the WGS reaction to lower temperature. When the

H2O/CO is increased to 2/1, the formation of CH4 and H2O occurs below 640 oC. Therefore, the

WGS reaction should be carried out at temperatures higher than 710 oC and 640 oC with H2O/CO

= 1/1 and 2/1, respectively.

4.3.2 Experimental Conversion

The CO conversion was measured under three situations: (1) blank reference, (2) with Ni

catalyst, and (3) with Ni catalyst and in situ H2 removal. In the first two situations, two different

gas compositions were applied: 3% CO + 3% H2O or 3% CO + 6% H2O, while maintaining a

constant flow rate of 20 sccm balanced by Ar. Only 3% CO + 6% H2O gas composition was

applied in situation (3). For the blank reference, CO and H2O were fed into an empty quartz

reactor. For the WGS reaction with Ni catalyst, the tubular membrane was installed in the quartz

reactor with the Ni in the porous support being exposed to the reactants and the permeated side

55

was blocked so that no produced H2 was removed. Therefore, the membrane only functioned as a

catalyst in this situation with the reactant gas mix passing by the Ni catalyst in the support tube

surface. For the WGS reaction with Ni catalyst and in situ H2 removal, the permeated side was

connected to a mass spectrometer, so that the permeated H2 concentration could be analyzed.

The mole fractions of H2, CO, H2O and CO2 under these three reactor configurations are

shown in Figure 4-3 to Figure 4-5. These mole fractions do not include the Ar diluent. The CH4

concentration was below detection limits and ignored here. For the blank reference, the mole

fraction of H2 is the same as that of CO2 under both H2O/CO = 1/1 and 2/1 (Figure 4-3). The

mole fractions of CO and H2O equal to each other for H2O/CO = 1/1 as well.

For WGS with the Ni catalyst (Figure 4-4), the mole fractions of species are similar with

the thermodynamic data shown in Figure 4-2. The mole fractions of H2 and CO2 agree with each

other at elevated temperature and there is similar deviation at low temperatures. This indicates

the WGS reaction approaches thermodynamic equilibrium in the presence of the Ni catalyst in

the tubular support.

The mole fractions in Figure 4-5 are just for the species in the reactor side effluent. The

permeated H2 is not included. It is important to point out that no Ar was observed in the

permeated gas so the membrane was leak free. The H2 and CO2 mole fractions were significantly

different comparing Figure 4-4 (B) and Figure 4-5. Both were under the WGS with Ni and

H2O/CO = 2/1, but the data in Figure 4-5 were achieved with in situ H2 removal. The CO2 mole

fraction was higher in Figure 4-5 and increased with increasing temperature due to the in situ

removal of H2 overcoming the inherent thermodynamic limitation at high temperature. In

addition, while the H2 mole fraction decreased with increasing temperature in Figure 4-4 (B), it

was almost independent of temperature in Figure 4-5.

56

CO conversion is a very important parameter in the WGS reaction and needs to be

maximized in order to increase H2 production. The CO conversion is defined as follows:

XCO (%)= %1002 ×inCO

outCO

FF

(4-2)

where outCOF

2 and in

COF are CO2 output flux and CO input flux, respectively.

Figure 4-6 shows the XCO temperature dependence for different reactor configurations with

a H2O/CO = 1/1 and 2/1, respectively and compares with their thermodynamic equilibrium data.

The thermodynamic conversion decreases as the reaction temperature increases. This is

consistent with the mole fraction decrease of H2 and CO2 as temperature increases in Figure 4-2.

At any temperature, the XCO under thermodynamic equilibrium increases with increasing

H2O/CO from 1/1 to 2/1. A higher feed steam concentration moves the reaction (4-1) forward,

resulting in higher XCO.

For the blank reference under both H2O/CO = 1/1 and 2/1, XCO was significantly lower

than the thermodynamic values. The reaction rate increased with increasing temperature and the

XCO approached theoretical at higher temperatures.

With the Ni catalyst, the XCO was comparable to and consistent with the thermodynamic

values. The slight deviation, especially at lower temperatures, was attributed to side reactions

which could take place during the WGS process [89].

According to the thermodynamic data, higher H2O/CO results in higher XCO. Therefore,

the effect of in situ H2 removal on the XCO was investigated under H2O/CO = 2/1 (Figure 4-6).

Compared to the XCO with only the Ni catalyst, much higher XCO was achieved especially at high

temperatures (46% increase at 900 oC). Since the permeated H2 lowered the H2 concentration in

the product stream it moved the reaction further toward the product side resulting in higher XCO.

57

Therefore, the H2 membrane can help overcome the thermodynamic limitations and improve

XCO. It simultaneously increases H2 yield as well. This can also be applied to reduce the CO level

in the H2 gas produced from hydrocarbon fuels for proton-exchange membrane (PEM) fuel cells.

4.3.3 H2 Production

H2 production for the blank reference increased with increasing temperature, in agreement

with the XCO under the same conditions as shown in Figure 4-6. Figure 4-7 shows the H2

production with H2O/CO = 2/1 as a function of temperature. The H2 production with only the Ni

catalyst decreased with increasing temperature consistent with the thermodynamic data due to

the exothermic nature of the WGS reaction. The H2 permeation flux increased with increasing

temperature due to the higher ambipolar conductivity of SrCe0.9Eu0.1O3-δ at elevated

temperatures [3]. In addition, the H2 production in the reactor side effluent was essentially

temperature independent with in situ H2 removal. The total H2 production with in situ H2

removal is the sum of the H2 in the reactor side effluent and the permeated H2. It increased with

increasing temperature. The improvement was more significant at elevated temperatures

compared to the thermodynamic value. A 46% increase in total H2 production was achieved at

900 oC.

The H2 yield is defined:

100% (%) yield H 22 ×= in

CO

outH

FF

(4-3)

where outHF

2 and in

COF are the H2 production and CO input flux, respectively.

The H2 yield was plotted in Figure 4-8 as well as the H2/CO in the reactor side effluent

with H2 in situ removal and H2O/CO = 2/1. The total H2 yield is the sum of the permeated H2

yield and the H2 yield in the reactor side effluent. The H2 yield was in similar trend with the H2

58

production (Figure 4-7). The permeated H2 yield and total H2 yield were 32% and 92% at 900

oC, respectively. The reactor side effluent consisted of H2 and CO2 rich gases together with the

residual CO and H2O. The H2/CO increased from 3.9 to 7.6 when the temperature increased from

700 to 900 oC.

4.4 Conclusions

WGS reaction is constrained by thermodynamic equilibrium limitations. A tubular

SrCe0.9Eu0.1O3-δ H2 transport WGS membrane reactor was fabricated. The XCO, H2 production,

H2 yield, and the H2/CO in the reactor side effluent increase significantly with the WGS

membrane. A 46% increase in XCO and total H2 yield was achieved at 900 oC under 3% CO and

6% H2O compared to the thermodynamic equilibrium calculation, resulting in a 92% single pass

H2 production yield and 32% single pass yield of pure permeated H2. These results demonstrate

the efficiency of H2 membranes for the WGS reaction.

59

A

B

C

Figure 4-1. Morphology of thin film membranes and experimental setup. A) Membrane tube, B) SEM image of sintered membrane cross section and C) Photo of the WGS reactor showing gas flow.

60

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

Mol

e Fr

actio

n

Temperature (oC)

CO2

H2

H2O

CO

CH4

C

A

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

Mol

e Fr

actio

n

Temperature (oC)

H2O

CO2

CO

CH4C

H2

B

Figure 4-2. Thermodynamic equilibrium of WGS under A) H2O/CO =1/1 and B) H2O/CO =2/1.

61

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

CO2

H2

CO

H2O

Mol

e Fr

actio

n

Temperature (oC) A

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

550 600 650 700 750 800 850 900 950

CO2

H2

CO

H2O

Mol

e Fr

actio

n

Temperature (oC) B

Figure 4-3. Blank reference effluent gas composition as a function of temperature under A) H2O/CO =1/1 and B) H2O/CO =2/1.

62

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

CO2

H2

CO

H2O

Mol

e Fr

actio

n

Temperature (oC) A

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

CO2

H2

CO

H2O

Mol

e Fr

actio

n

Temperature (oC) B

Figure 4-4. Catalytic effluent gas composition as a function of temperature under A) H2O/CO =1/1 and B) H2O/CO =2/1.

63

0

0.1

0.2

0.3

0.4

0.5

550 600 650 700 750 800 850 900 950

CO2

H2

CO

H2O

Mol

e Fr

actio

n

Temperature (oC)

Figure 4-5. Catalytic effluent gas composition with in situ H2 removal as a function of temperature for H2O/CO =2/1 feed gas.

64

0

20

40

60

80

100

550 600 650 700 750 800 850 900 950

X CO (%

)

Temperature (oC)

Figure 4-6. Temperature dependence of XCO under 3% CO + 3 % H2O and 3% CO + 6% H2O (○ blank reference (H2O/CO=1/1), ● blank reference (H2O/CO=2/1), – — thermodynamic data (H2O/CO=1/1), -- thermodynamic data (H2O/CO=2/1), □ WGS with catalyst (H2O/CO=1/1), ■ WGS with catalyst (H2O/CO=2/1), ♦ WGS with catalyst and in situ H2 removal (H2O/CO=2/1)).

65

0

0.1

0.2

0.3

0.4

0.5

0.6

550 600 650 700 750 800 850 900 9500

0.01

0.02

0.03

0.04

0.05

H2 P

rodu

ctio

n (c

m3 /m

in)

Temperature (oC)

H2 Flux (cc/cm

2min)

Figure 4-7. H2 production under 3% CO + 6% H2O as a function of temperature for three reactor configurations (-- Thermodynamic H2 production ● Catalytic H2 production without H2 removal, ▼ Pure permeated H2 through membrane ■ H2 production in the reactor side effluent with in situ H2 removal ♦ Total catalytic H 2 production with in situ H2 removal).

66

0

20

40

60

80

100

550 600 650 700 750 800 850 900 9500

2

4

6

8

10

Temperature (oC)

H2 y

ield

(%)

H2 /C

O

permeated H2

H2 in reactor side effluent

total H2 yield

reactor side effluent H2/CO ratio

Figure 4-8. H2 yield and syngas H2/CO ratio as a function of temperature under 3% CO + 6% H2O and with in situ H2 removal.

67

CHAPTER 5 STABILITY OF SrCe1-XZrXO3-δ UNDER WATER GAS SHIFT REACTION CONDITIONS

5.1 Introduction

In chapter 4, a supported SrCe0.9Eu0.1O3-δ thin film membrane was used to incorporate the

WGS reaction and H2 separation into one unit and significantly increased CO conversion and H2

yield was achieved [87]. We demonstrated that H2 flux for Ni-SrCeO3-δ supported SrCe1-xEuxO3-δ

(0.05≤x≤0.2) thin films is proportional to [PH2]1/4 under wet and dry H2 conditions [6]. The H2

flux of SrCe0.9Eu0.1O3-δ is stable under wet H2 atmospheres but degrades under dry H2 conditions

[7]. The degradation under dry H2 is caused by phase decomposition forming CeO2 under low

PO2 conditions. The tubular SrCe0.9Eu0.1O3-δ membrane is also unstable under hydrocarbon

conditions possibly due to phase change and/or coking below 700 oC as shown in Figure 5-1.The

temperature gradient in the reactor furnace as a function of axial distance and the set point

temperature is caused by our experimental setup for these tubular membranes [6]. In contrast, the

SrCe0.7Zr0.2Eu0.1O3-δ membrane maintains good integrity after exposure in methane with 18 %

steam. This indicates the partial substitution of Zr for Ce can improve the chemical stability of

SrCeO3-δ which is attributed to the higher stability of SrZrO3-δ against carbon dioxide than

SrCeO3-δ below 800 oC [90, 91]. In addition, Zr has been used to improve the chemical stability

of BaCeO3-δ system [8-12] and SrCe0.95Yb0.05O3-δ [13]. The thermo-chemistry of SrCeO3-δ [92-

94] and its stability under CO2 [93, 95] and water [96, 97] have been reported. However, the

stability of Zr-doped SrCeO3-δ hasn’t been systemically studied yet especially under various

hydrocarbon gas conditions.

In this chapter, four different zirconium dopant concentrations were examined to improve

the stability of SrCeO3-δ under WGS reaction atmospheres (CO/CO2, H2/H2O). The possible

SrCeO3-δ decomposition mechanism was investigated using SrCe0.8Zr0.2O3-δ (denoted SCZ82

68

hereafter) powder. We fabricated a supported SrCe0.7Zr0.2Eu0.1O3-δ (denoted SCZE721 hereafter)

thin film membrane using NiO-SCZ82 as the supported structure to increase H2 permeation.

Finally, its chemical stability was investigated by measuring the H2 permeation under WGS

conditions as a function of time.

5.2 Experimental

To investigate the zirconium dopant effect on the chemical stability under CO/CO2, five

different compositions (Zr=0, 2.5 at%, 5 at%. 10 at% and 20 at%) of SrCe1-xZrxO3-δ powders

were synthesized using conventional solid-state reaction by mixing stoichiometric amounts of

SrCO3-δ (99.9%, Alfa-Aesar), CeO2 (99.9%, Alfa-Aesar), and ZrO2 (99.7%, Alfa-Aesar)

powders, followed by ball milling and calcining at 1300 oC. These powders were exposed to

2.8% CO and 5.6% H2O (balanced by Ar) atmosphere at different temperatures for 19 hours. The

atmosphere finally became a mixture of CO, H2, H2O and CO2 due to the WGS reaction. In

addition, SCZ82 was exposed to 3.0% dry CO, 2.8% CO2/5.6% H2O, 3.0% dry CO2, 6.0% H2O,

and 2.8% H2/2.8% H2O atmospheres for 19 hours to investigate the decomposition mechanism.

Phase compositions were analyzed by X-ray diffraction (XRD) (XRD Philips APD 3720) with

CuKα radiation. The 2θ value chosen was from 20 to 80 degree with a step size of 0.03 degree/s.

A detailed preparation process of the supported thin film was discussed in chapter 3 and

our previous work [74]. For H2 permeation measurements, the tubular H2 membranes were

installed in a high temperature reactor apparatus which has been described previously [6]. H2

permeability dependence on time was investigated under 5.0% CO and 3.0% H2O atmosphere at

900 oC to further evaluate stability under operating conditions. The permeated gases were

analyzed by a mass spectrometer (Q100MS Dycor Quadlink).

69


5.3.1 Stability under Wet CO Conditions

Figure 5-2 displays the XRD patterns of SrCeO3-δ before and after the stability experiment.

The as-calcined powder had an orthorhombic perovskite structure. SrCeO3-δ phase remained

predominant at 900 oC, but decomposed into SrCO3 and CeO2 at temperatures between 600 oC

and 800 oC. XRD results show that SrCeO3-δ is unstable at low temperatures with respect to

SrCO3 under CO/H2O atmospheres, which is in agreement with Shirsat’s results [92].

Figure 5-3 (A) shows the XRD patterns of the as-calcined powder for the different

zirconium dopant concentrations at 1300 oC. No significant secondary phase was detected. All

the compositions were orthorhombic perovskite structure. The lattice parameters decrease with

increasing Zr dopant concentration (Figure5- 3 (B)) due to the smaller ionic radius of Zr4+. The

lattice parameters for SrCeO3-δ are a=8.575 Å, b=6.122 Å, c=6.000 Å, essentially identical to

Wei et al. results [98]. The lattice parameters are approximately a linear function of Zr dopant

concentration, following Vegard’s law. Figure 5-4 shows the XRD patterns of SrCe1-xZrxO3-δ at

800 oC after the stability experiment, exposed to 2.8% CO and 5.6% H2O for 19 hours. For Zr

dopant concentrations of 2.5 at% and 5 at%, the intensity of the CeO2 (111) reflection (2θ≈28.7o)

was comparable to the main peak of the perovskite (2θ≈29.7o), which indicates a significant

amount of CeO2 was formed. For Zr dopant concentration of 10 at% and 20 at%, the intensity of

CeO2 peak was much lower, showing higher stability under CO and CO2 atmospheres. This

demonstrates higher stability can be achieved with higher Zr dopant concentration. Therefore, 20

at% Zr dopant was used for further study.

To specify the temperature effect on SrCeO3-δ decomposition, the stability experiment was

carried out on SCZ82 at 800, 825, 850, 865, 875, 900 and 940 oC and their XRD patterns are

70

shown in Figure 5-5. The relative intensity of the strongest peaks of CeO2 and perovskite phase

is calculated and shown in table 5-1. The intensity of the CeO2 peak became much lower at

higher temperatures, with the relative intensity decreasing from 34% to 5% when temperature

was increased from 800 oC to over 865 oC. For the as-calcined SCZ82 powder, the relative

intensity of the secondary phase, CeO2, is 5%, which is in agreement with Mather’s results [97].

Generally, X-ray phase-pure material could not be obtained for nominally stoichiometric

SrCeO3-δ or Sr deficient samples synthesized by solid state reaction [42, 97]. At temperatures

higher than 865 oC, the relative intensity of CeO2 (111) is approximately equal to that in the as-

calcined powder, indicating the SCZ82 is stable at that temperature. In view of thermodynamics,

the formation reaction of SrCO3 and CeO2 from SrCeO3 and CO2 associates with a decrease in

entropy; thus, it is not favored at elevated temperatures. Therefore, to avoid SCZ82

decomposition, the WGS reaction should be carried out at temperatures higher than 865 oC.

The improvement in chemical stability of strontium zirconate can be explained by basicity

of the oxide and tolerance factor. A high basicity of the oxide is advantageous for the stability of

protonic defects but basic oxides are expected to react easily with acidic gases such as CO2 [99].

The stability against the formation of carbonates and hydroxides increases in the order cerate

zirconate with decreasing the stability of protonic defects.

The tolerance factor, t introduced by Goldschmidt, is a measure of the “cubic-ness” of the

perovskite [100]:

)(2 OB

OA

rrrrt+

+= (5-1)

where Ar , Br , Or are the ionic radii of the A-site cation, the B-site cation and the oxygen anion,

respectively. The tolerance factor has been used to explain the stability of many perovskites [95,

71

101-103]. When t equals unity, the structure is predicted to be perfect cubic. Lower values of t

correspond to lower symmetry. Using ionic radii found in Shannon’s paper [104], the tolerance

factor was calculated for SrCe1-xZrxO3-δ (x=0, 2.5%, 5%, 10% and 20%), Table 5-2. The

tolerance factor is less than 1 for all compositions studied but increases with increasing

zirconium dopant concentration.

The improvement of chemical stability with increasing tolerance factor is in accordance

with Yokokawa’s result [105]. They reported the stabilization energy of AIIMIVO3 perovskite,

which was related to the tolerance factor. Generally, the closer the value of the tolerance factor to

unity, the higher the chemical stability of the perovskite structure.

5.3.2 Decomposition Mechanism

The WGS atmosphere consists of CO, CO2, H2 and H2O gases. All of these gases might

react with the powder and cause decomposition. To identify which species, CO, CO2, H2 and

H2O, contribute to the decomposition, SCZ82 was exposed to 3% CO, 2.8% CO2/5.6% H2O, 3%

CO2, 6% H2O and, 2.8% H2/2.8% H2O, for 19 hours at 800 oC, respectively. Figure 5-6 shows

the XRD patterns. The pattern (a) is similar with that of the as-calcined powder, pattern (d) in

figure 5-3. No significant secondary phase was detected, indicating SCZ82 is stable under dry

CO atmosphere. Therefore, CO is unlikely to cause the decomposition. For patterns (b) and (c),

significant SrCO3 and CeO2 phases were detected and the intensity of CeO2 phase increased

dramatically compared to the pattern of the as-calcined powder. The intensity ratios between the

strongest peaks of CeO2 and perovskite phase in pattern (b) and (c) are, 41% and 47%,

respectively. They are higher than the ratio (34%) under 2.8% CO/5.6% H2O at 800oC. The

CO2% among these three conditions follows the relation: 2.8% CO/5.6% H2O < 2.8% CO2/5.6%

72

H2O < 3% CO2. The intensity ratio between the strongest peaks of CeO2 and the perovskite

phase follows this same order. Therefore, CO2 is responsible for the decomposition.

H2O and H2 may also contribute to decomposition under water gas shift (WGS) conditions.

Therefore, the effect of steam and wet H2 on the stability of SCZ82 was investigated. Patterns (d)

and (e) show the results after the stability experiment under 6% H2O and, 2.8% H2/5.6% H2O at

800oC, respectively. Both XRD patterns are similar with that of the SCZ82 as-calcined powder,

indicating SCZ82 is stable under these conditions. This is in agreement with Mather’s result

[106]. They reported SrCeO3 was stable under PH2O= 1.6 atm at much lower temperatures, 120-

174 oC. The formation reaction of CeO2 and Sr(OH)2 from SrCeO3 and H2O associates with a

decrease in entropy and is not favored at elevated temperatures. Therefore, steam and wet H2 do

not cause SCZ82 decomposition and CO2 plays the dominant role. Moreover, since no SrO and

Sr(OH)2 phases were detected under steam and wet H2 atmospheres, decomposition takes place

most likely by SrCeO3 reacting directly with CO2 through reaction (5-2).

2323 CeOSrCOCOSrCeO +↔+ (5-2)

5.3.3 Hydrogen Permeability

Figure 5-7 shows the H2 permeation versus time at 900 oC for SrCe09Eu0.1O3-δ and

SCZE721 membranes under 5% CO and 3% H2O atmospheres with a total flow rate of 20

cm3/min, balanced by Ar. In both cases, the H2 permeation increases in the first couple of hours,

which we attributed to the reduction of the NiO to Ni in the support structure. After that, the H2

permeation through the SrCe09Eu0.1O3-δ membrane degraded significantly with a degradation rate

of 1.8 %/hr. In contrast, the H2 permeation flux through the SCZE721 membrane was essentially

stable with a degradation rate of only 2.4×10-3 %/hr. The stable H2 permeation flux verifies the

73

stability enhancement by zirconium dopant. However, this comes at the expense of lower

permeation flux since the protonic conductivity of SrCeO3 is greater than SrZrO3 [37].

5.4 Conclusions

SrCeO3 is unstable under WGS atmospheres. CO2 is found to be the main cause of

decomposition, forming SrCO3 and CeO2. The H2 permeation flux of SrCe0.7Zr0.2Eu0.1O3-δ

membrane was essentially stable under WGS conditions at 900 oC. Its degradation rate is 1000

times lower than that of a SrCe0.9Eu0.1O3-δ membrane under the same conditions. This confirms

the stability of SrCeO3-δ can be improved by partially substituting Ce with Zr, which increases its

tolerance factor and decreases its basicity.

74

Figure 5-1. Temperature profile and photograph of the membranes A) temperature profile along the membrane [6] and B) Photograph of the SrCe0.9Eu0.1O3/Ni-SrCeO3 and SrCe0.9Eu0.1O3 /Ni-SCZ82 H2 membranes after exposure to methane with 18% steam.

A

B

75

Figure 5-2. XRD pattern of SrCeO3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h

(*) perovskite phase, (o) CeO2 phase, (+) SrCO3 phase.

76

A

0 5 10 15 20

Latti

ce P

aram

eter

s (A

)

ac

b

Zr (at%)

Lattice Parameter (A

)

6.15

6.10

6.05

5.95

6.00

5.90

8.65

8.60

8.55

8.50

8.40

8.45

B

Figure 5-3. XRD Pattern and their lattice parameters of SrCe1-xZrxO3-δ as-calcined at 1300 oC A) XRD pattern and B) lattice parameters as a function of dopant concentrations.

77

Figure 5-4. XRD Pattern of SrCe1-xZrxO3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h at

800 oC (o CeO2 phase, + SrCO3 phase).

78

Figure 5-5. XRD Pattern of SrCe0.8Zr0.2O3-δ after exposure to 2.8% CO and 5.6% H2O for 19 h (o CeO2 phase).

79

Figure 5-6. XRD Pattern of SrCe0.8Zr0.2O3-δ after stability experiment at different atmospheres at

800 oC (a) 3% CO, (b) 2.8% CO2 & 5.6% H2O, (c) 3% CO2, (d) 6% H2O, (e) 2.8% H2 & 5.6% H2O (o CeO2 phase, + SrCO3 phase).

80

0 2 4 6 8 10 12

H2 P

rodu

ctio

n (c

m3 /m

in)

Time (hrs)

SrCe0.9

Eu0.1

O3

(1.8 %/hr)

SrCe0.7

Zr0.2

Eu0.1

O3

(2.4*10-3 %/hr)

0.25

0.20

0.15

0.10

0.05

Figure 5-7. H2 permeation as a function of time under 5% CO and 3% H2O at 900 oC.

81

Table 5-1. Intensity ratios between the strongest peaks of CeO2 and SCZ82. Temperature (oC)

800 825 850 865 875 900 940 1300(as- calcined)

ICeO2/ISCZ82 34% 30% 24% 6% 5% 5% 5% 5% Table 5-2. Tolerance factors of SrCe1-xZrxO3-δ. x value 0 2.5% 5% 10% 20% t value 0.885 0.886 0.888 0.891 0.897

82

CHAPTER 6 HYDROGEN PERMEATION OF THIN SUPPORTED SrCe0.7Zr0.2Eu0.1O3-δ MEMBRANES

UNDER DIFFERENT OXYGEN PARTIAL PRESSURE

6.1 Introduction

In recent years, high-temperature mixed protonic and electronic conducting ceramics have

attracted considerable attention as potential membranes for H2 gas separation from hydrocarbon

fuels [21, 25, 71, 95]. Numerous SrCeO3-δ and BaCeO3-δ based perovskite oxides with

multivalent cation dopants have been reported [3, 26-30, 48, 107]. In order to enhance the

ionization of H2, cermet membranes with a continuous metallic phase were also studied [31, 32].

In addition, complex perovskites in the form of A2B1+x′B1-x″MxO6-δ or A3B1+x′B2-x″Mx2O9-δ [5,

33-36] have been developed to increased stability of the perovskite oxides. We previously

reported the H2 permeation of the SrCe0.9Eu0.1O3-δ [6] and demonstrated that its thermodynamic

stability was improved by partial substitution of Zr onto the Ce-site [108]. In chapter 5, the

stability of SrCeO3-δ was improved using 20 mol% Zr dopant even under WGS reaction

conditions [108, 109].

In this chapter, the H2 permeation of the SrCe0.7Zr0.2Eu0.1O3-δ thin film on NiO-

SrCe0.8Zr0.2O3-δ support was investigated in details as a function of temperature, H2 flow rate,

membrane thickness, and H2 and/or steam partial pressures. The activation energies under

different oxygen partial pressure were discussed. The long term stability of H2 permeation under

wet H2, WGS reaction, and SRM was investigated as well.

6.2 Experimental

6.2.1 Membrane Fabrication

A detailed preparation process of the SrCe0.7Zr0.2Eu0.1O3-δ thin film on NiO-SrCe0.8Zr0.2O3-

δ support has been discussed in chapter 3 and our previous work [74].

83

6.2.2 Membrane Morphology

The membrane tube was ~17 cm long and 0.48 cm in diameter (Figure 6-1 (A)). SEM

images after experiment show that the membrane is dense and pinhole free (Figure 6-1 (B)) on

the surface. It is ~33 µm thick (Figure 6-1 (C)). The effective area for H2 permeation in the

tubular membrane is 12 cm2.

6.2.3 Membrane Permeation

For H2 permeation measurements, the tubular H2 membranes were installed in a high

temperature reactor apparatus (Figure 6-1 (D)) which was previously described [74]. The outer

side of the membrane (feed) was exposed to H2 (99.999%) diluted to the desired concentration

using Ar (99.999%) with a 20 cm3/min total flow rate. The total flow rate was variable when the

flow rate effect on the H2 permeation was investigated. For wet gas flow, 3 vol % water vapor

was picked up by flowing the feed gas through a water bubbler at 25 oC. Greater concentration of

water vapor was achieved by gasifying a desired amount of water provided by a syringe pump.

The inner side (sweep) of the membrane was flushed with He at 20 cm3/min. The flow rates of

H2, Ar, and He were controlled by mass flow controllers. Ar was used as a tracer to determine

whether the membrane were pinhole/leak free [6, 74]. The components of the permeated gases

on the sweep side were measured using a mass spectrometer. The gases from the exhaust on the

feed side were measured using GC to verify the mass balance.

6.3 Result and Discussion

6.3.1 Heat Treatment

Before the as-sintered membrane was used for permeation, the NiO in the support had to

be reduced to create porosity. In this work, the outer side of the membrane was treated by 5

cm3/min H2 and 15 cm3/min Ar mixed with 3% H2O at 900 oC overnight. Meanwhile, 20

84

cm3/min He was fed to the inner side of the membrane. Finally, a flat H2 permeation was

observed indicating the NiO was completely reduced.

6.3.2 Flow Rate Effect on H2 Permeation

Figure 6-2 shows H2 permeation and recovery under different flow rates at 900 oC. The

feed gas composition was 48.5% H2, 48.5% Ar and 3 % H2O to maintain a constant feed H2

partial pressure. The H2 permeation increased initially with total feed flow rate but was

essentially constant when the feed flow rate was over 20 cm3/min. Tong et al [110] and Li et al.

[111] have observed similar results except that their H2 permeation reached a plateau level at

higher total feed flow rates. This is due to the higher permeability of their membranes. The

amount of the permeated H2 at lower feed flow rates can cause a larger decrease of H2 partial

pressure in the feed gas, resulting in a bigger drop in the driving force for H2 permeation. In

contrast, at higher feed flow rates, the same amount of permeated H2 only causes a smaller

decrease in the driving force. Therefore, the H2 permeation flux barely changed. H2 recovery was

defined as follows:

%100H feed

H permeated[%]recovery H2

22 ×= (6-1)

As shown in Figure 6-2, the H2 recovery decreased as the flow rate increased, probably due

to the decrease in the resident time. Thus, the opportunity of H2 atoms to be absorbed on the

membrane surface and diffused through the membrane is smaller. Consequently, the H2 recovery

decreased.

6.3.3 H2 Permeation as a Function of Thickness

Figure 6-3 shows the H2 permeation flux of SrZr0.2Ce0.7Eu0.1O3-δ as a function of thickness

and feed H2 concentration at 900 oC [112]. The H2 permeation flux linearly increases with

decreasing membrane thickness. This is consistent with the Wagner equation (6-2), which shows

85

that H2 permeation through a proton conducting membrane is inversely proportional to the

membrane thickness [5].

++−= ∫∫ •••••••

//2

/2

2/

//2

/2

2ln)(

2ln

41

22

H

H OO

O

O OOO

P

P HeVOHt

P

P OVOHtOH PdtttF

RTPdttF

RTL

J σσ (6-2)

This indicates that the H2 permeation of SrZr0.2Ce0.7Eu0.1O3-δ membranes is controlled by bulk

diffusion down to 17 micrometer thick, in agreement with the H2 permeation through the dense

SrCe0.95Yb0.05O3-δ membranes, which has been reported to be controlled by bulk diffusion at 950

K even for 2 µm films [73].

6.3.4 Effect of Temperature, H2 and H2O Partial Pressure in the Feed Side on H2 Permeation

Figure 6-4 shows the H2 permeation flux of the SrCe0.7Zr0.2Eu0.1O3-δ membrane as a

function of feed gas H2 partial pressure and temperature under dry H2 and 97% H2/3% H2O

conditions, respectively. fHP

2and P

HP2are the H2 partial pressure at the feed side and permeated

side. H2 permeation flux increased with temperature under both conditions which is attributed to

the increase in ambipolar conductivity of the SrCe0.7Zr0.2Eu0.1O3-δ membrane[3, 48]. In addition,

the H2 permeation was proportional to the transmembrane H2 partial pressure gradient with a 1/4

dependence.

When the membrane is exposed to a H2 atmosphere, protons and electrons are the

dominating defects [64]. We obtain equation (6-3) after integrating the Wagner equation

))()((1 4/14/122

PH

fHOH PP

LJ

O−∝• (6-3)

A maximum H2 permeation flux of 0.23 and 0.21 cm3/cm2 min was obtained at 900 oC for

100% H2 and 97% H2/3% H2O conditions, respectively. This is very close to the ~0.26 cm3/cm2

86

min maximum H2 permeation flux through the 30 µm thick SrCe0.7Zr0.2Eu0.1O3-δ membrane

under 97% H2/3% H2O conditions at 900 oC [112].

Compared to the dry H2 condition, the presence of H2O increases the O2 partial pressure

resulting in decrease in the electronic conduction of the SrCe0.7Zr0.2Eu0.1O3-δ. Therefore, the H2

permeation decreases. To further address the steam concentration effect, the H2 permeation was

investigated under three steam concentrations and dry H2 condition with a constant of H2 flow

rate of 20 cm3/min (Figure 6-5). It decreased consistently with increasing steam concentration.

The maximum H2 permeation flux decreased from 0.23 to 0.18 cm3/cm2 min when the steam

partial pressure increased from 0 to 30% at 900 oC.

6.3.5 Activation Energy

The activation energy can be obtained from the Arrhenius plot using the data in Figure 6-2

and Figure 6-4 and are listed in Table 6-1 and 6-2. For both dry and wet H2 conditions, the

activation energy decreases with increasing H2 partial pressure. In addition, when the ratio of

steam partial pressure to H2 partial pressure decreases, the activation energy decreases as well.

Since the H2 permeation is proportional to H2 partial pressure to the 1/4 power, the wet H2

condition corresponds to region V in the defect equilibrium diagram according to Song et al.

[27]. In this region, the electron concentration equals to the proton concentration and they are the

dominating defect species instead of oxygen vacancy. According to Guan et al. [72], the

activation energies of proton and electron mobility in perovskite are 0.4-0.6 eV and 1 eV,

respectively. In addition, the H2 permeation through SrCe0.9Eu0.1O3 was limited by electronic

conduction [6]. Therefore, the activation energies at high H2O concentration and low H2 partial

pressure are close to 1 eV. However, O2 partial pressure decreases as increasing H2 partial

pressure, resulting in an increase in electron concentration and electronic conduction. Therefore,

87

the H2 permeation is less limited by electronic conduction resulting in decrease in the activation

energy. Similarly, when the ratio of steam partial pressure to H2 partial pressure decreases, the

activation energy decreases.

6.3.6 Long Term Stability

The membrane’s long term stability was investigated by measuring the H2 permeation as a

function of time under wet H2, and conditions of WGS reaction and SRM (Figure 6-6). The H2

permeation was stable throughout the experiment for all three conditions. The H2 permeation is

~0.98 cm3/min under 5 cm3/min H2, 15 cm3/min Ar and 3% H2O at 850 oC, almost identical to

the result (~0.94 cm3/min) in Figure 6-2 (B). The stable H2 permeation under the conditions of

WGS reaction and SRM confirms the stability improvement of the SrCe0.7Zr0.2Eu0.1O3-δ using Zr

as a dopant.

6.4 Conclusions

H2 permeation through supported thin-film SrCe0.7Zr0.2Eu0.1O3-δ membranes was

investigated. Permeation flux was proportional to the transmembrane H2 partial pressure gradient

with a 1/4 dependence and controlled by bulk diffusion. A maximum H2 permeation flux of 0.23

and 0.21 cm3/cm2 min was obtained for the 33 μm thick SrCe0.7Zr0.2Eu0.1O3-δ membrane at 900

oC and 100% H2 and 97% H2/3% H2O in the feed gas, respectively. Permeation flux decreased

with increasing steam partial pressure. The activation energy decreased with increasing H2

partial pressure and/or decreasing steam partial pressure. Permeation flux through the

SrCe0.7Zr0.2Eu0.1O3-δ membrane was stable under wet H2, and conditions of WGS reaction and

SRM.

88

A

B C

D

Figure 6-1. SrCe0.7Zr0.2Eu0.1O3-δ membrane and experimental setup A) membrane after H2 permeation experiment, B) Surface morphology, C) Cross section and D) Experimental setup.

89

10

15

20

25

30

35

5 10 15 20 25 30 35 40 45Flow Rate (cm3/min)

J H2 (c

m3 /c

m2 m

in) Percentage (%

)

0.05

0.15

0.10

0.20

0.30

0.25

Figure 6-2. H2 permeation and H2 recovery as a function of feed flow rates.

90

0.01 0.02 0.03 0.04 0.05 0.06

J H2 (c

m3 /c

m2 m

in)

1/Thickness (µm-1)

24.3% H2

48.5% H2

72.8% H2

97.0% H2

0.40

0.35

0.30

0.25

0.20

0.15

0.10

Figure 6-3. H2 permeation vs thickness at 900oC.

91

0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5 0.550

0.04

0.08

0.12

0.16

0.2

0.24

J H

2 (cm

3 /min

)

PH2f

1/4-PH2p

1/4 (atm1/4)

J H2 (cm

3/cm2 m

in)

900 oC

850 oC

800 oC

750 oC

700 oC

0.20

3.0

2.0

1.0

1.5

0.5

0

2.5

A

0.2 0.3 0.4 0.5 0.60

0.04

0.08

0.12

0.16

0.2

0.24

J H

2 (cm

3 /min

)

PH2f

1/4-PH2p

1/4 (atm1/4)

J H2 (cm

3/cm2 m

in)

900 oC850 oC

800 oC

750 oC

700 oC

0.20

3.0

2.5

2.0

1.5

1.0

0.5

0

B

Figure 6-4. H2 permeation as a function of H2 partial pressure and temperature A) under dry H2 condition and B) under 3% H2O.

92

650 700 750 800 850 900 950

dry H2

3% H2O+H

210% H

2O+H

230% H

2O+H

2

0

0.04

0.08

0.12

0.16

0.2

0.24J

H2 (c

m3 /m

in)

Temperature (oC)

J H2 (cm

3/cm2 m

in)

0.20

3.0

2.5

2.0

1.5

1.0

0.5

0

Figure 6-5. H2 permeation as a function of feed steam concentration and temperature.

93

Figure 6-6. H2 permeation as a function of time (a) under 8 cm3/min CH4, 12 cm3/min Ar and CH4/H2O=1:2, at 850 oC, (b) under 5 cm3/min H2, 15 cm3/min Ar and 3% H2O at 900 oC and (c) under 10 cm3/min CO, 10 cm3/min Ar and CO/H2O=1:2, at 900 oC.

94

Table 6-1. Activation energy as a function of H2 partial pressure under dry H2 and H2/3% H2O conditions balanced by Ar.

Dry H2 Ea (eV) H2/3% H2O Ea (eV) 25% 0.86 24% 0.95 50% 0.74 48% 0.84 75% 0.63 72% 0.73

Table 6-2. Activation energy as a function of H2O partial pressure with a constant H2 flow rate of 20 cm3/min.

30% H2O+70% H2 10% H2O+90% H2 3% H2O + 97% H2 100% H2 1.06 eV 0.91 eV 0.65 eV 0.58 eV

95

CHAPTER 7 SrCe0.7Zr0.2Eu0.1O3-δ-BASED HYDROGEN TRANSPORT WATER GAS SHIFT REACTOR

7.1 Introduction

In chapter 4, tubular SrCe0.9Eu0.1O3-δ membrane on NiO-SrCeO3-δ support was applied to

incorporate the WGS reaction and H2 separation. CO conversion and H2 yield were significantly

improved compared to the thermodynamic equilibrium. However, that work just focused on the

comparison of the CO conversion under three different reactor configurations. The CO

concentrations and flow rates were very low. In addition, this membrane is unstable under the

WGS reaction conditions [108]. In chapter 5, the chemical stability of SrCeO3-δ was significantly

improved using Zr dopant [108, 109]. H2 permeation property of tubular SrCe0.7Zr0.2Eu0.1O3-δ

membranes on NiO-SrCe0.8Zr0.2O3-δ supports [86] was discussed in chapter 6, In this chapter, the

effect of the SrCe0.7Zr0.2Eu0.1O3-δ membrane reactor on the WGS reaction was investigated in

various conditions. Its performance was investigated under high CO concentrations, high flow

rates, and different CO concentration and H2O/CO. Its long term stability under WGS reaction

condition was investigated as well.

7.2 Experimental

Polycrystalline SrCe0.8Zr0.2O3-δ and SrCe0.7Zr0.2Eu0.1O3-δ powders were prepared by

conventional solid-state reaction. NiO-SrCe0.8Zr0.2O3-δ tubular supports were fabricated by tape-

casting and rolling techniques. Detailed powder synthesis and membrane fabrication have been

discussed in chapter 3 and our previous works [74, 86].

The dense SrCe0.7Zr0.2Eu0.1O3-δ membrane used in this experiment was ~33 µm thick and

coated on NiO-SrCe0.8Zr0.2O3-δ support. Figure 7-1(A) shows the SEM image of the membrane

cross section after experiment. Based on our previous experience, its active area is about 12 cm2

96

[6, 86]. The influence of temperature, feed CO concentration, feed flow rate and H2O/CO were

evaluated in terms of CO conversion and H2 recovery.

The experimental setup is the same as that shown in Figure 4-1 (C). Figure 7-1 (B) shows

the schematic view of the membrane reactor. Outer side of the membrane (feed) was exposed to

CO and steam diluted to the desired concentration using Ar. Ar was also used as a tracer to

detect leakage. It is noted that no Ar leakage was detected during the experiment. Steam was

achieved by gasifying desired amount of water provided by a syringe pump. The inner side

(sweep) of the membrane was flushed with He at 20 cm3/min, in co-current flow with the feed

gas. The flow rates of CO, Ar and He were controlled by mass flow controllers. The reactants,

CO and H2O, were flowed into the quartz chamber and exposed to the Ni catalyst on the outside

of the membrane. The unreacted steam in the reactor side effluent was condensed by a cold trap

filled with ice prior to being analyzed by GC. The concentrations of the permeated H2 in the

sweep gas, He, were analyzed by a mass spectrometer.


7.3.1 Heat Treatment of the Membranes

Before the WGS reaction experiment, the membrane was heat treated to reduce the NiO in

the support by exposing it to 5 cm3/min H2 and 15 cm3/min Ar mixed with 3% H2O under 900 oC

overnight. This heat treatment has been described in reference [86].

7.3.2 H2O/CO Effect on CO Conversion

The temperature and H2O/CO effect on the WGS reaction were investigated. Figure 7-2

shows the mole fractions of the species in the reactor side effluent as a function of temperature

with various H2O/CO. The solid lines are experimental results and the dashed lines are

thermodynamic calculation data from reference [87]. Feed gas composition was 10 cm3/min CO,

10 cm3/min Ar and desired amount of steam based on the H2O/CO. To compare with the

97

thermodynamic data, Ar diluent is not included when the mole fraction was plotted. H2 and CO2

have the same profile for the thermodynamic calculation data regardless of the H2O/CO since

they are produced with the same rate. However, the mole fraction of CO2 is higher than that of

the H2 in experimental data due to the in situ removal of H2 through the H2 membrane. In

addition, the ratio of products, H2 and CO2, to reactants, H2O and CO, is greater for the

experimental results than for the thermodynamically calculated data. This indicates the WGS

reaction moves to the product side with the H2 membrane. A larger amount of CH4 than the

thermodynamic data was detected which is probably due to the Ni catalyst since Ni is an

effective catalyst for the methanation reaction [113, 114]. We didn’t detect CH4 in chapter 4 and

our previous work [87] because low concentration of CO was used and the CH4 concentration

was below detection limit.

CO conversion can be calculated from the mole fraction data and is plotted in Figure 7-3.

The CO conversion is defined in equation (4-2).

The dashed lines are thermodynamic calculation data from reference [87]. Based on the

thermodynamic calculation, the XCO decreases with increasing temperature due to the exothermic

nature of the WGS reaction. In contrast, it increases with increasing temperature in the

experimental results for any H2O/CO because of the in situ H2 removal through the H2

membrane. It also increases with increasing H2O/CO. A XCO of 83.6% and 90.2% was achieved

under 900 oC with H2O/CO = 1/1 and 2/1, respectively, 77% and 44% increase compared to the

thermodynamic calculation data.

7.3.3 H2O/CO Effect on H2 Production

Figure 7-4 shows the H2 production as a function of temperature under various H2O/CO

with 10 cm3/min CO, cm3/min Ar and desired amount of steam as feed gas. The H2 production

from thermodynamic calculation decreases with temperature. In contrast, the H2 permeation

98

increases with temperature due to the higher ambipolar conductivity of the membrane at elevated

temperature. The H2 flux in the reactor side effluent is not a significant function of temperature.

The H2 total production is the sum of permeated H2 and H2 flux in the reactor side effluent. It

increases with temperature. The H2 production curve is in a similar trend for any H2O/CO. The

permeated H2 decreases with increasing H2O/CO. A maximum H2 production of 8.1, 8.7 and 8.9

cm3/min was achieved under 900 oC with H2O/CO = 1/1, 1.5/1 and 2/1, respectively. Compared

to the thermodynamic calculation data, they are 73% and 42% improvement for H2O/CO = 1/1

and 2/1, respectively. If there is no side reaction, the improvement in XCO and H2 production

should be equal. However, the improvement of the XCO is higher compared to that of the H2

production. This is attributed to the methanation reaction consuming part of the produced H2,

which is clearly shown in Figure 7-2. In addition, the XCO is derived from GC measurement. The

H2 production is derived from both GC and MS measurement simultaneously. The XCO and H2

production data are pretty consistent within experimental error.

7.3.4 H2O/CO Effect on H2 Production and H2/CO

The H2O/CO effect on H2 yield and H2/CO in the reactor side effluent is plotted in Figure

7-5. The H2 yield is defined in equation (4-3). The H2 total yield is the sum of permeated H2

yield and H2 yield in the reactor side effluent. The H2 yield curves are in a similar trend with the

H2 production curves in Figure 7-4. Both the H2 yield and H2/CO increase with H2O/CO. The

H2/CO from thermodynamic calculation decreases with temperature, in agreement with the XCO.

When temperature is increased from 600 to 900 oC, it decreases from 1.3 to 0.9 and 2.7 to 1.7 for

H2O/CO = 1/1 and 2/1, respectively. In contrast, the H2/CO in the reactor side effluent increases

from 2.4 to 4.4 and 3.8 to 8.3 at the same temperature region for H2O/CO = 1/1 and 2/1,

respectively. The main composition of the products from the WGS reaction is syngas. The H2

99

membrane supplies an option to adjust the H2/CO prior to the Fischer-Tropsch process based on

desired products.

7.3.5 Flow Rate Effect on WGS Reaction

The flow rate effect on the WGS reaction was carried under 900 oC with a total flow rate

from 15 to 60 cm3/min and H2O/CO = 2/1. As shown in Figure 7-6, the XCO slightly decreases

with increasing flow rate, which is possibly due to the shorter residence time at higher flow rate.

Whereas the H2/CO in the reactor side effluent is not a significant function of total flow rate, the

permeated H2, H2 in the reactor side effluent, and H2 total production increases with total flow

rate.

7.3.6 CO Concentration Effect on WGS Reaction

Furthermore, the CO concentration effect on WGS reaction was investigated with a

H2O/CO = 2/1 under 850 and 900 oC. The total flow rate was 60 cm3/min and Ar was used as the

balanced gas. Figure 7-7 shows the H2 production, XCO, and H2/CO in the reactor side effluent as

a function of CO concentration. Compared to the results under 850 oC, a better performance was

obtained under 900 oC. For both temperatures, the XCO slightly decreases with increasing CO

concentration which can be explained by the percentage of the permeated H2 and H2 total

production. For example, the permeated H2/total H2 decreases from 14% to 8% under 900 oC

when CO concentration increases from 8.33% to 33.33%. Larger portion of produced H2

permeates through the membrane at lower CO concentration, moving the WGS reaction further

forward to the product side. As a result, the XCO increases. In contrast, the permeated H2, H2 in

the reactor side effluent, and H2 total production increases with CO concentration due to the

increase of the reactants amount.

100

7.3.7 Long Term Stability

Finally, the long term stability of the membrane was investigated. This experiment was

carried out under 900 oC with a feed gas composition of CO (10 cm3/min), H2O (20 cm3/min),

and Ar (10 cm3/min), respectively. As shown in Figure 7-8, the XCO is quite stable, only slightly

decreasing from 90% to 86% over 200 hours. Whereas the H2 total production slightly decreases

as the XCO decreases, the permeated H2 flux is essentially stable. The total carbon in gas phase is

the sum of detected CO, CO2 and CH4. It is essentially equal to the input amount (10 cm3/min),

indicating the carbon deposition is negligible. In addition, good integrity of the membrane

reactor remained after the experiment, showing high stability under the WGS reaction

conditions. This further confirms the stability of SrCe0.9Eu0.1O3-δ is improved by substitution 20

mol% Zr on Ce sites.

7.4 Conclusions

WGS reaction is constrained by thermodynamic equilibrium limitations. A tubular

SrCe0.7Zr0.2Eu0.1O3-δ H2 membrane reactor was fabricated. The XCO, H2 production, H2 yield and

the H2/CO in the reactor side effluent increased with increasing temperature and H2O/CO. A XCO

of 83.6% and 90.2% was achieved under 900 oC with H2O/CO = 1/1 and 2/1, respectively, 77%

and 44% increase compared to the thermodynamic calculation data. The respective improvement

in H2 production was 73% and 42%. In contrast to the XCO, the permeated H2, H2 in the reactor

side effluent and H2 total production increased with increasing flow rate and CO concentration.

The H2/CO in the reactor side effluent is variable through the SrCe0.7Zr0.2Eu0.1O3-δ H2 membrane.

This membrane is stable under the WGS reaction conditions.

101

A

B

Figure 7-1. Membrane morphology and experiment setup A) SEM image of the membrane and B) Schematic View of the membrane reactor.

102

0

0.1

0.2

0.3

0.4

0.5

650 700 750 800 850 900 950

Mol

e Fr

actio

nCO

2

H2

COH

2O

CH4

Temperature (oC)

CO2H

2

COH

2O

CH4

H2O/CO=1/1

A

650 700 750 800 850 900 9500

0.1

0.2

0.3

0.4

0.5

CO2

H2

H2O

CO

CH4

Temperature (oC)

H2O/CO=1.5/1

Mol

e Fr

actio

n

B

Figure 7-2. Gas compositions of the reactor side effluent as a function of temperature (dashed lines are from thermodynamic calculation) A) H2O/CO=1/1; B) H2O/CO=1.5/1 and C) H2O/CO=2/1.

103

0

0.1

0.2

0.3

0.4

0.5

650 700 750 800 850 900 950

Mol

e Fr

actio

n

CO2 H

2

CO

H2O

CH4

Temperature (oC)

CO2

H2

CO

H2O

CH4

H2O/CO=2/1

C

Figure 7-2.. Continued

104

Figure 7-3. XCO as a function of temperature.

40

50

60

70

80

90

100

650 700 750 800 850 900 950

X CO (%

)

Temperature (oC)

2/11.5/1

1/1

H2O/CO

Thermodynamic data (2/1)

Thermodynamic data (1/1)

105

0

2

4

6

8

10

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

650 700 750 800 850 900 950

H2 P

rodu

ctio

n (c

m3 /m

in)

H2 permeation

thermodynamic H2 production

H2 total production

H2 flow in reactor side effluent

Temperature (oC)

H2 Flux (cm

3/cm2 m

in)

A

0

2

4

6

8

10

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

650 700 750 800 850 900 950

H2 P

rodu

ctio

n (c

m3 /m

in)

Temperature (oC)

H2 Flux (cm

3/cm2 m

in)

H2 ptotal roduction

H2 flow in reactor side effluent

H2 permeation

B

Figure 7-4. H2 production as a function of temperature A) H2O/CO=1/1; B) H2O/CO=1.5/1 and C) H2O/CO=2/1.

106

0

2

4

6

8

10

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

650 700 750 800 850 900 950

H2 P

rodu

ctio

n (c

m3 /m

in)

H2 permeation

thermodynamic H2 production

H2 total production

H2 flow in reactor

side effluent

Temperature (oC)

H2 Flux (cm

3/cm2 m

in)

C


107

0

20

40

60

80

100

0

0.8

1.6

2.4

3.2

4

4.8

5.6

6.4

650 700 750 800 850 900 950

H2 y

ield

(%)

Temperature (oC)

total H2 yield

H2 in feed side effluent

H2/CO ratio in feed side effluent

permeated H2

thermodynamic H2/CO ratio

thermodynamic H2 yield H

2 /CO

4.0

A

0

20

40

60

80

100

650 700 750 800 850 900 9500

2

4

6

8

10

Temperature (oC)

H2 y

ield

(%) H

2 /CO

permeated H2


H2 total yield

H2/CO ratio in reactor side effluent

B

Figure 7-5. H2 yield and H2/CO in the reactor side effluent as a function of temperature A) H2O/CO=1/1; B) H2O/CO=1.5/1 and C) H2O/CO=2/1.

108

0

20

40

60

80

100

0

5

10

15

650 700 750 800 850 900 950

H2 y

ield

(%)

Temperature (oC)

H2 total yield


H2/CO ratio in reactor side effluent

permeated H2

thermodynamic H2/CO ratio

thermodynamic H2 yield H

2 /CO

C


109

0

5

10

15

20

0

2

4

6

8

1080

85

90

95

100

10 20 30 40 50 60 70

H2 P

rodu

ctio

n (c

m3 /m

in)

Total Flow Rate (cm3/min)

permeated H2

H2/CO in reactor side effluent


total H2

CO conversion

H2 /C

OX

CO (%

)

Figure 7-6. The XCO, H2 production and H2/CO in the reactor side effluent as a function of flow rates under 900 oC.

110

0

5

10

15

20

0

2

4

6

8

1080

85

90

95

100

5 10 15 20 25 30 35

H2 P

rodu

ctio

n (c

m3 /m

in)

XC

O (%)

CO Concentration (%)

H2 /C

O

permeated H2

H2/CO in reactor side effluent

total H2


CO conversion

Solid symbol--900 oCHollow symbol--850 oC

Figure 7-7. The XCO, H2 production and H2/CO in the reactor side effluent as a function of CO concentrations with H2O/CO=2/1.

111

Figure 7-8. The performance of the membrane reactor as a function of time under 900 oC.

112

CHAPTER 8 HIGH TEMPERATURE SrCe0.7Zr0.2Eu0.1O3-δ MEMBRANE REACTOR FOR H2

PRODUCTION AND SEPARATION USING THE STEAM REFORMING OF METHANE

8.1 Introduction

Steam reforming of methane (SRM) has been the most important chemical process in

producing H2 and syngas [115]. About 95% of the H2 produced in the U.S. is through the SRM

process [1]. This process includes two reversible reactions [116]:

242 3 HCOCHOH +↔+ 206=∆ oH kJ/mol (8-1)

222 HCOCOOH +↔+ 41−=∆ oH kJ/mol (8-2)

The reforming reaction (8-1) is an endothermic reaction, thermodynamically favored by high

temperature and low pressure. The WGS reaction (8-2) is an exothermic reaction, independent of

pressure and favored by low temperature. Heat generated by the WGS reaction is not sufficient

for the reforming reaction (8-1), as can be seen from the overall reaction:

2242 4 2 HCOCHOH +↔+ 165=∆ oH kJ/mol (8-3)

The conventional industrial process is carried out in furnaces at about 850 oC, a few atmospheres

and with Ni/Al2O3 as the catalyst. 80% of CH4 conversion is achieved under these operating

conditions [116].

The SRM reaction is limited by thermodynamic equilibrium. It should be carried out at

high temperature to achieve high CH4 conversion. Therefore, the SRM reaction is highly capital

intensive accounting for about 70% of the total investment and operating cost in methanol

production based on natural gas [117]. Therefore, the development of a membrane based

separation process gives the possibility of increasing the CH4 conversion. Continuous removal of

H2 through H2 permeable membranes moves the equilibrium toward the products, resulting in

higher CH4 conversion. In addition, SRM in a membrane reactor becomes a transfer-limited

113

reaction related with membrane porosity and diffusivity, rather than an equilibrium-limited

reaction [118].

Enhanced performance of SRM with a palladium membrane was first reported by Oertel

et al. [119]. Previous research has been focused on Palladium-based membranes [119-124]. The

Palladium-based membranes are deposited on porous glass or metal substrates. However, the

porous glass substrate is fragile and difficult in connecting to metallic application [125]. The

rough surface of the porous metal substrate usually results in pinholes on the membrane [126].

Catalytic ceramic membranes supply another option to incorporate H2 separation and SRM into

one unit. The critical features for successfully membrane reactors are high separation selectivity,

high permeability and stability. In this chapter, SCZE721 membranes coated on the inner side of

tubular NiO-SCZ82 supports are used to incorporate H2 separation and SRM into one unit. The

separation selectivity of H2 is very high since they are dense membranes. Thin film membranes

are fabricated to improve permeability.

8.2 Experimental

The dense SCZE721 membrane used in this experiment was ~33 µm thick with an active

area of 12 cm2 [6, 86] and coated on a NiO-SCZ82 support. Figure 8-1 (A) shows the cross

section image of the membrane after experiment. The Influence of temperature, CH4/H2O, CH4

flow rate, and CH4 concentration on the SRM are investigated.

The experimental setup (Figure 8-1 (B)) is the same as that in reference [86] . The outer

side of the membrane (feed side) was exposed to CH4 and steam. Steam was achieved by

gasifying the desired amount of water provided by a syringe pump. The inner side (sweep side)

of the membrane was flushed with He at 20 cm3/min, in co-current flow with the feed gases. The

flow rates of CH4, Ar and He were controlled by mass flow controllers. The reactants were

114

flowed into the quartz chamber and exposed to the Ni catalyst on the outside of the membrane.

The unreacted steam in the reactor side effluent was condensed by a cold trap filled with ice

prior to being analyzed by GC. The concentrations of the permeated H2 in the sweep gas (He)

were analyzed by a mass spectrometer.

The membranes were heat treated in H2 to reduce NiO to Ni in the support before

experiments [86].

The conversion of CH4, the selectivity of H2 and CO, and the ratio of H2/CO in the reactor

side effluent were defined:

%1004

44

4 ×−

= inCH

outCH

inCH

CH FFF

X (8-4)

%1002

×+

= outCO

outCO

outCO

CO FFFS (8-5)

%1002

22 ×

+= out

COout

CO

outCO

CO FFF

S (8-6)

%1004

22 ×= in

CH

outH

FF

yieldH (8-7)

outCO

outH FFCOH //

22 = (8-8)

where iX , iS , iniF and out

iF (i=CH4, CO2, H2, and CO) are the conversion, selectivity, input and

output flux of i, respectively.


8.3.1 Thermodynamic Calculation Results

There are seven possible species in the SRM system: CH4, H2O, H2, CO, CO2, C and O2.

Three equations can be obtained through mass balance of C, H and O atoms. Four more

equations are needed to solve the seven unknowns. These equations can be obtained through the

115

relation between thermodynamic equilibrium constant (K) and Gibbs free energy of reactions

between 298-2000 K [127]:

KRTG ln−=∆ (8-9)

The following four reactions were used in the calculation:

224 3HCOOHCH +=+

)( 15.78ln25.22204920 JTTTGo −−=∆ (8-1)

222 HCOOHCO +=+

)( 3236000 JTGo +−=∆ (8-2)

24 2HCCH +=

)( 35.65ln25.2269120 JTTTGo +−=∆ (8-10)

OHOH 222 2/1 =+

)( 85.55247500 JTGo +−=∆ (8-11)

The influence of CH4/H2O and CH4 and H2O concentrations on the SRM was calculated with a

total pressure of 1 atm.

Figure 8-2 shows the influence CH4/H2O on the calculated XCH4 as a function of

CH4/H2O and temperature. As expected, the XCH4 is strongly dependent on both temperature and

CH4/H2O. For both with and without inert diluent, the XCH4 increases with increasing

temperature since the SRM reaction (equation (8-1) and (8-3)) is endothermic reaction. The XCH4

increases with decreasing CH4/H2O as well. Similar findings were reported by Liu [128] and

Rakas et al [129].

Figure 8-3 shows the influence of CH4 and H2O concentrations on the calculated XCH4 as a

function of temperature with CH4/H2O=1/2. The XCH4 increases with decreasing the CH4 and

116

H2O concentrations. The volume expands in the SRM process. Therefore, lower CH4 and H2O

concentrations move the reaction forward to the product side and increase the XCH4.

The condition for carbon formation was calculated as a function of CH4/H2O and

temperature and is shown in Figure 8-4. The carbon formation is suppressed by lower CH4/H2O.

The CH4/H2O should be below 0.6 to avoid possible coking. CH4/H2O=1/2 is used in the

following sections except for the discussion of the CH4/H2O effect on the SRM.

8.3.2 Experimental Results

8.3.2.1 Influence of CH4/H2O on the SRM

The effect of CH4/H2O on the SRM was investigated using min/ 4 34

cmF inCH = ,

min/ 16 3cmF inAr = with a desired amount of steam. Figure 8-5 shows the XCH4 as a function of

temperature and CH4/H2O. The XCH4 increases with increasing temperature and decreasing

CH4/H2O, in agreement with the thermodynamic calculation results. A XCH4 of 95.0%, 92.0%

and 89.3% was achieved at 900 oC with CH4/H2O/Ar=1/3/4, 1/2/4 and 1/1/4, respectively.

However, the experimental XCH4 is lower than thermodynamic calculation results, especially at

low temperatures. This indicates the SRM process is rate limited by kinetic reaction rate which is

possibly caused by the short residence time and/or the inadequacy of Ni catalyst. In this

experimental setup, the Ni catalyst is embedded in the porous substrate. The reactants, CH4 and

H2O, flow between the quartz reactor and the H2 membrane. It is possible that part of the

reactants flows out the system without being exposed to the Ni, especially with the volume

expansion in the SRM process. This situation will be solved by building a Ni catalyst bed

between the quartz reactor and the H2 membrane, which will be discussed in more detailed in

chapter 10. The lower experimental XCH4 than thermodynamic equilibrium results is common in

literatures. 80% of XCH4, for instance, was reported at 850 oC [116].

117

Figure 8-6 shows the SCO, SCO2 and H2/CO in the reactor side effluent as a function of

temperature and CH4/H2O. The SCO increases with increasing temperature whereas the SCO2 and

H2/CO decrease with increasing temperature which is ascribed to the higher XCH4 and exothermic

nature of the WGS reaction (equation (8-2)). The SCO decreases with decreasing CH4/H2O since

lower CH4/H2O moves the WGS reaction forward and produces more CO2 and H2.

Consequently, the SCO2 and H2/CO increase with decreasing CH4/H2O. If the effluent is used to

synthesize liquid fuels through the Fischer-Tropsch process, the H2/CO is too high since the ideal

value is ~2. However, it can be adjusted by combing the SRM and carbon dioxide reforming of

methane, which will be discussed in chapter 9.

The H2 production is shown in Figure 8-7 as a function of temperature and CH4/H2O. The

H2 permeation decreases with decreasing CH4/H2O due to the higher PO2 in lower CH4/H2O. It

increases with increasing temperature due to the higher ambipolar conductivity of the H2

membrane at higher temperatures. A maximum H2 permeation of 0.18 cm3/cm2 min was

achieved at 900 oC and CH4/H2O/Ar=1/1/4. The H2 flow in the reactor side effluent is not a

significant function of temperature and increases with decreasing CH4/H2O. The total H2 is the

sum of the H2 permeation and the H2 flow at the reactor side effluent. It increases with increasing

temperature and decreasing CH4/H2O, which is consistent with the XCH4 in Figure 8-5. A

maximum H2 production of 11.8 cm3/ min was achieved at 900 oC with CH4/H2O/Ar=1/3/4.

8.3.2.2 Influence of CH4 concentration on the SRM

The effect of CH4 concentration on the SRM was investigated with a total flow rate of 60

cm3/ min and CH4/H2O=1/2. Ar was used as the balanced gas. Figure 8-8 shows the XCH4, SCO,

H2/CO and H2 production as a function of CH4%. The XCH4 slightly decreases with increasing

CH4% in agreement with the thermodynamic calculation results (Figure 8-3).This can be

explained by the volume expansion of the SRM. Higher CH4% results in lower XCH4 and higher

118

PH2O. As a result, the SCO decreases with increasing CH4%. In contrast with the XCH4 and SCO, the

permeated H2, H2 flow in the reactor side effluent and total H2 production increase with CH4%

due to the increase of the reactants amount. The H2/CO increases with increasing CH4% as well

due to the higher H2 production and lower SCO. In addition, compared to the results under 850

oC, the XCH4, SCO and H2 production are higher at 900 oC whereas the H2/CO is lower. This is

consistent with the discussion in section 8.3.2.1.

8.3.2.3 Influence of total flow rate on the SRM

The total flow rate effect on the SRM was investigated from 700 to 900 oC with

CH4/H2O=1/2. As shown in Figure 8-9, the XCH4 decreases with increasing total flow rate due to

the shorter residence time at higher total flow rate. Lower XCH4 results in higher PH2O which

favors the WGS reaction and produces more CO2. Consequently, the SCO decreases with

increasing total flow rate as well. In contrast, the permeated H2, H2 flow in the reactor side

effluent and total H2 production increase with increasing total flow rate due to the increase of the

reactants amount. The H2/CO increases with increasing total flow rate as well due to the higher

H2 production and lower SCO. In addition, the XCH4, SCO and H2 production increase and the

H2/CO decreases with increasing temperature in agreement with the discussion in section 8.3.2.1.

8.3.2.4 Influence of the H2 membrane reactor on the SRM

To investigate the influence of the H2 membrane reactor on the SRM, experiment was

carried out with min/ 20 34

cmF inCH = and min/ 40 3

2cmF in

OH = under three reactor configurations:

(1) blank quartz reactor, (2) with Ni catalyst and (3) with Ni catalyst and in situ H2 removal. For

the blank quartz reactor, CH4 and H2O were fed into the empty quartz reactor. For the

configuration (2), a H2 membrane was installed in the quartz reactor but the permeated side

outlet was blocked so that no produced H2 was removed. Therefore, the membrane only

119

functioned as a catalyst in this situation. For the configuration (3), the permeated side of the H2

membrane was open and connected to a mass spectrometer, so that the permeated H2

concentration could be analyzed. These three configurations are similar with those configurations

discussed in the section 4.3.2.

As shown in Figure 8-10, the XCH4 in the blank quartz reactor is very low and increases

with increasing temperature which is due to the higher kinetic reaction rate at higher

temperature. The XCH4 of the configuration (3) is much higher than that of the configuration (2)

even though both XCH4 are lower than the thermodynamically calculated XCH4. The maximum

XCH4 of the configuration (3) is 89% at 900 oC, 15% increase compared to the 77.5% of the

configuration (2) under the same conditions. Similarly, the total H2 production is much higher in

the configuration (3). The maximum H2 production of configuration (3) and (2) are 57.6 and 50.3

cm3/min at 900 oC, 15% increase for the configuration (3) with in situ H2 removal. In contrast,

the SCO and H2/CO are higher in the configuration (2). This demonstrates the H2 membrane can

improve the SRM performance by increasing the XCH4 and H2 production.

8.3.2.5 Long term stability

The long term stability of the H2 membrane was investigated under 850 oC with

min/ 8 34

cmF inCH = and min/ 16 3

2cmF in

OH = . As shown in Figure 8-11, the XCH4, the H2

permeation and total H2 production are all quite stable over 70 hours. This demonstrates the high

stability of the H2 membrane under the SRM conditions.

8.4 Conclusions

The XCH4 under thermodynamic equilibrium increases with increasing temperature and

decreasing CH4/H2O and CH4%. The experimental XCH4 is lower than the thermodynamic data

and limited by kinetic reaction rate. However, the H2 membrane can still enhance the SRM

120

performance by increasing 15% of both the XCH4 and total H2 production at 900 oC compared to

that with only Ni catalyst. The experimental XCH4 increases with increasing temperature and

decreasing CH4/H2O, CH4% and total flow rate. In contrast with the H2/CO, the SCO increases

with increasing temperature and decreases with decreasing CH4/H2O and increasing CH4% and

total flow rate. In contrast with the total H2 production, the H2 permeation decreases with

decreasing CH4/H2O. The H2 permeation and production increase with increasing temperature,

CH4% and total flow rate. The SrCe0.7Zr0.2Eu0.1O3-δ hydrogen membrane is stable under the SRM

conditions.

121

A

B Figure 8-1. Membrane morphology and experimental setup A) Cross section of the membrane

after experiment and B) Experimental setup.

122

0

20

40

60

80

100

300 400 500 600 700 800 900 1000 1100

X CH

4 (%)

Temperature (oC)

CH4/H

2O

1/3

1/1

1/2

A

0

20

40

60

80

100

300 400 500 600 700 800 900 1000 1100

X CH

4 (%)

Temperature (oC)

CH4/H

2O/Ar

1/3/4

1/1/4

1/2/41/2/4

B

Figure 8-2. Influence of CH4/H2O on XCH4 under thermodynamic equilibrium A) without a diluent and B) with a diluent (Ar).

123

0

20

40

60

80

100

300 400 500 600 700 800 900 1000 1100

X CH

4 (%)

Temperature (oC)

CH4/H

2O/Ar

1/2/91/2/41/2/31/2/11/2/0

CH4 & H

2O

Concentration increasing

Figure 8-3. Influence of CH4/H2O concentrations on XCH4 under thermodynamic equilibrium

with CH4/H2O=1/2 and Ar as the diluent.

124

Figure 8-4. Thermodynamic calculation of carbon formation as a function of temperature and CH4/H2O [130].

125

70

75

80

85

90

95

100

650 700 750 800 850 900 950

X CH

4 (%)

Temperature (oC)

1/1/4

1/2/41/2/4

1/3/41/3/4

1/1/4

CH4/H

2O/Ar

Experimental

Thermodynamiccalculation

Figure 8-5. Influence of CH4/H2O on XCH4.

126

10

20

30

40

50

60

70

80

90

2.5

3

3.5

4

4.5

5

5.5

6

650 700 750 800 850 900 950

S CO &

SC

O2 (%

)H

2 /CO

Temperature (oC)

1/1/4

1/1/4

1/2/4

1/2/4

1/2/4

1/3/4

1/3/41/3/4

1/1/4

CH4/H

2O/Ar

6.0

5.0

4.0

3.0

Figure 8-6. Influence of CH4/H2O on SCO, SCO2 and H2/CO in reactor side effluent (solid symbol—SCO2, hollow symbol—SCO and dashed line—H2/CO).

127

650 700 750 800 850 900 950

H2 P

rodu

ctio

n (c

m3 /m

in)

1/1/4

1/1/4

1/1/4

1/2/4

1/2/4

1/2/41/3/4

1/3/4

1/3/4

CH4/H

2O/Ar

Temperature (oC)

12

11

10

9

8

2

1

0.20

0.160.18

0.140.120.100.080.06

H2 Perm

eation (cm3/cm

2 min)

Figure 8-7. Influence of CH4/H2O on H2 production (solid symbol—total H2 production, hollow symbol—H2 production in reactor side effluent and dashed line—H2 permeation).

128

65

70

75

80

85

90

95

3.2

3.4

3.6

3.8

4

4.2

4.4

5 10 15 20 25 30 35

X CH

4 & S

CO (%

)

H2 /C

O

CH4 Concentration (%)

XCH4

SCO

4.0

H2/CO

Solid symbol-900 oCHollow symbol-850 oC

A

0

10

20

30

40

50

60

5 10 15 20 25 30 350

0.8

1.6

2.4

3.2

4

4.8

H2 P

rodu

ctio

n (c

m3 /m

in)

CH4 Concentration (%)

Solid symbol-900 oCHollow symbol-850 oC

H2 Permeation

Total H2


H2 Perm

eation (cm3/cm

2 min)

4.0

B

Figure 8-8. Influence of CH4 concentration on SRM A) XCH4, SCO & H2/CO vs CH4% and B) H2 production vs CH4%.

129

70

75

80

85

90

95

0

10

20

30

40

50

60

20 30 40 50 60 70

X CH

4 & S

CO (%

)

H2 Production (cm

3/cm2 m

in)


XCH4

SCO H

2 permeation


Total H2

CH4/H

2O=1/2 & 900 oC

A

65

70

75

80

85

90

0

10

20

30

40

50

60

20 30 40 50 60 70

X CH

4 & S

CO (%

)

H2 Production (cm

3/cm2 m

in)


XCH4

SCO

H2 permeation


Total H2

CH4/H

2O=1/2 & 850 oC

B

Figure 8-9. Influence of total flow rate on SRM A) 900 oC, B) 850 oC, C) 800 oC, D) 750 oC, E) 700 oC and F) H2/CO in reactor side effluent.

130

60

65

70

75

80

85

90

0

10

20

30

40

50

60

20 30 40 50 60 70

X CH

4 & S

CO (%

)

H2 Production (cm

3/cm2 m

in)


XCH4

SCO

H2 permeation


Total H2

CH4/H

2O=1/2 & 800 oC

C

60

65

70

75

80

85

0

10

20

30

40

50

60

20 30 40 50 60 70

X CH

4 & S

CO (%

)

H2 Production (cm

3/cm2 m

in)


XCH4

SCO

H2 permeation


Total H2

CH4/H

2O=1/2 & 750 oC

D

Figure 8-9. Continued

131

55

60

65

70

75

80

0

10

20

30

40

50

60

20 30 40 50 60 70

X CH

4 & S

CO (%

)

H2 Production (cm

3/cm2 m

in)


XCH4

SCO

H2 permeation


Total H2

CH4/H

2O=1/2 & 700 oC

E

20 30 40 50 60 70

H2/C

O


700 oC

750 oC

800 oC

850 oC

900 oC

CH4/H

2O=1/2

6.0

5.5

5.0

4.0

3.0

4.5

3.5

F

Figure 8-9. Continued

132

0

20

40

60

80

100

4

4.5

5

5.5

650 700 750 800 850 900 950

X CH

4 & S

CO (%

), H

2 Pro

duct

ion

(cm

3 /min

)H

2 /CO

XCH4

XCH4

Total H2 production

Total H2 production

SCO

SCO

XCH4

H2/CO

H2/CO

4.0

Temperature (oC)

Thermodynamic XCH4

Figure 8-10. Influence of reactor configurations on SRM (solid symbol—with catalyst and H2 removal, hollow symbol—with catalyst and dashed line—blank quartz reactor).

133

Figure 8-11. The performance of the membrane reactor as a function of time under 850 oC.

134

CHAPTER 9 HIGH TEMPERATURE SrCe0.7Zr0.2Eu0.1O3-δ PROTON CONDUCTING MEMBRANE

REACTOR FOR CARBON DIOXIDE REFORMING OF METHANE

9.1 Introduction

Fossil fuels are likely to play a major role in H2 energy in the near to medium-term future

with their inherent advantages, such as availability, relatively low cost and the existing

infrastructure for delivery and distribution [1]. Currently, the majority of H2 is produced through

SRM process with significant CO2 emission. The global warming potential (GWP) of H2

production via the SRM process is estimated to be 13.7 kg CO2 (equiv.) per kg of net H2

produced (CO2 accounts for 77.6% of the system’s GWP) [131]. 0.3-0.4 million cubic meters of

CO2 will be produced when one million cubic meters of H2 is produced through a typical SRM

H2 plant. The amount of CO2 emission would be double if H2 is to be produced by coal

gasification [1]. Therefore, CO2 sequestration has drawn lots of interest. The capture and

disposal of CO2 costs about 25-30% of the total cost of H2 production by the SRM process [132].

The net cost of CO2 disposal, however, could be significantly reduced if CO2 sequestration is

accompanied by enhanced oil recover [133].

9.1.1 Carbon Dioxide Reforming of Methane (CDRM)

There is a growing interest in catalytic reforming of methane with carbon dioxide

(equation (9-1)) because of the great benefit to both the economy and the environment. The

reforming reaction is:

224 2 2 HCOCOCH +↔+ 247=∆ oH kJ/mol (9-1)

This conversion consumes two undesirable greenhouse gases, CO2 and CH4, to generate

syngas. The produced syngas has a low H2/CO ratio, i.e., 1:1 or less. A low ratio is preferred for

synthesis of oxygenated compounds and long-chain hydrocarbons [134-136]. It also introduces

135

the possibility of combining the steam reforming, partial oxidation, and dry reforming reactions

to get the desired H2/CO [137] for different applications. Several technologies have been applied

to CO2 reforming of CH4 including catalysis conversion [138-143], plasma conversion [144-146]

and combination of catalyst and plasma [147, 148]. Nickel is a typical component in catalytic

reforming catalysts due to its wide availability, low cost and high catalytic activity [149-152].

9.1.2 Membrane Reactors for the CDRM

While plenty of research has been focused on the catalysts, there are only a few works on

the membrane reactor effect on the CO2 reforming of CH4 [153-155]. With in situ removal of H2,

membrane reactors can increase conversion of CH4 and CO2 compared to traditional reactors, or

the reforming process can be operated at lower temperature. Most of those membrane reactors in

the literatures above are Pd based membranes and they are expensive. Here, we investigate the

CDRM using ceramic membranes.

SrCeO3-δ and BaCeO3-δ based perovskite oxides with multivalent cation dopants are

promising for H2 membranes and have been reported by several groups [5, 6, 26, 27, 47, 73, 86].

We previously reported the H2 permeation of SrCe0.9Eu0.1O3-δ and SCZE721 [6, 86, 112]. Zr was

used to improve the stability of SrCe0.9Eu0.1O3-δ [7, 108]. We also carried out the water gas shift

(WGS) reaction using a SCZE721 H2 membrane reactor and found CO conversion and H2

production were significantly increased compared to the thermodynamic calculation [156]. The

SCZE721 membrane was coated on the inner side of a tubular NiO-SCZ82 support. In this

chapter, we investigate the SDRM through H2 permeable SCZE721 membrane reactors in terms

of CH4 and CO2 conversion, CO and H2 selectivity, and H2/CO in syngas product. Whereas the

H2/CO in syngas product through the CDRM is 1/1 or below, it is much higher through the steam

136

reforming of methane (SRM). The ideal H2/CO is ~2 for Fischer-Tropsch process to produce

liquid fuels. Therefore, we adjust the H2/CO by combining the CDRM and SRM processes.

9.1.3 Reaction Mechanism and Kinetics

The reaction of CO2 and CH4 is expected to proceed by three steps: (1) dehydrogenation of

methane to form surface carbon and H2, (2) dissociative of adsorption of CO2 and H2, and (3)

reduction of CO2 to CO. Possible reaction mechanism has been proposed [135]:

)(4)(4 aHaCCH += (9-2)

)()()(2 aOaCOgCO += (9-3)

)()()( aCOaOaC =+ (9-4)

)()( gCOaCO = (9-5)

)()(2 2 gHaH = (9-6)

A simplified reaction sequence for the CO2 reforming may involve two irreversible steps,

namely, the activation of methane followed by the surface reaction with adsorbed oxygen atoms:

**34 * HCHCH +=+ (9-7)

**2

*3 * HCHCH +=+ (9-8)

***2 * HCHCH +=+ (9-9)

*** * HCCH +=+ (9-10)

*2)2/( 2** ++=+ HxCOOCH x (9-11)

COOCO +=+ *2 * (9-12)

It has been proved from experiments that CH4 promotes the dissociation of CO2 on

catalysts. The promotion is attributed to the effect of H2 in the decomposition of CH4. It has been

demonstrated that a small amount of H2 can significantly facilitate this process [157]. Assuming

137

the above effect, the reaction mechanism of the CO2 reforming of CH4 on Rh and Pd was

proposed as follows:

HCHCH += 34 (9-13)

OHCOHCO +=+2 (9-14)

OHCHOCH +=+ 34 (9-15)

HCHCH += 23 (9-16)

HCHCH +=2 (9-17)

xHCCH x += (9-18)

xHCOOCH x +=+ (9-19)

xHCOCOCH x +=+ 22 (9-20)

22 HH = (9-21)

OHOH 22 = (9-22)

The kinetics for the CO2 reforming of CH4 depends on the catalyst and the mechanism is

probably changing with temperature. Previous investigations of the CO2 reforming of CH4

mainly dealt with the screening test of catalysts. Little research has focused on the reaction

mechanism and the kinetics. Thus, no general expression has been derived so far.

Sakai et al. [158] considered that the half-order and zero-order dependence of the reaction

rates on the partial pressure of CO2 or CH4 suggested that CO2 participates in the reaction via the

dissociative adsorption mechanism, being described as equation (9-3), whereas CH4 participates

in the reaction as strongly adsorbed species dehydrogenated to CHx and (4-x)H. A rate equation

was obtained by Richardson and Paripatyadar [159] using linear regression analysis:

2)1/(44224242 CHCHCOCOCHCOCHCOr PKPKPPKKKR ++= (9-23)

138

A regression coefficient of 0.988 was from calculated and measured rates. Experimental results

and the model fit well with each other.

9.2 Experimental

The dense SCZE721 membrane used in this experiment was ~33 µm thick (Figure 9-1 (A))

with an active area of 12 cm2 [6, 86] and coated on a NiO-SCZ82 support. The influence of

temperature, feed flow rate, and CH4/CO2 and CH4/CO2/H2O were evaluated in terms of CH4 and

CO2 conversion, H2 production and H2/CO.

The experimental setup (Figure 9-1 (B)) is the same as that in reference [86] . The outer

side of the membrane (feed side) was exposed to CH4 and CO2 or/and steam. Steam was

achieved by gasifying the desired amount of water provided by a syringe pump. The inner side

(sweep side) of the membrane was flushed with He at 20 cm3/min, in co-current flow with the

feed gases. The flow rates of CH4, CO2 and He were controlled by mass flow controllers. The

reactants were flowed into the quartz chamber and exposed to the Ni catalyst on the outside of

the membrane. The produced and/or unreacted steam in the reactor side effluent was condensed

by a cold trap filled with ice prior to being analyzed by GC. The concentrations of the permeated

H2 in the sweep gas (He) were analyzed by a mass spectrometer.

The membranes were heat treated in H2 to reduce NiO to Ni in the support before

experiments [86].

The conversion of CH4 and CO2, the selectivity of H2 and CO, and the ratio of H2/CO in

the reactor side effluent were defined:

%1004

44

4 ×−

= inCH

outCH

inCH

CH FFF

X (9-24)

%1002

22

2×

−= in

CO

outCO

inCO

CO FFF

X (9-25)

139

%100)(2

44

2

2×

−×= out

CHin

CH

outH

H FFF

S (9-26)

%100)()(

2244

×−+−

= outCO

inCO

outCH

inCH

outCO

CO FFFFFS (9-27)

outCO

outH FFCOH //

22 = (9-28)

where iX , iS , iniF and out

iF (i=CH4, CO2, H2, and CO) are the conversion, selectivity, input and

output flux of i, respectively.


9.3.1 CH4/CO2 Effect on Conversion, H2/CO and H2 Production

The effect of CH4/CO2 on the CO2 reforming of CH4 was investigated using

min/ 10 34

cmF inCH = with a desired amount of CO2. Figure 9-2 shows the XCH4 and XCO2 as a

function of temperature and CH4/CO2. Both XCH4 and XCO2 increase with increasing temperature

since the CO2 reforming of CH4 is endothermic. The XCH4 increases with decreasing CH4/CO2 as

well. The XCH4 and XCO2 significantly depends on catalysts and their supports [135]. Various

results have been reported [138, 139, 151, 160-162] and our results are in the range in the

literatures. At 900 oC, the XCH4 is 87%, 89% and 93% for CH4/CO2 = 1/1, 1/1.5 and 1/2,

respectively.

The XCH4 and XCO2 should be equal to each other with CH4/CO2 = 1/1 without any side

reactions. However, the measured XCH4 is higher. This is ascribed to carbon deposition through

CH4 decomposition and/or the Boudouard reaction:

24 2 HCCH +↔ 75=∆ oH kJ/mol (9-29)

CCOCO +↔ 2 2 172−=∆ oH kJ/mol (9-30)

140

The carbon deposition is confirmed by the less detected total gas phase carbon amount,

sum of CH4, CO and CO2 in the output. It is 2-5% lower than the input amount for CH4/CO2 =

1/1. In contrast, the total gas phase carbon amount essentially remains constant for CH4/CO2 =

1/1.5 and 1/2 indicating that carbon deposition is negligible under these conditions. This is in

agreement with Wang’s results [135]. Carbon deposition is thermodynamically possible for

CH4/CO2 = 1/1 at temperature up to 870 oC at 1 atm. Lower CH4/CO2 can suppress carbon

deposition. The lowest temperature for CH4/CO2 = 1/1.5 and 1/2 are 760 oC and 710 oC,

respectively.

Controversial results of the XCH4 and XCO2 have been reported [151, 160, 163]. Higher

XCO2 is ascribed to the reverse water gas shift (RWGS) reaction:

OHCOCOH 222 +↔+ 6.40=∆ oH kJ/mol (9-31)

The difference may be due to different catalysts and higher operating temperature in this

work. Reaction (9-7) and (9-8) are highly dependent on catalysts. According to Sacco et al [164],

the primary source of surface carbon on Ni catalyst is CH4 indicating that most of the carbon

deposition is through Reaction (9-7). In addition, according to Gibbs free energy, RWGS and the

Boudouard reactions could not occur spontaneously over 820 oC [135] while CH4 decomposition

is favorable at high temperature. In addition, a part of the produced H2 permeates through the

membrane reactor, lowering the H2 partial pressure and further limiting the RWGS reaction. As a

result, the XCH4 is higher than the XCO2 due to carbon deposition with CH4/CO2 = 1/1.

A low CH4/CO2 not only suppresses the carbon deposition but also enhances the RWGS

reaction (Reaction 9-31), resulting in higher XCO2. For example, the XCO2 is 52.4% with

CH4/CO2 = 1/2 at 900 oC, higher than half of the XCH4 (93.2%).

141

Figure 9-3 shows the SH2 and SCO as a function of temperature and CH4/CO2. It is noted

that H2 total production is used in calculating the H2 selectivity. At any given temperature and

CH4/CO2, the SCO is higher than the SH2. This is due to the RWGS reaction consuming some

produced H2. The SH2 decreases with increasing temperature in contrast to the SCO since the

RWGS reaction is endothermic. Lower SH2 and higher SCO are obtained at lower CH4/CO2.

Lower CH4/CO2 means more amount of oxygen atoms are in excess, which eventually are in

forms of H2O, CO and CO2 through reaction (9-9). Therefore, SH2 decreases as CH4/CO2

decreases in contrast to the SCO.

Figure 9-4 shows the H2 production. H2 total production is the sum of the H2 permeation

and the H2 in the reactor side effluent. All of them increase with increasing temperature due to

the higher ambipolar conductivity of the membrane and higher XCH4 at higher temperatures. A

maximum H2 permeation of 2.2 cm3/min (~0.2 cm3/cm2 min) was achieved at 900 oC with

CH4/CO2 = 1/1. The H2 permeation decreases with decreasing the CH4/CO2. This is due to the

higher PO2 in the system, which causes lower electronic conductivity of the membrane material

and results in lower H2 permeation flux. A maximum H2 total production of 17.1 cm3/min was

achieved at 900 oC with CH4/CO2 = 1/2.

Figure 9-5 shows the H2/CO in the reactor side effluent. The H2/CO is less than 1 due to

the RWGS reaction. The H2/CO decreases with increasing temperature due to the endothermic

RWGS reaction and higher H2 permeation at elevated temperatures. The decrease in the H2/CO

with lower CH4/CO2 is ascribed to the RWGS reaction as well. When temperature is increased

from 700 oC to 900 oC, the H2/CO decreases from 1.00 to 0.91, from 0.95 to 0.83 and from 0.92

to 0.77 for CH4/CO2 = 1/1, 1/1.5 and 1/2, respectively.

142

9.3.2 Flow Rate Effect on Conversion, H2/CO and H2 Production

To prevent carbon deposition, operating temperatures must be high enough. However,

nickel carbide may form on the surface of Ni-based catalysts at high temperatures [135].

Accordingly, an upper-temperature limit is needed to prevent such formation. For example, the

optimum temperature is between 870 oC and 1040 oC with CH4/CO2 = 1/1 at 1 atm. As

previously described, lower CH4/CO2 can suppress carbon deposition and enhance the RWGS

reaction as well. Therefore, an intermediate CH4/CO2 = 1/1.5 was chosen to investigate the flow

rate effect on the XCH4 and XCO2, H2 production and H2/CO at 850 oC and 900 oC. The total flow

rates are 25, 37.5 and 50 cm3/min, respectively.

As shown in Figure 9-6, the XCH4, XCO2 and SCO are higher at 900 oC than 850 oC. The SH2

is lower at 900 oC. This agrees with the results in section 9.3.1. In addition, the XCH4 and XCO2

slightly decrease with total flow rate due to shorter residence time. Whereas the SCO increases

with increasing total flow rate, the SH2 decreases in a similar trend with Raybold’s results [154].

The permeated H2 percentage of the H2 total production decreases with increasing total flow rate.

It decreases from 13.2% to 9.3% and from 10.4% to 8.2% at 900 oC and 850 oC, respectively,

when the total flow rate is increased from 20 to 50 cm3/min (Figure 9-8). Higher H2 partial

pressure in the reactor side effluent favors the RWGS reaction. As a result, the H2 selectivity

decreases with increasing total flow rate.

The H2/CO in the reactor side effluent (Figure 9-7) increases with increasing total flow rate

due to the lower percentage of the H2 permeation in the total H2 production. It is also higher at

850 oC than 900 oC which is ascribed to the WGS reaction.

Figure 9-8 shows the H2 production as a function of total flow rate. The H2 production

increases with total flow rates since more mass of reactants react even though the XCH4 and XCO2

slightly decrease. The H2 permeation fraction in H2 total production decreases with increasing

143

total flow rate. A maximum H2 permeation of 3.1 cm3/min (~0.26 cm3/cm2 min) and a maximum

H2 total production 32.8 cm3/min were achieved at 900 oC with a total flow rate of 50 cm3/min.

9.3.3 CH4/CO2/H2O Effect on XCH4, XCO2, H2/CO and H2 Production

The theoretical H2/CO in syngas through reaction (9-1) is 1/1 and it is less than 1/1 with

this H2 membrane reactor due to the in situ removal of H2 and the RWGS reaction. The ideal

H2/CO is ~2 to produce liquid fuels through the Fischer-Tropsch process. In addition, the H2/CO

is high in syngas produced through SRM:

242 3 HCOCHOH +↔+ 206=∆ oH kJ/mol (9-32)

Therefore, H2O was added to CO2 and CH4, combining the CDRM and SRM, to increase

the H2/CO. CH4/CO2/H2O = 2/1/1 and 2/1/1.5 were investigated in terms of H2/CO and XCH4 and

XCO2 with min/ 20 34

cmF inCH = , min/ 10 3

2cmF in

CO = , and desired amount of steam.

Figure 9-9 shows the XCH4 and XCO2 as a function of temperature. Both the XCH4 and

XCO2 increase with temperature. The XCH4 increases with increasing H2O concentration as well.

A higher H2O concentration favors the SRM and enhances the XCH4. A higher H2O concentration

also means lower concentration of CH4 and CO2 since the CH4/CO2 is fixed at 2/1 and favors the

WGS reaction as well. As a result, the XCO2 decreases. The XCH4 and XCO2 are 85% and 70% at

900 oC with CH4/CO2/H2O = 2/1/1, which means more than half of the converted CH4 is through

the SRM (reaction (9-32)). The SRM is even more dominant at low temperature since it is less

endothermic and the WGS reaction is exothermic.

Figure 9-10 shows the H2 production and H2/CO as a function of temperature. The H2

production increases with temperature. The H2 flux in reactor side effluent and H2 total increase

with increasing H2O concentration as well due to the higher XCH4. In contrast, the H2 permeation

decreases with increasing H2O since higher H2O concentration means higher PO2. As a result, the

144

H2 permeation decreases. A maximum H2 permeation of 4.7 cm3/min (~0.39 cm3/cm2 min) and a

maximum H2 total production of 48.0 cm3/min were achieved at 900 oC with CH4/CO2/H2O =

2/1/1 and 2/1/1.5, respectively. The H2/CO in the reactor side effluent decreases with

temperature in similar trend as that in Figure 9-5 and is attributed to the RWGS reaction. The

H2/CO between 700 oC to 900 oC is between 1.9-1.7 and 2.5-2.0 for CH4/CO2H2O = 2/1/1 and

2/1/1.5, respectively. This demonstrates the H2/CO is variable in CH4/CO2/H2O system.

9.4 Conclusions

Carbon dioxide reforming of methane was investigated using a tubular

SrCe0.7Zr0.2Eu0.1O3-δ H2 membrane reactor. The XCH4 and XCO2, CO selectivity and H2

production increase with increasing temperature. 93.2% of XCH4 was achieved with CH4/CO2 =

1/2 at 900 oC. A maximum H2 permeation of 2.2 cm3/min (~0.2 cm3/cm2 min) and a maximum

H2 total production of 17.1 cm3/min were achieved at 900 oC with the CH4/CO2 = 1/1 and 1/2,

respectively. In contrast, the H2 selectivity and H2/CO decrease with increasing temperature.

The XCH4, XCO2, SH2, and H2/CO decrease with total flow rate. In contrast, the SCO and H2

production increases with increasing total flow rate. A maximum H2 permeation of 3.1 cm3/min

(~0.26 cm3/cm2 min) and a maximum H2 production 32.8 cm3/min were achieved at 900 oC with

a total flow rate of 50 cm3/min.

The H2/CO in the reactor side effluent for carbon dioxide reforming of methane is less than

1. However, it can be increased by adding steam to the system. The H2/CO between 700 oC to

900 oC is between 1.9-1.7 and 2.5-2.0 for CH4/CO2/H2O = 2/1/1 and 2/1/1.5, respectively.

145

A

B

Figure 9-1. Membrane morphology and experimental setup A) Cross section of the membrane after experiment and B) Experimental setup.

146

40

50

60

70

80

90

100

650 700 750 800 850 900 950

X CH

4 & X

CO

2 (%)

CH4/CO

2

Solid line-XCH4

Dashed line-XCO2

1/2

1/2

1/1.5

1/1.5

1/1

1/1

Temperature (oC)

Figure 9-2. XCH4 and XCO2 as a function of temperature and CH4/CO2.

147

92

93

94

95

96

97

98

99

100

650 700 750 800 850 900 950

S H2 &

SC

O(%

)

Temperature (oC)

CH4/CO

2

Solid line-SH2

Dashed line-SCO

1/2

1/1.5

1/1

1/2

1/1

1/1.5

Figure 9-3. SH2 and SCO as a function of temperature and CH4/CO2.

148

650 700 750 800 850 900 950

H2 p

rodu

ctio

n (c

m3 /m

in)

Temperature (oC)

CH4/CO

2

H2 permeation

solid symbol- total H2 production

hollow symbol- H2 in reactor side effluent

1/2

1/2

1/2

1/1.5

1/1.5

1/1.5

1/1

1/1

1/1

18

14151617

111213

103

2

1

Figure 9-4. H2 production as a function of temperature and CH4/CO2.

149

650 700 750 800 850 900 950

H2/C

O

Temperature (oC)

1/1

1/2

1/1.5

CH4/CO

2

1.00

0.95

0.90

0.85

0.80

0.75

Figure 9-5. H2/CO in the reactor side effluent as a function of temperature and CH4/CO2.

150

50

55

60

65

70

75

80

85

90

93

94

95

96

97

98

99

100

20 25 30 35 40 45 50 55

X CH

4 & X

CO

2(%) S

H2 &

SC

O (%

)

Total flow rate (cm3/min)

XCH4

900oC

XCH4

850oC

XCO2

900oC

XCO2

850oC

SH2

850 oC

SH2

900 oC

SCO

900 oC SCO

850 oC

CH4/CO

2=1/1.5

Figure 9-6. XCH4, XCO2, SH2 and SCO as a function of total flow rate.

151

0.82

0.83

0.84

0.85

0.86

0.87

0.88

20 25 30 35 40 45 50 55

H2/C

O


850 oC

900 oC

CH4/CO

2=1/1.5

Figure 9-7. H2/CO in reactor side effluent as a function of total flow rate.

152

8

9

10

11

12

13

14

20 25 30 35 40 45 50 55

H2 p

rodu

ctio

n (c

m3 /m

in)

H2p percentage of total H

2 production (%)

H2 permeation

solid symbol--900oChollow symbol--850oC


CH4/CO

2=1/1.5

H2 total production


H2p

fraction

35.0

30.0

25.0

20.0

15.0

10.0

3.5

3.0

2.5

2.0

1.5

Figure 9-8. H2 production as a function of total flow rate.

153

10

20

30

40

50

60

70

80

90

650 700 750 800 850 900 950

X CH

4 & X

CO

2 (%)

Temperature (oC)

XCH4

XCO2

2/1/1.52/1/1

2/1/1

2/1/1.5

CH4/CO

2/H

2O

Figure 9-9. XCH4 and XCO2 as a function of temperature.

154

2

3

4

5

6

7

830

35

40

45

50

650 700 750 800 850 900 950

H2 p

rodu

ctio

n (c

m3 /m

in)

Temperature (oC)

H2 /C

O

H2 permeation

H2 total production


H2/CO

solid symbol--2/1/1hollow symbol--2/1/1.5

CH4/CO

2/H

2O

2.6

1.5

2.0

Figure 9-10. H2 production and H2/CO as a function of temperature.

155

CHAPTER 10 CONCLUSIONS AND FUTURE WORKS

10.1 Conclusions

In this dissertation, H2 production and separation is demonstrated using supported

SCZE721 thin film membranes through WGS, SRM and CDRM processes. My conclusions are

as followings:

Supported SrCe0.9Eu0.1O3-δ and SCZE721 thin film membranes are successfully fabricated

by tape casting followed by a rolling step. NiO-SrCeO3-δ and NiO-SrCe0.8Zr0.2O3-δ are used as

the substrates to maintain mechanical integrity.

CO2 is the main reason for the decomposition of SrCe0.9Eu0.1O3-δ under hydrocarbon

conditions. However, the chemical stability of SrCe0.9Eu0.1O3-δ can be improved by partially

substituting Ce with Zr, which increases its tolerance factor and decreases its basicity. The H2

permeation of the SCZE721 membranes was essentially stable under WGS and SRM conditions.

H2 permeation through supported SCZE721 membranes is proportional to the

transmembrane H2 partial pressure gradient with a 1/4 dependence and controlled by bulk

diffusion with thickness down to 17 µm. A maximum H2 permeation flux of 0.23 and 0.21

cm3/cm2 min was obtained for the 33 μm thick SCZE721 membrane at 900 oC with the total flow

rate of 20 cm3/min and the feed gas composition of 100% H2 and 97% H2/3% H2O, respectively.

H2 permeation decreases with increasing steam partial pressure. The activation energy decreases

with increasing H2 partial pressure and/or decreasing steam partial pressure. Permeation flux

through the SCZE721 membrane is stable under wet H2, and conditions of WGS reaction and

SRM.

WGS reaction is exothermic and constrained by thermodynamic equilibrium limitations.

SrCe0.9Eu0.1O3-δ and SCZE721 thin film membranes incorporate H2 separation and WGS reaction

156

into one unit. Continuously removal of produced H2 moves the equilibrium toward the product

side, overcoming the thermodynamic equilibrium limitation. The XCO, H2 production, H2 yield

and the H2/CO in the reactor side effluent increase with increasing temperature and H2O/CO. A

XCO of 83.6% and 90.2% was achieved under 900 oC with H2O/CO = 1/1 and 2/1, respectively,

77% and 44% increase compared to the thermodynamic calculation data. The respective

improvement in H2 production was 73% and 42%. In contrast to the XCO, the permeated H2, H2

in the reactor side effluent and H2 total production increase with increasing flow rate and CO

concentration. The H2/CO in the reactor side effluent is variable through the SCZE721 H2

membrane. The SCZE721 membrane is stable under the WGS reaction conditions.

The XCH4 under thermodynamic equilibrium increases with increasing temperature and

decreasing CH4/H2O and CH4%. The experimental XCH4 is lower than the thermodynamic data

and limited by kinetic reaction rate. However, the H2 membrane can still enhance the SRM

performance by increasing 15% of both the XCH4 and total H2 production at 900 oC compared to

that with only Ni catalyst. The experimental XCH4 increases with increasing temperature and

decreasing CH4/H2O, CH4% and total flow rate. In contrast with the H2/CO, the SCO increases

with increasing temperature and decreases with decreasing CH4/H2O and increasing CH4% and

total flow rate. In contrast with the total H2 production, the H2 permeation decreases with

decreasing CH4/H2O. The H2 permeation and production increase with increasing temperature,

CH4% and total flow rate. The SCZE721 H2 membrane is stable under the SRM conditions.

The CH4 and CO2 conversion, CO selectivity and H2 production in the CSRM increase

with increasing temperature. 93.2% of CH4 conversion is achieved with CH4/CO2 = 1/2 at 900

oC. A maximum H2 permeation of 2.2 cm3/min (~0.2 cm3/cm2 min) and a maximum H2 total

production of 17.1 cm3/min are achieved at 900 oC with the CH4/CO2 = 1/1 and 1/2, respectively.

157

In contrast, the H2 selectivity and H2/CO decrease with increasing temperature. The CH4 and

CO2 conversion, H2 selectivity and H2/CO decrease with total flow rate. In contrast, the CO

selectivity and H2 production increases with increasing total flow rate. A maximum H2

permeation of 3.1 cm3/min (~0.26 cm3/cm2 min) and a maximum H2 production 32.8 cm3/min

were achieved at 900 oC with a total flow rate of 50 cm3/min. The H2/CO in the reactor side

effluent can be adjusted to desired value by combing the SRM and CDRM. The H2/CO between

700 oC to 900 oC is between 1.9-1.7 and 2.5-2.0 for CH4/CO2H2O = 2/1/1 and 2/1/1.5,

respectively.

10.2 Future Work

The objective of this research is to fabricate supported tubular thin film membranes by tape

casting followed by a rolling step, to improve the chemical stability of SrCe0.9Eu0.1O3-δ and to

incorporate H2 production and separation into one unit using SCZE721 membranes through

WGS reaction, steam reforming of CH4 and CO2 reforming of CH4. All these have been achieved

and demonstrated in chapter 3 to chapter 9. However, overall performance of the membrane

reactors can be improved in future work including material compositions and reactor design.

One future work is to modify the substrate composition. NiO-SCZ82 is the substrate in this

work. NiO is used to create porosity and to serve as the catalyst, by reduction to Ni when the

membrane is subsequently exposed to H2. However, Ni catalyst is embedded in the substrate. It

is difficult to replace the Ni since replacing the Ni usually means replacing the whole membrane.

The Ni particles are inevitable to grow during the sintering process of the membranes which

decreases its surface area. In addition, as shown in chapter 8, the XCH4 is lower than the

thermodynamic calculation result which is partially due to the inadequacy of the catalytic

activity. In addition, Ni is a catalyst for CH4 decomposition as well. The carbon deposited inside

158

the substrate will block the porous structure and even cause mechanically instability. Therefore, I

propose to substitute NiO with graphite as pore former and place Ni catalyst outside the

membrane. Ni catalyst can be synthesis separately to achieve optimized catalytic activity with

desired particle size and on different supports. In addition, much less amount of Ni is needed

compared to the amount used in the substrate above while maintaining enough activity for the

WGS, SRM and CDRM processes. Actually, I have successfully fabricated SCZE721 thin film

membranes coated on 30 vol% graphite/cellulose + 70 vol% SCZ82 substrates (Figure 10-1).

Another future work is to change the experimental setup and put the whole membrane into

the hot zone of the furnace. In the current setup, the membrane is installed at the bottom of the

quartz reactor. The bottom part of the tubular membrane is not in the hot zone and there is

temperature gradient along side with the membrane (Figure 5-1). This might cause mechanical

and chemical stability problem. Carbon deposition, for instance, is favorable at low temperature

in the WGS reaction, SRM and CDRM. Therefore, carbon deposition is more likely to happen in

that part of membranes outside the hot zone where the temperature is lower. In addition,

carbonate can be formed at low temperature which is detrimental to the chemical stability of the

membranes. One concern for the experimental setup change proposed above is the sealing. One

possible solution is sealing with metal rings with mechanic strings applying force on top.

159

Figure 10-1. A SrCe0.7Zr0.2Eu0.1O3-δ thin film membrane coated on graphite-SrCe0.8Zr0.2O3-δ substrate.

160

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BIOGRAPHICAL SKETCH

Jianlin Li was born in a small village in Zhaoqing, China. As early as he could remember,

Jianlin always wanted to attend elementary school with his sister. In 1984, he couldn’t wait any

longer. He registered at Fuluo Elementary School and paid the tuition fee with the money, which

he borrowed from his grandmother. In 1991, Jianlin attended the Xinqiao Middle School where

he spent six years and living by himself until he graduated high school. Jianlin excelled in

mathematics, physics and chemistry. Jianlin was ranked 2nd in mathematics among the three

hundred thousand students taking the college entrance exam.

In 1997, Jianlin attended the University of Science and Technology of China in Hefei

which is about 1000 miles away from his hometown. It was his first time taking a train. He

couldn’t forget the 26-hour experience he had on the crowded train for the rest of his life. He was

extremely excited when he finally arrived after such a long trip. Jianlin lived by himself and only

went home once a year during the winter holidays. Another challenge was having to speak

Mandarin when your native language is Cantonese. Jianlin exposed him to different disciplines.

He took classes in mathematics, physics, chemistry, computer science, electronic engineering

and international trade. He received his dual bachelor’s degrees in materials chemistry and

electronic information engineering in June 2001. Then, he found himself more interested in

materials science and engineering. He pursued his master’s degree under the supervision of Dr.

Chusheng Chen and senior engineer Pinghua Yang. He received his Master of Engineering

degree in materials science in June 2004. During his master studies, he worked on hydrogen

separation using mesoporous silicon membranes and syngas production through partial oxidation

of methane using oxygen permeable membranes such as Ba0.5Sr0.5Co0.8Fe0.2O3-δ.

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In August 2004, Jianlin joined Dr. Eric Wachsman’s group at the Department of Materials

Science and Engineering at the University of Florida.

It didn’t take long for Jianlin get used to the life in Gainesville, however, the language

barrier was a challenge at times. The last thing he wanted to do was to place an order by phone.

He really struggled with it for the first year.

Jianlin fell in love with college sports quickly, especially football. He had never seen a

football game before coming to America, but it only took him one season to become a football

fan. Jianlin felt lucky to witness the Gators win four national championships in basketball and

football in the last four years. He enjoyed the awesome experience cheering for the Gators in the

O’Dome and SWAMP. He also loved the experience of marching down University Avenue

every time the Gators were crowned with a national championship.

Jianlin was very active in student organizations. He was a two-term senator in the Student

Government at the University of Florida. He was the president of the Friendship Association of

Chinese Students and Scholars (FACSS) in 2005 and served as a consultant the following three

years. As president, Jianlin managed to get $15,000 from 14 sponsors to organize the 2006

FACSS Chinese New Year Show at the Phillips Center of Performing Arts. This is the biggest

event he ever organized and he felt proud of himself. In 2006, he was awarded the President of

the Year by their parent organization, the Volunteers for International Students Affairs (VISA).

Jianlin represented the University of Florida at the 6th Annual Mayor’s Summit in Tallahassee in

2006 and the 1st Florida International Leadership Conference in Ocala in 2007.

Jianlin enjoyed working with the group members in Dr. Wachsman’s group and

appreciated all the valuable discussion and comments from them. He received his Ph.D. from the

University of Florida in the fall of 2009.

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