The Chemical Looping Gasification of Biomass for Syngas

The Chemical Looping Gasification of Biomass for Syngas Utilization in a

Solid Oxide Fuel Cell System Simulated in Aspen Plus

by

Stephen Ganesh Gopaul

A Thesis

presented to

The University of Guelph

In partial fulfillment of requirements

for the degree of

Master of Applied Science

in

Engineering

Guelph, Ontario, Canada

© Stephen G. Gopaul, April, 2014

ABSTRACT

The Chemical Looping Gasification of Biomass for Syngas Utilization in a

Solid Oxide Fuel Cell System Simulated in Aspen Plus

Stephen G. Gopaul

University of Guelph, 2014

Advisor: Dr. Animesh Dutta

Co-Advisor: Dr. Ryan Clemmer

This thesis bridges the green energy fields of high-purity H2 synthesis gas (syngas) production

from biomass chemical looping gasification (CLG) and power generation via solid oxide fuel cell

(SOFC) operation. Two distinct CLG processes were simulated in Aspen Plus using the abundant,

nonconventional poultry litter biomass type. The first process (CLG 1) involved CO2 capture using

a CaO sorbent and generated lower yields of higher-purity syngas. The second process (CLG 2)

did not involve CO2 capture but used iron-based oxygen carriers to produce higher yields of lower-

purity syngas. The resulting syngas from either process was directly fed as fuel to a simulated

SOFC to determine operational viability. Both syngas types proved effective in the SOFC,

however CLG 1 syngas exhibited relatively higher performance overall. The results contribute to

both fields through novel approaches to the respective goals of each and by outlining the benefits

of an integrated CLG-SOFC system.

iii

Dedicated to my:

Parents: Ganesh Gopaul and Rozeena Gopaul

Fiancée: Leanna Harnarain

and

Siblings: Jason Gopaul and Anesha Gopaul

iv

Acknowledgements

The author wishes to express deep gratitude to, and appreciation for, both his advisor and co-

advisor, Dr. Animesh Dutta and Dr. Ryan Clemmer, respectively, who played key roles at every

stage of his study via conceptual discussion and input. Their guidance was vital for the successful

completion of the research stages of the author’s study and instrumental in bringing this thesis to

completion. The author would also like to thank them for their constructive comments and

invaluable suggestions throughout the progression of his graduate studies.

Additionally, the author acknowledges financial support from the Discovery Grant

Program funded by the National Science and Engineering Research Council (NSERC) of Canada,

the Dean’s Scholarship from the University of Guelph, and the Queen Elizabeth II Graduate

Scholarship in Science and Technology from the Government of Ontario in partnership with

private sector donors. The author also received financial support from his parents, Mr. Ganesh

Gopaul and Mrs. Rozeena Gopaul.

Finally, the author gratefully acknowledges the enduring inspiration and patience from his

parents, named above; fiancée, Ms. Leanna K. Harnarain; and siblings, Mr. Jason J. Gopaul and

Ms. Anesha Gopaul, whose love and support have encouraged the author throughout the duration

of his graduate studies.

v

Table of Contents

ABSTRACT .................................................................................................................................. ii

LIST OF TABLES ....................................................................................................................... ix

LIST OF FIGURES ...................................................................................................................... x

NOMENCLATURE ................................................................................................................... xii

CHAPTER I: Introduction ........................................................................................................ 1

1.1 Background ............................................................................................................... 1

1.2 Objectives ................................................................................................................. 4

1.3 Contribution .............................................................................................................. 5

1.4 Co-Authorship .......................................................................................................... 6

1.5 Organization of Thesis .............................................................................................. 6

References ....................................................................................................................... 8

CHAPTER II: Biomass Gasification Literature Review – A Review of Operating

Parameters for the Production of Hydrogen via Biomass Gasification ................................. 9

2.1 Introductory Remarks ............................................................................................... 9

2.2 Biomass Gasification .............................................................................................. 11

2.3 Steam Reforming .................................................................................................... 18

2.4 Chemical Looping Gasification .............................................................................. 19

2.5 Sorption-Enhanced Reaction .................................................................................. 22

2.6 Research Gaps ........................................................................................................ 32

2.7 Concluding Remarks .............................................................................................. 32

References ..................................................................................................................... 34

CHAPTER III: Solid Oxide Fuel Cell Literature Review – A Review of the Effect of Fuel

Type and Composition on Solid Oxide Fuel Cell Performance ............................................ 37

3.1 Introductory Remarks ............................................................................................. 37

3.2 Literature Review ................................................................................................... 39

3.2.1 Performance Comparison of H2 and Hydrocarbons .............................. 40

3.2.2 Performance Comparison of H2 and CO ............................................... 44

3.2.3 Performance Comparison of H2 and Syngas ......................................... 48

3.2.4 Performance in the Presence of Sulfur .................................................. 49

vi

3.3 Research Impact ...................................................................................................... 50


References ..................................................................................................................... 52

CHAPTER IV: Chemical Looping Gasification for Hydrogen Production – A Comparison

of Two Unique Processes Simulated Using Aspen Plus ......................................................... 53


4.1.1 Simulated Processes .............................................................................. 55

4.1.1.1 Chemical Looping Gasification Type 1 ............................... 55

4.1.1.2 Chemical Looping Gasification Type 2 ............................... 57

4.2 Feedstock Used ....................................................................................................... 59

4.2.1 Composition of Biomass Types ............................................................. 59

4.2.2 Chemical Equations for Gasification of Biomass Types ....................... 59

4.3 Simulation Input Parameters and Description ........................................................ 60

4.3.1 Setup and Calculation Methods ............................................................. 60

4.3.2 Component Definition and Input ........................................................... 60

4.3.3 CLG 1 Flowsheet Description ............................................................... 61

4.3.4 CLG 2 Flowsheet Description .............................................................. 64

4.4 Results and Discussion ........................................................................................... 66

4.4.1 Determining the Optimal Operating Conditions ................................... 67

4.4.1.1 CLG 1 Results ...................................................................... 67

4.4.1.2 CLG 2 Results ...................................................................... 71

4.4.2 Comparison of Simulation Results ........................................................ 74

4.4.2.1 Syngas Yield Comparison .................................................... 74

4.4.2.2 Syngas Composition Comparison ........................................ 75

4.5 Potential for Future Research ................................................................................. 76


References ..................................................................................................................... 78

vii

CHAPTER V: Tubular Solid Oxide Fuel Cell Operation on Syngas from Two Unique

Biomass Chemical Looping Gasification Processes – A Performance Comparison

Simulated Using Aspen Plus .................................................................................................... 80

Nomenclature ................................................................................................................ 80


5.1.1 Solid Oxide Fuel Cells ........................................................................... 81

5.1.2 Syngas from Biomass CLG and SOFC Operation ................................ 85

5.2 Simulation Description and Input Parameters ........................................................ 85

5.2.1 Simulated Process .................................................................................. 85

5.2.2 Setup and Component Definition .......................................................... 88

5.2.3 Flowsheet Description ........................................................................... 88

5.2.4 Cell Performance Calculation Methods ................................................. 89

5.2.4.1 Cell Voltage .......................................................................... 89

5.2.4.2 Electrical Efficiency ............................................................. 91

5.2.4.3 Power Output ........................................................................ 91

5.3 Results and Discussion ........................................................................................... 92

5.3.1 Syngas Performance Comparison .......................................................... 92

5.3.1.1 Effect of Syngas CO Composition ....................................... 95

5.3.1.2 Effect of Syngas CO2 Composition ...................................... 96

5.3.2 Anode Temperature Sensitivity Analysis .............................................. 99

5.3.3 Anode Pressure Sensitivity Analysis ..................................................... 99

5.3.4 Fuel Utilization Factor Sensitivity Analysis ........................................ 100

5.3.5 Current Density Sensitivity Analysis .................................................. 101

5.4 Potential for Future Research ............................................................................... 103

5.5 Concluding Remarks ............................................................................................ 104

References ................................................................................................................... 105

CHAPTER VI: Integrated CLG-SOFC System .................................................................. 107

CHAPTER VII: Conclusions and Recommendations ......................................................... 109

6.1 Overall Conclusions .............................................................................................. 109

6.2 Limitations of Research ........................................................................................ 110

6.3 Recommendations for Future Research ................................................................ 111

viii

APPENDICES ......................................................................................................................... 112

APPENDIX TABLE OF CONTENTS ....................................................................... 112

APPENDIX LIST OF TABLES ................................................................................. 113

APPENDIX LIST OF FIGURES ............................................................................... 114

Appendix A: Biomass Chemical Formula Calculations ............................................. 115

Appendix B: Raw Data for Biomass CLG Simulation ............................................... 118

Appendix C: Sensitivity Analyses for Biomass CLG Simulation .............................. 123

Appendix D: Raw Data for SOFC Simulation ........................................................... 127

ix

List of Tables

Chapter II

Table 2.1. Comparison of H2 production processes from biomass sources ................................ 10

Table 2.2. Optimal operating parameters from biomass gasification studies ............................. 15

Table 2.3. Optimal operating parameters from steam reforming studies ................................... 18

Table 2.4. Optimal operating parameters for CaO sorption-enhanced studies ........................... 24

Table 2.5. Chemical reactions in the steam gasification of biomass .......................................... 27

Chapter III

Table 3.1. Chemical reactions for SOFC operation on various fuels ........................................ 38

Table 3.2. Summary of performance data at 700 °C for H2 and various hydrocarbons ............ 42

Chapter IV

Table 4.1. Chemical reactions in the steam gasification of biomass .......................................... 54

Table 4.2. Chemical reactions in syngas chemical looping ........................................................ 58

Table 4.3. Ultimate analysis of poultry litter in both the presence and absence of sulfur and

nitrogen ....................................................................................................................................... 59

Table 4.4. Proximate analysis of poultry litter ............................................................................ 59

Table 4.5. Feed stream input conditions for CLG 1 simulation .................................................. 63

Table 4.6. Block unit operating conditions for CLG 1 simulation ............................................. 63

Table 4.7. Feed stream input conditions for CLG 2 simulation .................................................. 66

Table 4.8. Block unit operating conditions for CLG 2 simulation ............................................. 66

Chapter V

Table 5.1. Feed stream input conditions ..................................................................................... 87

Table 5.2. Block unit operating conditions ................................................................................. 88

Table 5.3. Reference conditions used in voltage calculations .................................................... 90

Table 5.4. Reference voltage as a function of current density .................................................... 90

Table 5.5. Lower heating value of fuels ..................................................................................... 91

Table 5.6. Standard operating conditions ................................................................................... 92

Table 5.7. Comparison of simulation results under standard conditions to the literature ......... 94

x

List of Figures

Chapter II

Fig. 2.1. Major pathways for H2 production from biomass sources ........................................... 11

Fig. 2.2. HyPr-RING biomass gasification schematic ................................................................ 13

Fig. 2.3. Effect of reaction temperature on product gas composition ......................................... 15

Fig. 2.4. Simulation of biomass gasification in interconnected fluidized beds .......................... 17

Fig. 2.5. Coal-direct chemical looping using Fe2O3 as an O2 carrier.......................................... 20

Fig. 2.6. Syngas chemical looping separating gasification and looping ..................................... 21

Fig. 2.7. Sorption-enhanced H2 production schematic ............................................................... 23

Fig. 2.8. H2 and CO temperature profiles for sorption-enhanced reaction at low pressure ........ 25

Fig. 2.9. Comparison of product gas composition for AER and standard gasification .............. 27

Fig. 2.10. Effect of Ca/C ratio on product gas composition ....................................................... 29

Fig. 2.11. H2 concentration at H2O/CH4 ratios for various CaO/CH4 ratios ............................... 31

Chapter III

Fig. 3.1. SOFC operation on H2 fuel ........................................................................................... 37

Fig. 3.2. Voltage and power output curve comparison for H2 and CH4 ........................................................ 40

Fig. 3.3. Voltage output curves for various CH4-H2O-N2 mixtures ........................................... 41

Fig. 3.4. Voltage output curves for ethane and ethene................................................................ 41

Fig. 3.5. Voltage output curves for H2 and propane at different temperatures ........................... 43

Fig. 3.6. Voltage and power output curve comparison for H2 and n-butane .............................. 43

Fig. 3.7. Voltage output curves comparing H2-H2O and CO-CO2 systems at 1000 °C with N2 as a

diluent ......................................................................................................................................... 45

Fig. 3.8. Effect of adding N2 diluent to H2 fuel on voltage and power output at 800 °C ........... 45

Fig. 3.9. Effect of adding CO2 diluent to CO fuel on voltage and power output at 800 °C ....... 46

Fig. 3.10. Voltage and power output for various H2-CO mixtures ............................................. 47

Fig. 3.11. Comparison of H2, CO, and syngas performance with Cu-CeO2-YSZ and Ni-YSZ

anodes ......................................................................................................................................... 47

Fig. 3.12. Effect of increasing CO content on voltage and power output using Cu-CeO2-coated

Ni-YSZ anode ............................................................................................................................. 48

Fig. 3.13. Concentration polarization effects from sulfur poisoning .......................................... 49

xi

Chapter IV

Fig. 4.1. CLG 1 simulation block diagram ................................................................................. 56

Fig. 4.2. CLG 1 process schematic ............................................................................................. 57

Fig. 4.3. CLG 2 simulation block diagram ................................................................................. 58

Fig. 4.4. CLG 1 simulation flowsheet ......................................................................................... 62

Fig. 4.5. CLG 2 simulation flowsheet ......................................................................................... 65

Fig. 4.6. CLG 1 reformer temperature sensitivity analysis ......................................................... 68

Fig. 4.7. CLG 1 reformer pressure sensitivity analysis............................................................... 69

Fig. 4.8. CLG 1 WGS reactor temperature sensitivity analysis .................................................. 71

Fig. 4.9. CLG 2 reducer temperature sensitivity analysis ........................................................... 72

Fig. 4.10. CLG 2 reformer syngas yield temperature sensitivity analysis .................................. 73

Fig. 4.11. Comparison of simulation syngas yields .................................................................... 75

Fig. 4.12. Comparison of simulation syngas compositions ........................................................ 76

Chapter V

Fig. 5.1. SOFC operation on H2 and CO fuels ............................................................................ 82

Fig. 5.2. Tubular and flat plate SOFC configurations................................................................. 84

Fig. 5.3. SOFC simulation block diagram .................................................................................. 86

Fig. 5.4. SOFC simulation flowsheet .......................................................................................... 87

Fig. 5.5. Cell voltage and electrical efficiency comparison under standard conditions ............. 93

Fig. 5.6. Total power output comparison under standard conditions .......................................... 93

Fig. 5.7. Effect of syngas CO composition on cell voltage and electrical efficiency ................. 95

Fig. 5.8. Effect of syngas CO composition on total power output.............................................. 96

Fig. 5.9. Effect of syngas CO2 composition on cell voltage and electrical efficiency ................ 97

Fig. 5.10. Effect of syngas CO2 composition on total power output .......................................... 97

Fig. 5.11. Effect of anode temperature on syngas performance under standard conditions ....... 99

Fig. 5.12. Effect of utilization factor on voltage and efficiency under standard conditions ..... 100

Fig. 5.13. Effect of utilization factor on total power output under standard conditions ........... 101

Fig. 5.14. Effect of current density on voltage and efficiency under standard conditions ....... 102

Fig. 5.15. Effect of current density on total power output under standard conditions .............. 102

xii

Nomenclature

Acronyms

AER Absorption Enhanced Reforming

AFC Alkaline fuel cell

AGC Advanced Gasification-Combustion

Ca/C Calcium-to-carbon ratio

CE Carbon efficiency

CLC Chemical looping combustion

CLG Chemical looping gasification

ESP Electrostatic precipitator

GE Gasification efficiency

GHG Greenhouse gas

HHV Higher heating value

HyPr-RING Hydrogen Production by Reaction-Integrated Novel Gasification

LHV Lower heating value

MCFC Molten carbonate fuel cell

NC No CO2

OCV Open circuit voltage

PAFC Phosphoric acid fuel cell

PEMFC Polymer electrolyte membrane fuel cell

PL Poultry litter

SBR Steam-to-biomass ratio

SCR Steam-to-carbon ratio

SCW Supercritical water

SEM Scanning electron microscopy

SOFC Solid oxide fuel cell

WC With CO2

WGS Water-gas shift

XRD X-ray diffraction

ZECA Zero Emission Coal Alliance

Chemical Formulae

Al2O3 Alumina

C Elemental carbon or graphite

C2H4 Ethene

C2H5OH Ethanol

C2H6 Ethane

C3H8 Propane

C4H10 n-Butane

CaCO3 Calcium carbonate

CaO Calcium oxide

CaSO4 Calcium sulfate

Ce0.9Gd0.1O1.98 Gadolinium-doped ceria (GDC)

CeO2 Ceria

CH3OH Methanol

CH4 Methane

CO Carbon monoxide

CO2 Carbon dioxide

Cr Chromium

Cu Copper

e- Electron

Fe Iron

Fe2O3 Hematite or iron (III) oxide

Fe3O4 Magnetite or iron (II,III) oxide

xiii

Gd Gadolinium

H Elemental hydrogen

H+ Hydrogen ion or proton

H2 Hydrogen gas

H2O Water or steam

H2S Hydrogen sulfide

He Helium

La0.8Sr0.2Cr0.95V0.02O3 LSCV

LaMnO3 Lanthanum manganate

Li Lithium

M Metal

MgO Magnesia

N Elemental nitrogen

N2 Nitrogen gas

Na Sodium

NH3 Ammonia

Ni Nickel

NiO Nickel oxide

O Elemental oxygen

O= Oxide ion

O2 Oxygen gas

Pd Palladium

Pt Platinum

Rh Rhodium

Ru Rubidium

S Elemental sulfur

Sm2O3 Samaria

SO2 Sulfur dioxide

Sr Strontium

TiO2 Titania

V Vanadium

Y2O3 Yttria

YSZ Yttria-stabilized zirconia

Zn Zinc

ZnCl2 Zinc chloride

ZrO2 Zinc oxide

Greek Alphabet

ΓH2 Fuel equivalent H2 content [kmol h-1]

ΔV Voltage difference [V = 1000 mV]

η Efficiency [-]

Latin Alphabet

Ac Active cell area [m2]

F Faraday’s constant [96 485 C mol-1]

i Current density [mA cm-2]

n Number of electrons transferred [mol e- per mol]

Nc Number of cells

�̇�𝑗 Molar flowrate of species j [kmol h-1]

P Pressure [atm = 1.01325 bar]

pc Cell power [W cell-1]

ptot Total power output [kW = 1000 W]

Q Thermal energy [kJ mol-1 = 1000 J mol-1]

T Temperature [°C]

xiv

Uf Fuel utilization factor [-]

V Voltage [V = 1000 mV]

W Work [kJ mol-1 = 1000 J mol-1]

Subscripts

an Anode

c Cell

cath Cathode

e Electric

(g) Gas

op Operating conditions

P Pressure

ref Reference conditions

(s) Solid

T Temperature

tot Total

1

Chapter I

Introduction

1.1 Background

Hydrogen (H2) has the potential to be a major contributor to the replacement of carbonaceous fossil

fuels as the primary global energy source, and ultimately become a significant benefactor to the

cause of climate change mitigation. World issues such as global atmospheric carbon dioxide (CO2)

levels and other greenhouse gas (GHG) emissions can be addressed with the use of H2

technologies. H2 also presents considerable advantages over other renewable technologies due to

its energy storage and transport capabilities [1].

Biomass sources can be processed for the production of H2 and other fuel gases such as

carbon monoxide (CO) and methane (CH4). Biomass is an umbrella term for organic materials

containing carbon (C), hydrogen (H), oxygen (O), nitrogen (N), and sulfur (S), and which have

stored sunlight in the form of chemical energy. Further, it is an abundant natural resource stemming

mainly from wood and wood residues, municipal solid wastes, aquatic plants, and agricultural and

animal wastes. Poultry litter is an abundant example of such animal wastes and is the focal biomass

type of this thesis. In 2012 the world chicken population exceeded twenty-one billion, or greater

than three chickens per person [2]. Thus, poultry excreta are an omnipresent and significant

potential source for solid biomass fuel. Biomass currently accounts for roughly 15 to 20% of fuel

utilization worldwide, though it is not a major fuel in contemporary industrial practises. However,

the use of biomass as an energy source is not a novel approach exclusive to modern society as it

has been used since prehistoric times for the purposes of heat generation. In addition to being an

effective, abundant fuel source, biomass sources are also considered to be CO2-neutral, and thus

aid in addressing the aforementioned issue of atmospheric CO2 concentration. Emissions of CO2

released during biomass conversion are equivalent to the amounts of CO2 absorbed by the organic

material via photosynthetic mechanisms. Biomass types are characterized through the use of

ultimate and proximate analyses. The former is an elemental analysis which considers the C, H,

O, N, S, and ash content of the biomass fuel, while the latter is a more qualitative analysis and

considers volatile matter, moisture, fixed carbon, and ash content [3].

2

The gasification of biomass is a widely used thermochemical process with the purpose of

converting the combustion value of the solid biomass to gaseous products. Common product gas,

or syngas, constituents include H2, H2O, CO, CO2, and CH4, as well as carbonaceous char and tar

by-products. The resulting gaseous fuel stream can be utilized in further downstream energy

production processing units, e.g. combustors, boilers, gas turbines, and fuel cells. Examples of

marketable products of gasification include H2, ammonia (NH3), methanol (CH3OH), gasoline,

oxo-alcohols (precursors for detergents and plasticizers), and various liquid oils and other fuels.

Biomass gasification encompasses four simultaneous processes in which the biomass particle

undergoes thermal decomposition. Drying occurs when moisture evaporates from the particle in

the form of steam; pyrolysis allows for the formation of gaseous components such as H2, CO, CO2,

steam, as well as liquid oils, tars, and solid char in the absence of oxygen (O2); gasification involves

endothermic reactions between solid char and the resultant gaseous products from pyrolysis; and

finally combustion includes exothermic reactions between O2 and the solid char and volatile matter

surrounding the particle. The chemical reactions involved in the endothermic stage of gasification

include the Boudouard, Water-Gas, Hydrogenation, Methanation, and Water-Gas Shift Reactions,

and are represented by Equations (1) to (5), respectively.

Boudouard reaction: C + CO2 ↔ 2 CO (1)

Water-gas reaction: C + H2O ↔ CO + H2 (2)

Hydrogenation reaction: C + 2 H2 ↔ CH4 (3)

Methanation reaction: CO + 3 H2 ↔ CH4 + H2O (4)

Water-gas shift reaction: CO + H2O ↔ CO2 + H2 (5)

Equations (1) and (4) proceed very slowly in the absence of catalysis. Equation (2) is prevalent in

the gasification mechanism when steam is present as the gasification medium, and Equation (3)

becomes important when H2 is used as the medium. Equation (5) is highly desirable in the system

when H2 is the desired gasification product, and its equilibrium is driven forward in-part by the

products of Equation (1) [3].

The gasification process, regardless of fuel type, is carried out in reactors called gasifiers.

A multitude of gasifier types exist which are optimized for various scenarios and desired products.

Fixed bed gasifiers, the oldest gasifier type, consist of the up-draft (countercurrent stream flow),

3

down-draft (co-current stream flow), and cross-draft schematics. These gasifiers produce low

purity H2 syngas that is highly diluted with nitrogen (N2) and CH4, thereby requiring downstream

reforming to convert the CH4 to H2. Entrained flow gasifiers are highly effective for coal

gasification and are capable of gasifying a vast array of coal types. However, the difficulty

involved in grinding biomass particles to sufficiently small particle sizes renders entrained flow

gasifiers less suitable for biomass gasification [4]. Moreover, fluidized bed gasifiers provide

greater versatility than the other gasifier types in terms of biomass gasification and have greater

high-purity H2 syngas production capabilities. These advantages stem from the greater degree of

solid mixing and uniform temperature of the fluidized bed, and the higher fuel flexibility of the

system. Consequently, H2 syngas purity can be further increased in the presence of a CO2 sorbent,

i.e. calcium oxide (CaO), by the formation of calcium carbonate (CaCO3). Desorption of CO2 from

the sorbent via CaCO3 dissociation and subsequent recovery and recycle of the sorbent constitutes

a syngas chemical looping gasification (CLG) system, which is the focus of a major portion of this

thesis. The sorbent particles are looped between the fluidized bed gasifier and the downstream

regenerator. Sorbent performance degradation over time is a major issue facing the field of biomass

CLG. Capture and sequestration of the separated CO2 inhibits its release to the atmosphere and

thus CLG processes are considered as CO2-negative when using CO2-neutral biomass fuels [3,4].

Fuel cells are electrochemical devices which directly convert the chemical energy stored

in the inlet fuel stream to electrical and thermal energy. Fuel cells generally consist of an anode

where electrochemical oxidation reactions occur, a cathode where reduction occurs, an electrolyte

which provides contact between the two electrodes, and an interconnecting material which

electrically joins the electrodes to allow for electron (e-) flow. Fuel cells generally operate on H2

fuel, though CO, CH4, CH3OH, and other carbonaceous fuels can be oxidized at the anode

depending on the fuel cell type. The anodic and cathodic electrochemical half-cell reactions for H2

operation are summarized in Equations (6) and (7) and the overall reaction for the system can be

seen in Equation (8).

Anodic oxidation: H2 → 2 H+ + 2 e- (6)

Cathodic reduction: ½ O2 + 2 H+ + 2 e- → H2O (7)

Overall reaction: H2 + ½ O2 → H2O (8)

Equations (6) through (8) vary depending on both the utilized fuel and fuel cell type.

4

Many fuel cell types exist; each with its own set of advantages and disadvantages. For

example, alkaline fuel cells (AFCs) are optimal for space applications as they exhibit high H2-O2

performance compared to other fuel cell types and do not require the use of precious metal

catalysts. However, the presence of CO2 in the fuel stream greatly reduces AFC performance.

Phosphoric acid fuel cells (PAFCs) were developed for the use of reformed hydrocarbon fuel,

thereby increasing CO2 tolerance, however precious metal catalysts such as platinum (Pt) are

required, and are easily poisoned by very small amounts of CO. Polymer electrolyte membrane

fuel cells (PEMFCs) exhibit no carbonaceous exhaust emissions and are scalable to a wide range

of applications, though PEMFCs are also greatly susceptible to CO poisoning as direct

consequence of precious metal catalyst requirements. Higher temperature fuel cells (operating

temperature > 600 °C) include molten carbonate fuel cells (MCFCs) and solid oxide fuel cells

(SOFCs), the latter of which is focused on in this thesis. MCFCs were developed to operate directly

on coal syngas. SOFCs operate between 800 and 1000 °C and are highly versatile in terms of fuel

operation [5]. They are capable of running on product streams from a vast array of syngas

production and reforming processes, and are therefore the optimal fuel cell choice for operation

on product syngas from biomass CLG.

The purpose of this thesis is to bridge the green energy fields of high-purity H2 syngas

production from biomass chemical CLG and power generation via SOFC operation. A simulation-

based approach was utilized to achieve the main objectives of the conducted research, and Aspen

Plus was the chosen simulation software. Aspen Plus is a commercially available software package

provided by AspenTech and developed in 1981 by the chemical engineering group at the

Massachusetts Institute of Technology (MIT) under a grant from the United States Department of

Energy. It includes a comprehensive thermodynamic and physical property database allowing for

simple process simulation and analysis. Built-in process unit modules (e.g. chemical reactors,

heaters, and separators) further simplify the user interface and usability of the program [6,7].

1.2 Objectives

The overall objective of the conducted research was the comparison of two novel, simulation-

based approaches to biomass CLG on the basis of their respective syngas composition and

production capabilities. SOFC operation on the resultant syngas types was simulated for the

purpose of performance comparison and was also a main objective of the research. Specific

objectives were also to:

5

Design and simulate two unique biomass CLG processes.

Determine optimal operating conditions for the simulated reactors in both CLG processes.

Compare resultant syngas from the CLG processes in terms of H2 yield and purity.

Simulate SOFC operation on the syngas from both CLG processes.

Compare the results of SOFC operation on syngas in terms of cell voltage, electrical

efficiency, and total power output.

Determine the effects of syngas CO composition on SOFC performance.

Study the effects of syngas CO2 composition on SOFC performance.

Research the effects of varying operating parameters on SOFC performance.

Compare the simulated results to values published in the literature.

1.3 Contribution

This thesis represents a valuable addition to the efforts of bridging the renewable energy fields of

biomass CLG and SOFC operation. Novel approaches to the respective goal of each field, H2

production and power generation, are designed, simulated, and the results compared. Specific

contributions also include the:

Development of a CLG process for H2-rich syngas production with in situ CO2 capture via

CaO sorbent, total sorbent recovery and tar reforming, and simulated using the Aspen Plus

software.

Design of a CLG process for high yields of majority-H2 syngas using Fe-based oxygen

carriers, with near-total carrier recycle, and simulated using Aspen Plus.

Direct comparison of aspects from the two resultant syngas types using the same,

nonconventional biomass feedstock, i.e. poultry litter, in both cases.

Direct tubular SOFC performance comparison of the two biomass CLG syngas types under

the same feed and operating conditions using Aspen Plus.

Investigation of the effects of CO levels in the syngas feed on simulated SOFC

performance.

Study of the effects of syngas CO2 composition on SOFC performance to determine if its

removal is a net benefactor to the combined CLG-SOFC system.

6

1.4 Co-Authorship

The following are aspects of the conducted research that underwent the peer-review process and

were published in a journal, or were submitted to a peer-reviewed journal. Although the listed

authors provided invaluable contributions in the form of suggestions, input, and critiquing, the

author of this thesis was the principal author of the listed works.

Chapter IV

Title: Chemical looping gasification for hydrogen production: A comparison of two unique

processes simulated using ASPEN Plus.

Authors: Stephen G. Gopaul, Animesh Dutta, and Ryan Clemmer.

Published in: The International Journal of Hydrogen Energy

Citation: Gopaul, S.G., A. Dutta, and R. Clemmer. “Chemical Looping Gasification for Hydrogen

Production – A Comparison of Two Unique Processes Simulated Using ASPEN Plus.” Int J

Hydrogen Energ 39 (2014): 5804-5817.

Chapter V

Title: Tubular solid oxide fuel cell operation on syngas from two unique biomass chemical looping

gasification processes: A performance comparison simulated using Aspen Plus.

Authors: Stephen G. Gopaul, Ryan Clemmer, and Animesh Dutta.

Submitted to: The International Journal of Hydrogen Energy

1.5 Organization of Thesis

A brief description of the main chapters of the thesis is provided below.

Chapter II

This chapter begins with the importance of H2 as a clean fuel and its potential to replace fossil

fuels as the dominant energy source. It then moves on to explain its connection to biomass, a

renewable energy source. Literature regarding thermochemical biomass conversion pathways such

as biomass gasification, steam reforming, CLG, and sorption-enhanced H2 production are

reviewed. The optimal operating parameters such as temperature, pressure, steam/carbon ratio, and

calcium/carbon ratio for a multitude of studies utilizing these biomass conversion pathways are

7

also summarized. Potential research gaps in the field of biomass conversion to H2 are identified

towards the end of the chapter.

Chapter III

This chapter begins with background information regarding SOFC operation and inherent

advantages and disadvantages to other fuel cell types posed by SOFC utilization. The review

proceeds to focus on determining the effects of fuel type and composition on SOFC performance

in terms of voltage output and power density. Performance from H2 operation is compared to other

fuels such as CO, CH4, coal syngas, biomass syngas, and other fuel mixtures, and the effects of

alternative anode materials are reviewed. Performance degradation due to sulfur poisoning is also

considered in this chapter.

Chapter IV

This chapter presents the comparison between two biomass CLG processes for the production of

H2 simulated using Aspen Plus. The optimal operating conditions are determined via temperature

and pressure sensitivity analyses of the constituent reactors for each process. The end of the chapter

discusses the H2 purity and production capability differences between the two processes.

Chapter V

This chapter presents the performance simulation of a tubular SOFC operating on syngas generated

from two unique biomass CLG processes using Aspen Plus. The effects of syngas CO and CO2

composition on performance are discussed. Anode temperature and pressure, fuel utilization

factor, and applied current density sensitivity analyses are conducted and shown towards the end

of the chapter. Performance values are compared to those reported in the literature.

Chapter VI

This chapter briefly summarizes the benefits of combining the CLG and SOFC processes into a

single, integrated system. The recycling of high quality heat from the SOFC stack exhaust stream

to high temperature reactors from the CLG processes is emphasized upon.

Chapter VII

The overall findings from the conducted research and limitations posed by the utilized methods

and assumptions are detailed in this chapter. Recommendations for further study are also

summarized.

8

References

[1] Levin, David B., and Richard Chahine. "Challenges for Renewable Hydrogen Production from Biomass." Inter J

Hydrogen Energ 35 (2009): 4962-969.

[2] Food and Agriculture Organization of the United Nations. “FAOSTAT – Production – Live Animals.” FAOSTAT.

7 Feb. 2014. Web. 26 Mar. 2014. <http://faostat.fao.org/site/573/DesktopDefault.aspx?PageID=573#ancor>.

[3] Fan, Liang-Shih. Chemical Looping Systems for Fossil Energy Conversions. Hoboken, NJ: Wiley-AIChE, 2010.

Print.

[4] Acharya, Bishnu. Chemical Looping Gasification of Biomass for Hydrogen-Enriched Gas Production. Thesis.

Dalhousie University, Halifax, Nova Scotia, 2011. Dalhousie University, Department of Mechanical Engineering.

Print.

[5] Li, X. Principles of Fuel Cells. New York: Taylor & Francis, 2006. Print.

[6] Aspen Plus 11.1 Users Guide, 2002. AspenTech Ltd., Cambridge, MA, USA.

[7] Zhang, W., E. Croiset, P.L. Douglas, M.W. Fowler, and E. Entchev. “Simulation of a Tubular Solid Oxide Fuel

Cell Stack using AspenPlusTM Unit Operation Models.” Energ Convers Manage 46 (2005): 181-196.

9

Chapter II

Biomass Gasification Literature Review – A Review of

Operating Parameters for the Production of Hydrogen via

Biomass Gasification

2.1 Introductory Remarks

Hydrogen (H2) has the potential to revolutionize the global energy industry and reduce and

subsequently eliminate our reliance on carbonaceous fossil fuels. H2 presents a viable alternative

to fossil fuels as its use in H2 fuel cells results in clean energy with minimal polluting or greenhouse

gas (GHG) emissions. Further, H2 is a viable alternative compared to other natural energies such

as solar and wind power due to its ability to store energy and be used as a medium for energy

transport. However, the production of H2 from current technologies requires H2 consumption and

subsequent generation of GHGs [1]. Biomass presents a renewable and environmentally friendly

alternative feedstock for H2 production.

Biomass is an organic fuel source generally consisting of carbon (C), hydrogen (H), oxygen

(O), nitrogen (N), and sulphur (S). Biomass can come in many forms including, but not limited to:

animal wastes, municipal solid wastes, crop residues and other agricultural wastes, and saw dust

[2]. The use of biomass as a fuel source is not a novel approach as it has been used since pre-

historic times as a source of heat generation. Burning wood to cook meat is exemplar of such

practices. Consequently, a more modern approach is the thermochemical or biological conversion

of biomass to H2 for use in subsequent energy-producing applications (e.g. fuel cells). Continually

rising prices for hydrocarbon-derived energy is causing H2 production from biomass to become

increasingly favourable as biomass becomes relatively cheaper. Moreover, in Canada, the most

rapidly growing demand for H2 comes from the upgrading of heavy oil in the oil sand industries

of Alberta [1].

Although the conversion of biomass to H2 does present a viable alternative to widespread

fossil fuel utilization, significant physicochemical and economic shortcomings do exist. For

example, on a weight basis, H2 is more efficient as a fuel than both oil and natural gas. However,

gaseous H2 is eight times lighter than methane and liquid H2 is ten times lighter than gasoline.

10

Thus, H2 has a relatively low volumetric energy density overall. Still, this fact has favourable

implications when considering logistics costs and energy transport. Furthermore, purification

issues exist when considering separating H2 out of product gas streams. Palladium (Pd)-based

membranes are viable technologies for H2 separation due to their high H2 permeability, although

they are rapidly deactivated by trace amounts of sulphur in the gas stream. The utilization of

palladium-metal (Pd-M) membranes has been found to improve membrane performance in the

presence of sulphur. Additionally, the degree of purification required depends on the market and

application type [1]. Economic barriers also exist. The recent contraction and relative decline of

the North American automotive industry has postponed the expected use of H2 in automobiles as

a replacement fuel for gasoline and diesel. Assuming full market penetration, roughly 40 million

tonnes of H2 per year would be required to run 100 million fuel-cell operated cars [1]. However,

the wide-scale use of H2 in automobiles requires investment and subsequent development of

necessary infrastructure and on-board storage techniques [3]. Energy costs also play a large role

in the economic barriers. Using current technologies, a 400 MWth input can produce H2 at US$ 8

to 11 per GJ given that biomass can be purchased for US$ 2 per GJ. This figure is roughly double

current gasoline production prices of US$ 4 to 6 per GJ. Reduction in biomass expense coupled

over the long-term with increased capital investment, technological advancements, and larger-

scale applications could potentially reduce the H2 production cost to approximately US$ 6 per GJ,

thus matching current gasoline production costs [3].

Table 2.1

Comparison of H2 production processes from biomass sources. Adapted from [1].

Process Biomass feedstocks Efficiency By-products

Thermochemical conversion processes

Steam reforming Methane, glycerol,

alcohols, polyols, sugars,

organic acids

70-85% CO, CO2, C10-C22 chains

Aqueous reforming Glycerols, alcohols,

polyols, sugars, organic

acids

35-100% CO, CO2, alkanes,

alcohols, polyols, organic

acids

Electrolysis H2O and electricity 50-60% None

Partial oxidation Hydrocarbons 60-75% No data

Biomass gasification Biomass 35-50% CO, CO2, CH4

Biological conversion processes

Photolysis H2O and sunlight 0.5% None

Photo-fermentation Organic acids and

sunlight

0.1% CO2

Dark fermentation Lignocellulosic biomass 60-80% CO2

11

The focus of this review is the outlining of the current processes for H2 production from

the thermochemical conversion of biomass. Consequently, Table 2.1 [1] summarizes current H2

production processes, their respective efficiencies and by-products, including both the

thermochemical and biological methods. Figure 2.1 shows the major pathways for H2 production

from biomass [4].

Fig. 2.1. Major pathways for H2 production from biomass sources [4].

2.2 Biomass Gasification

Biomass gasification is the thermochemical conversion of biomass for the production of a

combustible gaseous product stream, i.e. synthesis gas (syngas). It is carried out in the presence of

a gasification agent and is generally applicable to biomass containing less than 35% moisture. Tar

formation during the gasification process is a major concern as it results in the generation of more

complex by-products downstream. These by-products hinder H2 production, reduce the purity, and

contaminate the product gas stream [5]. The use of a rhodium (Rh)/ceria (CeO2)/M (where M is

silica, alumina, or zirconia) catalyst has been found to reduce tar formation, however this catalyst

is relatively costly and more research is required to develop a more cost-effective one [2]. CeO2-

12

zirconia (ZrO2) catalysts were found to perform better than Al2O3-supported ones in terms of H2

production [6].

Gasification is a favourable method to pyrolysis for H2 production as it aims to generate

the products in a gaseous state [5]. Further, H2 yield from gasification is generally greater than that

from pyrolysis [7]. To elaborate upon the aforementioned moisture content restriction, biomass

gasification is in fact possible for biomass containing > 35% moisture, given that it is carried out

in supercritical water (SCW) conditions. SCW is that which is subjected to temperatures and

pressures exceeding 374.3°C and 221.2 bar, respectively [4]. The properties of liquid and vaporous

H2O at the critical temperature and pressure are indistinguishable. SCW gasification eliminates the

need for any biomass drying due to its ability to withstand relatively high moisture contents [8]

and can be carried out both at lower (350 to 600 °C) and higher (> 600 °C) temperatures [4].

Consequently, SCW gasification has some advantages inherent to the thermophysical properties

of SCW itself. These include, but are not limited to [9]:

Process reactions finish rapidly and completely due to lack of mass transfer limitations.

Water is more readily separated from the products after the process is complete. This can

be done simply by changing operating parameters, and thus is advantageous when

compared to other separation methods.

The life of the active catalyst is lengthened due to decreased generation and subsequent

deposition of coke.

Moreover, 100% gas conversion can only be achieved at lower temperatures via utilization of a

bimetallic rubidium (Ru) or nickel (Ni) catalyst supported on titania (TiO2), ZrO2, or carbon.

Catalyst use is not required for high-temperature SCW gasification [4].

A more recent biomass gasification technology is the Hydrogen Production by Reaction-

Integrated Novel Gasification (HyPr-RING) method, in which the H2 production and gas

separation reactions are carried out in the same reactor at relatively lower temperatures. The

process is depicted in Figure 2.2 [5,7].

Variation of gasification operating parameters also has a significant effect on the

thermophysical properties of the product gas stream. For example, higher temperatures are

generally favourable to H2 production by gasification. In addition, reducing gasification pressure

by 10% was found to result in negligible increases in H2 production – amounting to less than a

13

0.2% increase [8]. Higher steam/biomass ratio (SBR) also favours H2 production. Methane (CH4)

and solid carbon that would be produced at lower SBRs are completely converted to H2 and CO at

higher steam flowrates. Furthermore, the use of O2 for gasification generates higher quality syngas

with a higher heating value (HHV) range of 10 to 15 MJ Nm-3. A high temperature range of 1000

to 1400 °C can be achieved using this method. In comparison, syngas generated by air gasification

has a HHV of only about 4 to 6 MJ Nm-3 with undesired by-products such as H2O, carbon dioxide

(CO2), hydrocarbons, and tars [8].

Fig. 2.2. HyPr-RING biomass gasification schematic [5,7].

14

Many studies have been conducted to determine various ways of improving H2 production

from biomass gasification. Some of the final results of the reviewed studies are summarized in

Table 2.2. For example, Moghtaderi [10] focused research on controlling operating parameters

such as reaction temperature and heating rate to determine the effect on product gas yield using

radiata pine dust. Higher temperatures and higher steam rates were found to increase gas yield.

Maximum H2 production from the low-temperature catalytic steam gasification of biomass was

observed at the optimal conditions of reaction temperature of 600 °C, steam content of roughly 90

mol-%, and a residence time of 20 min. This result was additionally improved with the use of a

Ni-based catalyst with a nickel oxide (NiO) molar ratio of 50%. Figure 2.3 shows the effect of

reaction temperature on product gas composition (at 90 mol-% steam and 20 min residence time)

from this study [10]. The figure on the left shows that 600 °C was the approximate critical point

where any reduction in reaction temperature would result in significant loss of H2 production.

Further, increasing reaction temperature beyond this point resulted in minimal (if any) increase in

H2 production accompanied by an increase in potentially adverse by-products such as CO. The

figure on the right shows the effects catalyst use on the product gas yield. It can be seen that by-

product generation was significantly reduced at temperatures above 600 °C.

Furthermore, Anuadala et al. [11] presented an analytical model of sawdust wood

gasification which predicted H2 production over the range of 10 to 32 kg biomass per second.

Biomass was fed to a gasifier at 727 to 1227 °C and was accompanied by steam at 227 °C. It was

found that input temperature and quantities of steam and biomass affected the H2 production rate.

80 to 130 g H2 / kg biomass was produced over the biomass feed range, whereas 80 g H2 / kg

biomass was produced at the gasifier operating temperature range. H2 constituency in the product

gas stream varied from 51 to 63% with 4.5 kg s-1 of steam and narrowed to 51 to 53% with 6.3 kg

s-1 of steam.

Guo et al. [12] conducted an investigation aimed at H2 production from biomass

gasification in SCW. It was found that H2 yield, gasification efficiency (GE), and carbon efficiency

(CE) all increased with an increase in temperature. Also, H2 yield, GE, and CE all slightly

increased with a pressure increase from 25 to 30 MPa. However, the increase in all three

parameters was too small to be considered significant. Further, the three parameters decreased with

increasing biomass content, as gasification became more difficult under these conditions. Higher

temperatures were required to cope with higher biomass content. Both H2 and CO2 yield increased

15

rapidly with increasing residence time. Conversely, CO production tended to behave in the

opposite manner. Thus, higher residence times were found to be favourable, albeit at the expense

of production time losses. H2 and CH4 yield decreased with increased oxygen (i.e. oxidizer)

addition. CO2 yield notably increased under the same conditions.

Table 2.2

Optimal operating parameters from biomass gasification studies.

Biomass type Maximum

H2 content

Gasifier

temperature

Steam /

biomass

Residence

time Catalyst used Reference

Radiata pine

sawdust

~1.50 m3/kg

of biomass 600 °C 0.9 20 min 50 mol-% NiO [10]

Sawdust wood 51-63% 727-1227 °C (6.3 kg

steam/s)a - - [11]

Olive oil waste 70 mol-% 900 °C - 7-10 min 5 wt.-%

ZnCl2b

[14]

- 60.5% 750-800 °Cc 0.6-0.7 - - [15]

Softwood,

Eucalyptus

globulus, and

hardwood

-d 830 °C 0.6-0.7 - - [16]

a Specific SBR value not given. b Catalyst use was optimal for H2 production at 800 °C and resulted in 69 mol-% H2. c Optimum combustor temperature was 920 °C. Also, recirculation of 4 to 14 bed particles was required to maintain

constant gasifier temperature. d No specific maximum H2 content value given; however H2 content increased by 10 to 20% with an increase in

temperature.

Fig. 2.3. Effect of reaction temperature on product gas composition [10].

16

Correspondingly, Furusawa et al. [13] also conducted research on biomass gasification in

SCW; however lignin was used in conjunction with Ni catalysts. From their research, it was

concluded that an optimal Ni particle size exists for the catalytic gasification of lignin biomass in

SCW. This conclusion was made on the basis that the carbon yield of gas products increased with

increasing Ni surface area, with the exception of 10 wt.-% Ni/magnesia (MgO) catalyst calcined

at 500 °C. Under the tested conditions, the 10 wt.-% Ni/MgO catalyst calcined at 600 °C produced

a carbon yield of 30%; the highest of the tested catalysts. Therefore, Furusawa et al. deemed this

catalyst optimal for gasification of lignin in SCW.

González et al. [14] studied the reactions influencing biomass air and air/steam gasification

for H2 production. The influence of zinc chloride (ZnCl2) and dolomite catalysts were also

investigated. It was found that maximum H2 production was 70 mol-% H2 in the product gas

stream. This was attained at a reactor temperature of 900 °C and at a residence time range of 7 to

10 min. The use of 5 wt.-% ZnCl2 catalysts were most effective at 800 °C and resulted in the

product gas stream being composed of 69 mol-% H2 with a residence time of 5 min. No significant

improvement in H2 production could be seen at 900 °C.

Shen et al. [15] simulated biomass gasification in interconnected fluidized beds. The

simulation schematic can be seen in Figure 2.4 [15]. The effect of gasifier temperature and SBR

on fuel gas composition, H2 yield, carbon conversion of biomass, recirculation of bed particles,

and other parameters were studied. The fluidized beds included a circulating bed for air-fed

combustion and a bubbling bed for steam-fed gasification. Optimum gasifier temperature was

found to be in the range of 750 to 800 °C. H2 content in this range reached a maximum of 60.5%

with an optimum combustor temperature of 920 °C and maximal SBR in the range of 0.6 to 0.7.

An increase in SBR showed a steady increase in H2 and CO2 content, while CO content decreased

steadily. Subsequently, this optimal value was found to decrease with increasing gasifier

temperature. The trend observed for carbon conversion was a decline with increasing gasifier

temperature and SBR. Furthermore, the recirculation of bed particles required to maintain a

constant gasifier temperature increased with an exponential trend for each incremental increase in

gasifier temperature. The optimal value was found to be between 4 and 14 particles at optimal

gasifier and SBR conditions.

Consequently, Franco et al. [16] investigated the reactions influencing biomass gasification

by varying biomass type. Softwood, Eucalyptus globulus, and hardwood were utilized in the study.

17

The optimum gasifier temperature was 830 °C. Increasing temperature resulted in greater product

gas generation and a reduction in the generation of various hydrocarbons. H2 content increased by

10 to 20% with the increase in temperature coupled with a reduction of hydrocarbon and tar

generation by 3 to 5%. An SBR of 0.6 to 0.7 was found to be optimal. Switching between biomass

types was not found to have a significant effect on product gas composition. This is generally a

good result since biomass availability varies seasonally depending on the regional climate.

Furthermore, the water-gas shift (WGS) reaction was dominant for eucalyptus and hardwood at

the optimal gasifier temperature. The Boudouard and WGS reactions were significant over the

entire tested temperature range for softwood.

Fig. 2.4. Simulation of biomass gasification in interconnected fluidized beds [15].

18

2.3 Steam Reforming

After the gasification process has taken place, the production of purer H2 can be achieved by steam

reforming followed by a WGS reactor. H2 is produced during the WGS reaction where CO and

steam are converted to CO2 and H2. Both high- and low-temperature WGS reaction methods

require the use of a catalyst: iron (Fe) and/or chromium (Cr)-based oxide catalysts at high

temperatures and copper (Cu)-zinc (Zn) oxide catalysts at low temperatures [4]. Although O2

would be ideal in steam reforming, O2 separation units are cost-intensive and are not practical for

smaller scale plants [2].

As with biomass gasification, many studies exist in which researchers utilized steam

reforming for the production of higher purity H2. The results from some of the reviewed studies

which use the steam reforming method are summarized in Table 2.3.

Table 2.3

Optimal operating parameters from steam reforming studies.

Biomass type Maximum

H2 content

Optimal

temperature

Steam /

carbon

Optimal

pressure

Calcium /

carbon Reference

1:1 biomass to

crude glycerin

mixture

- a 700-750 °C 1.7-2.25 100 kPa 1 [18]

Acetic acid,

ethylene glycol,

and acetone

80-90% 627 °C 9 b 1 atm - [19]

a No specific maximum H2 content value given; however H2 yield increased linearly from 0.053 to 0.059 mol H2 / kg

biomass with an increase in temperature from 650 to 825 °C. b Increasing SCR from 1 to 9 improved H2 production by 20% at 627 °C and saw a maximum value of 64.4 mol-%.

The specific SCR for the tests yielding 80 to 90% H2 content was not given.

Wang et al. [17] found that variation in temperature had the most significant effect on H2

yields. Moreover, at 600 °C, varying residence time from 0.04 to 0.15 s and increasing

steam/carbon ratio (SCR) from 4.5 to 7.5 did not have a significant effect on H2 production.

Furthermore, Chen and Zhao [18] investigated the co-steam-reforming of a 1:1 biomass to crude

glycerin mixture and determined that higher temperatures and lower pressures favoured H2

production. It was further determined that the optimal temperature range for H2 production was

700 to 750 °C. A temperature increase from 650 to 825 °C resulted in a linear increase in H2

production from 0.053 to 0.059 mol H2 / kg biomass. Optimal pressure was roughly atmospheric

at 100 kPa. The SCR that maximized H2 production was determined to be between 1.7 and 2.25

and a clear increase in H2 production was observed with increasing SCR. Finally, H2 production

19

was maximized at a calcium/carbon ratio (Ca/C) of unity, with no significant changes in gas

production beyond this value.

Thermodynamic analysis of H2 production via steam reforming of acetic acid, ethylene

glycol, and acetone was carried out by Vagia and Lemonidou [19]. It was found that maximum H2

yield ranged from 80 to 90% at 627 °C. Increased H2 content at higher operating temperatures was

accompanied by the increased presence of CO. Increasing SCR favoured H2 production over the

temperature range studied such that an increase from 1 to 9 improved H2 production by 20% at the

optimal operating temperature. Maximum H2 composition of 64.4 mol-% was obtained at a SCR

of 9. Higher pressures decreased H2 content in the product stream. Increasing pressure from 1 to

20 atm decreased H2 content from 68 to 49 mol-%. Thus, optimum pressure was found to be 1 atm.

The production of 1 kmol/s of H2 via bio-oil steam reforming required roughly the same amount

of energy as with natural gas reforming. This was demonstrated with the utilization of material

and energy balances over the entire system.

2.4 Chemical Looping Gasification

In general, chemical looping systems utilizing solid fuels (e.g. biomass) consist of a solid carrier

in a loop between two reactors. These reactors may be a combination of fixed and fluidized bed

reactors. The two main types of chemical looping systems include chemical looping combustion

(CLC) and chemical looping gasification (CLG). However, the latter will be focused on in this

review. CLG combines high-purity H2 production with CO2 capture, making it an attractive

thermochemical biomass conversion process. Two types of CLG exist. They differ based on the

type of solid carrier used – either an O2 or CO2 carrier [20].

CLG using an O2 carrier includes the cycling of the carrier between a fuel reactor and an

air reactor. A metal oxide is used as the carrier and is reduced in the fuel reactor thereby freeing

up oxygen for use by the fuel to generate the product gases. Subsequently, the reduced metal oxide

is cycled to the air reactor where it is oxidized upon contact with the air stream. Inert N2 is thus

prevented from diluting or combining with the product gas stream, ultimately eliminating any

downstream separation requirements. CLG using an O2 carrier is further divided into two

processes: biomass directly gasified in the fuel reactor and biomass gasified in a separate gasifier

prior to entering the looping system [20].

In the process where biomass is directly gasified in the fuel reactor, O2 levels are

maintained sufficiently low, below the stoichiometric level, so as to avoid biomass combustion.

20

Fan demonstrated that iron oxide (Fe2O3) was a suitable O2 carrier for this process [21]. They

developed the coal-direct chemical looping system utilizing Fe2O3 as an O2 carrier for the

production of pure H2. This process is depicted in Figure 2.5 [21]. Conversely, Scott et al. found

that significant problems occurred when utilizing Fe2O3 as an O2 carrier. Gasification was found

to be a limiting step [20]. Moreover, it was not well understood if the Fe2O3 carrier would retain

its efficiency upon increased oxidation-reduction cycling in the long-term. Furthermore, Leion

proposed that tar, char, and ash particle deposition on the Fe2O3 carrier surface reduced its efficacy

[22].

Fig. 2.5. Coal-direct chemical looping using Fe2O3 as an O2 carrier. Adapted from [21].

21

Subsequently, separating the biomass gasification and looping system processes was

developed to mitigate the shortcomings of the aforementioned process. The syngas chemical

looping system proposed by Fan is exemplar of this process, and is illustrated in Figure 2.6 [21].

In this process, pure streams of both H2 and CO2 were generated via the condensation of a steam

and CO2 stream.

Fig. 2.6. Syngas chemical looping separating gasification and looping. Adapted from [21].

In contrast, CLG using a CO2 carrier involves the use of a sorbent, usually calcium oxide

(CaO), being cycled between gasifying and regenerating reactors. As such, the sorbent undergoes

a series of calcination-carbonation cycles whereby CO2 is absorbed by CaO for the generation of

solid calcium carbonate (CaCO3). This facilitates the production of a pure CO2 product stream.

This process is reiterated and further explained in a later section of this review. Processes such as

HyPr-RING, Zero Emission Coal Alliance (ZECA), ALSTOM hybrid gasification-combustion,

22

and Advanced Gasification-Combustion (AGC) are all exemplar of CLG with the utilization of

CaO as a CO2 carrier. The major difference between the processes is their respective method for

providing sufficient energy for the calcination reaction. Note, however, that these methods are best

adapted for coal gasification under current technological conditions. Utilizing coal, these processes

were able generate H2 in the product stream at levels of 65 to 80% [21].

2.5 Sorption-Enhanced Reaction

Sorption-enhanced H2 production combines hydrocarbon reforming, WGS, and CO2 capture and

separation processes into a single step, thereby greatly enhancing H2 purity in the product gas

stream. Furthermore, CO and CO2 are reduced to parts-per-million amounts. Sorption-enhanced

H2 production is an inherently transient process as the sorbent is consumed during reaction. This

requires that the sorbent be regenerated in a downstream process to make sorption-enhanced

reaction economically viable. Consequently, maximization of sorbent life is paramount for process

efficiency [23].

The sorbent is generally mixed with the reforming catalyst so both processes can occur

simultaneously [24]. Sorbents can be calcium-based oxides, K-promoted hydrotalcites, and mixed

metal oxides of lithium (Li) and sodium (Na). The mechanism of CO2 capture is that the CO2 is

absorbed by the solid sorbent, subsequently generating a CO2-solid complex. For example: if

calcium oxide (CaO) is used as the sorbent, CaO and CO2 will react to form solid CaCO3,

effectively removing CO2 from the product gas stream [23].

Some general advantages presented by sorption-enhanced H2 production when compared

with typical biomass gasification include, but are not limited to [25]:

Only two processing steps are required.

Lower temperatures and pressures are utilized.

Predicted energy savings of 20 to 25%.

Dolomite can be used as a sorbent source; it is relatively inexpensive and abundant.

Sorbent regeneration produces pure CO2 that can either be utilized in further applications

or sequestered.

A general schematic of sorption-enhanced reaction can be seen in Figure 2.7 [25]. The majority of

the reviewed studies utilize CaO as the sorbent of choice. Grasa and Abanades [26] studied the

23

effect of continuous cycling on the ability of CaO sorbent to capture CO2. It was found that capture

capacity decreased significantly within the first 20 cycles, but tended to become steady with further

cycling. This residual conversion range was stable for up to 500 cycles. Consequently, the residual

conversion range was reached more rapidly for both calcination temperatures exceeding 950 °C

and longer calcination times, as the deactivation rate constant was favoured under these conditions.

Sorption-enhanced H2 production has been extensively studied and well-reported in the

literature. The results of select studies are summarized in Table 2.4. Ortiz and Harrison [27]

obtained a purity of > 95 mol-% H2 due to sufficiently rapid rates of the reforming, WGS, and CO2

removal reactions. The majority of the Ni-based reforming catalyst and Ca-based sorbent mixture

regeneration conditions resulted in minor activity loss, even after 25 cycles. Furthermore, it was

found that most of the activity loss was attributed to the Ca-based sorbent which is inexpensive

relative to the catalyst.

Fig. 2.7. Sorption-enhanced H2 production schematic [25].

24

Table 2.4

Optimal operating parameters for CaO sorption-enhanced studies.

Biomass

type Temperature Pressure

Steam /

biomass

Steam /

carbon

Ca /

carbon H2 content CO2 content CO content Reference

CH4: steam

mix (1:4) 480 °C 5 atm - - - 97.8% - 20 ppmv [25]

- 400-460 °C 1-5 bar - 3 - > 96 mol-% - < 7 ppmv [28]

Wood pellets 841 °C - - 0.63 - 73.9 mol-% 6.0 mol-% 6.1 mol-% [29]

- 690 °C - - - - 50.6% - - [30]

Cellulose 527-627 °C 1 atm - 1.5 0.9 83 mol-% - - [31]

Bio-oil

aqueous

fractions

500-700 °C - - 1 - 83 vol.-% - - [32]

Coal / CaO

mix 600 °C 6 MPa - - - 84.8% 1.6% 1.1% [33]

Wood 700 °C 0.6 MPa - - 2 - a - - [34]

Wet 650-700 °C - - - 0.5 51.5% inc. b 28.4% red. b - [35]

Pine bark 600 °C 1 atm - - - 48.6% inc. c - - [36]

White fir 670 °C - 0.83 - 1.5-2 ~54% 1% - [37]

Glycerol 577 °C - 9 - - 95% - - [38]

Bio-oil 600-850 °C,

800 °C d

10-20 bar, 1

atm d - 2 - > 95% - - [39]

Biogas 800 °C - - - - 55-65% - - [40]

Ethanol 500-700 °C 1 atm 4 - - - a - - [41]

a No specific maximum H2 content value given; however H2 content was maximized under the listed conditions. b No specific H2 or CO2 content values given; however H2 content was increased by 51.5% and CO2 content reduced by 28.4% after increasing moisture content

by a factor of 10. c H2 content increased by 48.6% with the use of CaO sorbent. d For the reforming and desorbing reactors, respectively. e No specific H2 or CO2 content values given; however H2 content was increased by 19% and CO2 content reduced by 50.2% when the CaO sorbent was present

in the reforming reactor.

25

Additionally, thermodynamic analysis of the sorption-enhanced reaction by Harrison and

Peng [25] using a mixture of Ni-based catalyst and Ca-based CO2 sorbent showed that H2 of >

95% purity could be produced with less than 20 ppmv CO in the product stream, on a dry basis,

given that equilibrium conditions could be closely approached. The study resulted in the

production of 97.8% H2 and 17 ppmv CO in the product gas stream. This was done under an

operating temperature of 480 °C, a pressure of 5 atm, and feed composition of 20% CH4 and 80%

steam.

Moreover, Yi and Harrison [28] investigated sorption-enhanced H2 production at lower

pressures which necessitate lower temperatures, allowing the sorbent to remain effective at CO2

removal. The results of the investigation were such that an H2 concentration > 96 mol-% and CO

concentration < 7 ppmv were obtained at reaction temperatures of 400 to 460 °C, pressures ranging

from 1 to 5 bar, and a SCR of 3. Calcined Arctic SHB dolomite was used as a source for the Ca-

based sorbent as it did not require sulphur removal prior to utilization. The H2 and CO

concentrations as a function of temperature at a SCR of 3 are depicted in Figure 2.8 [28].

Fig. 2.8. H2 and CO temperature profiles for sorption-enhanced reaction at low pressure [28].

In addition, Pfeifer et al. [29] compared H2 production capacities of steam gasification

using a dual fluidized bed system both using, and in the absence of, a calcite sorbent. The use of

the sorbent was termed the absorption enhanced reforming (AER) process. Wood pellets were used

as the test biomass in both cases. The SCRs with and without calcite sorbent were 0.63 and 0.79

kg/kg, respectively. Optimal gasifier temperatures with and without calcite sorbent were found to

26

be 841 and 645 °C, respectively, with combustion temperatures of 920 and 894 °C. Overall, it was

determined from the study that the use of calcite sorbents greatly enhanced H2 content in the

product stream, as well as significantly reducing the CO and CO2 content. H2 content with and

without the calcite sorbent were 73.9 and 37.7 mol-%, respectively; CO content with and without

were 6.1 and 29.1 mol-%; and CO2 content with and without were 6.0 and 19.6 mol-%. A downside

noted in the study was that hydrocarbon content slightly increased with the use of the sorbent.

Subsequently, the aforementioned AER process was tested on an industrial scale by

Koppatz et al. [30] using an 8 MW heat and power plant and CaO sorbent. The results were also

compared to the standard gasification process. Optimal gasification temperature was found to be

at or slightly below 690 °C. Combustion temperature for AER was roughly 850 °C and for standard

gasification was approximately 950 °C. The tested gasification temperature range for standard

gasification was 850 to 900 °C. It was observed that the AER process produced higher purity H2

(50.60%) than standard gasification while simultaneously reducing gaseous by-product content.

The differences between the AER process and standard gasification are illustrated in Figure 2.9

[30].

Florin and Harris [31] investigated the consequences of thermodynamic equilibrium on H2

production utilizing steam gasification of biomass in the form of cellulose. Analytical model

predictions were subsequently compared to experimental results. Information was also presented

which summarizes all chemical reactions that take place during the steam gasification of biomass.

This is summarized in Table 2.5 [31]. The results of the experiment showed that a maximum of 83

mol-% H2 was produced using CaO as the sorbent material. The optimal temperature range was

found to be 527 to 627 °C, a relatively lower range. The most favourable pressure conditions for

maximal H2 generation were atmospheric. Furthermore, the optimal SCR was found to be 1.5 on

a molar basis. A Ca/C ratio of 0.9 was found to be optimal for successful removal of the subsequent

concentration of CO2. Moreover, it was found that the analytical equilibrium model was neither

able to predict the formation of hydrocarbon tar by-products nor the degradation of reactivity of

the CaO sorbent.

27

Fig. 2.9. Comparison of product gas composition for AER and standard gasification [30].

Table 2.5

Chemical reactions in the steam gasification of biomass. Adapted from [31].

Reaction Chemical equation ΔHo923 (kJ mol-1)

Water-gas shift CO + H2O → CO2 + H2 -35.6 (exothermic)

Methane reforming CH4 + H2O →CO + 3 H2 225 (endothermic)

Water-gas (i) C + H2O → CO + H2 136 (endothermic)

Water-gas (ii) C + 2 H2O → CO2 + 2 H2 100 (endothermic)

Oxidation (i) C + O2 → CO2 -394 (exothermic)

Oxidation (ii) C + ½ O2 → CO -112 (exothermic)

Boudouard C + CO2 → 2CO 171 (endothermic)

Methanation C + 2 H2 → CH4 -89.0 (exothermic)

28

Yan et al. [32] studied the steam reforming of bio-oil aqueous fractions using CaO and

calcined dolomite sorbents to capture CO2. A maximum H2 content of 83 vol-% was achieved

using lower temperatures. However, maximum H2 yield was 75%, achieved at higher temperatures

in the range of 500 to 700° C. Furthermore, it was observed that H2 content was found to vary only

slightly with SCR, although it was optimized at a ratio of 1. The use of the calcined dolomite

sorbent at 600 °C and a particle diameter of 250 to 500 µm were determined to be optimal for both

H2 content and yield.

Additionally, Lin et al. [33] studied the pyrolysis of a coal/CaO mixture with steam

utilizing a flow-type reactor. The results were compared to pyrolysis of pure coal. While the

pyrolysis of pure coal resulted in product gas comprised of only about 15% H2, the mixture

generated 84.8% H2. Furthermore, the use of the mixture greatly reduced the CO2 content from

12.0% to 1.6%, and the CO content from 12.0% to a mere 1.1%. Furthermore, the small amount

of remaining CO in the product gas stream could be completely converted to H2 and CO2 via WGS.

Higher gasifier temperatures were found to favour H2 production. The coal/CaO mixture generated

twice as much H2 upon gasification as pure coal did at 700 °C. Consequently, this H2 production

value was as much as four times larger than the coal/CaO mixture performance at 600 °C.

Moreover, an increase in pressure from 1 to 6 MPa increased H2 production by a factor of about

1.5, showing that increasing pressure also favoured H2 production. However, any increase beyond

6 MPa showed minimal and relatively insignificant increases in H2 generation.

Hanaoka et al. [34] produced H2 from woody biomass while testing the effects of

temperature, pressure, and Ca/C ratio. It was observed that maximum H2 yield was obtained at a

Ca/C ratio of 2. The trend can be observed in Figure 2.10 (at 700 °C temperature, 0.6 MPa pressure,

and 10 min holding time) [34]. Optimal pressure was found to be 0.6 MPa. This value was

considerably lower when compared to other forms of biomass such as coal and heavy oil.

Furthermore, increasing temperature was favourable to H2 production, with 700 °C producing

optimal results.

Further, Guoxin and Hao [35] studied the effects of gasifying wet biomass on the

production of H2. It was found that increasing moisture content by a factor of 10 (i.e. from 0.09 to

0.90) increased H2 yield by 51.5% and reduced CO2 generation by 28.4%. Increasing temperature

also favoured H2 generation. The optimal operating temperature range was found to be 650 to 700

°C. Conversely, rising temperature tended to restrict CO2 absorption by the CaO sorbent. This

29

result was further confirmed utilizing x-ray diffraction (XRD) and scanning electron microscopy

(SEM) imaging techniques. Optimal H2 production was observed at a Ca/C ratio of roughly 0.5.

While CO2 capture was not exactly maximized at this value, only minimal increases were observed

beyond a Ca/C ratio of 0.5.

Fig. 2.10. Effect of Ca/C ratio on product gas composition [34].

Subsequently, Mahishi and Goswami [36] presented a technique to enhance H2 production

from steam gasification of pine bark biomass by integrating the gasification and absorption steps.

H2 yield was found to increase by 48.6% in the presence of a CaO sorbent, as compared with in its

absence. Moreover, overall gas yield improved by 62.2% and carbon conversion efficiency by

83.5%. These optimal conditions were produced at a gasification temperature of 600 °C and

atmospheric pressure. The use of the CaO sorbent led to reduced concentrations of CO and CH4 in

the product gas stream. CaO acted as both a sorbent and catalyst in that it aided in the reforming

of hydrocarbons and tars in the product gas stream. It is important to note that no sorbent

regeneration was conducted in this study.

Dutta et al. [37] studied the effect of varying SBR, CaO/biomass ratio, and temperature on

H2 production from the steam gasification of white fir biomass. A maximum H2 concentration of

30

about 54% was obtained at an optimal SBR of 0.83. Further, varying SBR at a CaO/biomass ratio

of 1.5 was found to produce a CO2 concentration of 1%. Increasing temperature favoured H2

production. Maximum H2 content was observed at 670°C, with a decline at higher temperatures.

However, maximum H2 yield of 315.08 mL H2 / g biomass was observed at 710 °C. A reduction

of 93.33% in CO2 concentration was observed at a CaO/biomass ratio of 2, as compared with the

no-CaO sorbent case.

Wang et al. [38] studied the effects of utilization of CaO sorbent on glycerol steam

reforming. It was found that a Ni/ZrO2 catalyst was unable to bring the system to equilibrium

conditions as maximum H2 content achieved was 64% – down from the theoretical maximum value

of 67%. A maximum H2 purity of 95% was found using the CaO sorbent with the balance being

CH4. This was conducted at 577 °C and steam/glycerol (i.e. steam/biomass) ratio of 9. The addition

of CaO was found to be highly effective in removal of CO from the product gas stream.

Furthermore, Kinoshita and Turn [39] investigated the production of transportation-fuel-

grade H2 using CaO as a CO2 sorbent via steam reforming of bio-oil. The study was simulated

using the Aspen Plus simulation software. It was found that optimal operating temperatures of the

reforming and desorbing reactors were 600 to 850 °C and roughly 800 °C, respectively. Further, a

pressure of 10 to 20 bar was utilized for the reforming reactor, while atmospheric pressure was

optimal for the desorbing reactor. At higher pressures, CO2 regeneration became difficult in the

desorbing reactor and thus lower pressures were favourable. The optimal SCR was found to be 2.

Consequently, the maximum H2 content in the product gas stream was > 95% under the

aforementioned conditions.

Assabumrungrat et al. [40] found that a mixture of CaO and Ni/SiO2·MgO was the optimal

sorbent/catalyst arrangement to maximize H2 yield for H2 production from biogas.

Thermodynamic analysis determined that thermal neutral conditions can be acquired by utilizing

higher steam/CH4 feed ratio and higher O2/CH4 or CaO/CH4 ratios. It was further observed that a

maximum H2 production value existed at an optimal steam/CH4 ratio and at 800 °C for each

CaO/CH4 ratio tested. This observation is depicted in Figure 2.11 [40].

31

Fig. 2.11. H2 concentration at H2O/CH4 ratios for various CaO/CH4 ratios [40].

Comas et al. [41] found that the utilization of a CaO sorbent greatly enhanced H2

production and reduced CO formation in the product gas stream. Ethanol steam reforming with

the use of CaO was optimal for H2 production in the temperature range of 500 to 700 °C – lower

than in the absence of the sorbent. The case utilizing CaO as sorbent also had a higher thermal

efficiency and removed the necessity for a WGS reactor. In addition to the previously mentioned

temperature range, H2 production was found to be optimized at atmospheric pressure and a

water/ethanol (i.e. steam/biomass) ratio of 4 (molar basis).

Subsequently, Mahishi et al. [42] used the Aspen Plus simulation software to simulate

ethanol sorbent-enhanced gasification for H2 production. Three cases were considered: no CaO

sorbent, reforming with CaO sorbent in the reformer, and reforming with CaO sorbent in the WGS

reactor. The optimal case was determined to be that of CaO sorbent in the reformer, with an optimal

gasification temperature of 777 °C and optimal pressure of 1 atm. The optimal SBR was

determined to be 4. Furthermore, it was observed that the sorbent utilization increased the

equilibrium H2 yield from conventional gasification by roughly 19% while simultaneously

reducing CO2 content by 50.2%.

32

2.6 Research Gaps

Future research proposed by Kinoshita and Turn [39] focuses on ascertaining a better

understanding of the kinetic limitations of the involved reactions, as well as further identification

of optimal operating parameters and assessment of CaO sorbent regeneration and compatibility

with any reforming catalysts that may be utilized. The solid form of the utilized sorbent (i.e.

powder or pellet) also leaves room for any potential research, as mentioned by Mahishi et al. [42].

Moreover, further research may be conducted on the kinetic mechanisms of the reaction

between the CaO sorbent and CO2 during the latter’s absorption and desorption processes.

Maximizing the rate of this equilibrium reaction would be the aim of such research. It is suggested

that the tar-reforming capabilities of the CaO sorbent be further addressed to enhance in situ tar

removal from the product gas stream [42,43].

In addition, Florin and Harris [43] suggested that more research is required in determining

the effect of biomass chemical composition (i.e. ultimate analysis) on H2 production. It was also

proposed that there is merit in research regarding the modification of the crystal structure of CaO

sorbent. Ion addition to enhance the CO2 capture capabilities of CaO is exemplar of such

adaptations. Subsequently, the effect of impurities in the CaO sorbent bed on H2 production should

also be considered.

2.7 Concluding Remarks

In conclusion:

1. H2 presents a clean, efficient alternative energy production source to fossil fuels and its

utilization has the potential to become widespread in the near future. The thermochemical

conversion of biomass is a viable H2 production technique capable of achieving high-purity

H2 syngas streams.

2. Biomass gasification in conjunction with the use of Ni- or other metal-based catalysts

resulted in H2 production purities upwards of 70 mol-%. Increasing gasification

temperature favoured H2 production, with an average optimal temperature range between

750 and 850 °C. Maximal H2 production occurred at steam/biomass ratios of 0.6 to 0.7 with

residence times varying from 10 to 20 min.

33

3. Following biomass gasification with the steam reforming method further enhanced H2

production and purified the product gas stream. H2 production using this method reached

upwards of 80 to 90%. Increasing temperature favoured H2 production and optimal

temperatures ranged from 627 to 750 °C. Steam/carbon ratios varied as some studies found

lower ranges of 1.7 to 2.25 to be optimal, while others found higher values of 9 to be

optimal. Optimal pressures were atmospheric or roughly so.

4. Chemical looping gasification (CLG) combines H2 production with CO2 capture and

purification, and can be achieved using either an O2 or CO2 carrier. Fe2O3 was found to be

a suitable O2 carrier, although reduced efficiency was an issue reported in some studies

when considering increased reduction-oxidation cycles. CaO was found to be an effective

CO2 carrier, permitting superior CO2 separation and downstream concentration.

5. Enhancing H2 production with the utilization of Ca-based sorbent for CO2 capture greatly

enhanced the purity of the product gas stream. Purities > 95 mol-% H2 were observed in

many studies, with an average upwards of 80%. CO2 and CO content in the gas stream

were also greatly reduced. Higher reaction temperatures favoured H2 production with an

optimal range of 600 to 700 °C. Atmospheric pressures were favourable. Optimal

steam/carbon and Ca/carbon ratios of 1.5 to 2 were characteristic.

6. Finally, potential research opportunities are rife. Major focuses include acquiring a better

understanding of the kinetic mechanisms and limitations of the reactions involved in

thermochemical biomass conversion processes as well as further identification of optimal

operating parameters for the maximization of H2 production. More research is required

regarding CaO sorbent performance, regeneration, and tar-reforming capabilities. The

effect of the chemical composition of biomass feed on H2 production and purity should

also be considered in future work.

34

References

[1] Levin, David B., and Richard Chahine. "Challenges for Renewable Hydrogen Production from

Biomass." International Journal of Hydrogen Energy 35 (2009): 4962-969.

[2] Holladay, J., J. Hu, D. King, and Y. Wang. "An Overview of Hydrogen Production Technologies." Catalysis

Today 139.4 (2009): 244-60.

[3] Hamelinck, Carlo N., and Andre P.C. Faaij. "Future Prospects for Production of Methanol and Hydrogen

from Biomass." Journal of Power Sources 111 (2002): 1-22.

[4] Tanksale, Akshat, Jorge Norberto Beltramini, and Gao Qing Max Lu. "A Review of Catalytic Hydrogen

Production Processes from Biomass." Renewable and Sustainable Energy Reviews 14 (2010): 166-82.

[5] Ni, M., D. Leung, M. Leung, and K. Sumathy. "An Overview of Hydrogen Production from Biomass." Fuel

Processing Technology 87.5 (2006): 461-72.

[6] Saxena, R., D. Seal, S. Kumar, and H. Goyal. "Thermo-chemical Routes for Hydrogen Rich Gas from

Biomass: A Review." Renewable and Sustainable Energy Reviews 12.7 (2008): 1909-927.

[7] Kirtay, Elif. "Recent Advances in Production of Hydrogen from Biomass." Energy Conversion and

Management 52 (2011): 1778-789.

[8] Kalinci, Yildiz, Arif Hepbasli, and Ibrahim Dincer. "Biomass-based Hydrogen Production: A Review and

Analysis." International Journal of Hydrogen Energy 34.21 (2009): 8799-817.

[9] Guo, Y., S.Z. Wang, D.H. Xu, Y.M. Gong, H.H. Ma, and X.Y. Tang. "Review of Catalytic Supercritical

Water Gasification for Hydrogen Production from Biomass." Renewable and Sustainable Energy Reviews 14

(2010): 334-43.

[10] Moghtaderi, B. "Effects of Controlling Parameters on Production of Hydrogen by Catalytic Steam

Gasification of Biomass at Low Temperatures." Fuel 86.15 (2007): 2422-430.

[11] Abuadala, A., I. Dincer, and G.F. Naterer. "Exergy Analysis of Hydrogen Production from Biomass

Gasification." International Journal of Hydrogen Energy 35 (2010): 4981-990.

[12] Guo, L., Y. Lu, X. Zhang, C. Ji, Y. Guan, and A. Pei. "Hydrogen Production by Biomass Gasification in

Supercritical Water: A Systematic Experimental and Analytical Study."Catalysis Today 129.3-4 (2007): 275-

86.

[13] Furusawa, T., T. Sato, H. Sugito, Y. Miura, Y. Ishiyama, M. Sato, N. Itoh, and N. Suzuki. "Hydrogen

Production from the Gasification of Lignin with Nickel Catalysts in Supercritical Water." International

Journal of Hydrogen Energy 32.6 (2007): 699-704.

[14] Gonzalez, J., S. Roman, D. Bragado, and M. Calderon. "Investigation on the Reactions Influencing Biomass

Air and Air/steam Gasification for Hydrogen Production." Fuel Processing Technology 89.8 (2008): 764-72.

[15] Shen, L., Y. Gao, and J. Xiao. "Simulation of Hydrogen Production from Biomass Gasification in

Interconnected Fluidized Beds." Biomass and Bioenergy 32.2 (2008): 120-27.

35

[16] Franco, C., F. Pinto, I. Gulyurtlu, and I. Cabrita. "The Study of Reactions Influencing the Biomass Steam

Gasification Process⋆." Fuel 82.7 (2003): 835-42.

[17] Wang, D., S. Czernik, D. Montané, M. Mann, and E. Chornet. "Biomass to Hydrogen via Fast Pyrolysis and

Catalytic Steam Reforming of the Pyrolysis Oil or Its Fractions."Industrial & Engineering Chemistry

Research 36.5 (1997): 1507-518.

[18] Chen, Guanyi, and Lixia Zhao. "Preliminary Investigation on Hydrogen-rich Gas Production by Co-steam-

reforming of Biomass and Crude Glycerin." International Journal of Hydrogen Energy 37 (2012): 765-73.

[19] Vagia, E., and A. Lemonidou. "Thermodynamic Analysis of Hydrogen Production via Steam Reforming of

Selected Components of Aqueous Bio-oil Fraction." International Journal of Hydrogen Energy 32.2 (2007):

212-23.

[20] Scott, S. A., Dennis, J. S., Hayhurst, A. N., and Brown, T., “In Situ Gasification of a Solid Fuel and CO2

Separation Using Chemical Looping”, AIChE Journal 52 (2006): 3325-328.

[21] Fan, Liang-Shih. Chemical Looping Systems for Fossil Energy Conversions. Hoboken, NJ; Wiley-AIChE,

2010. Print.

[22] Leion, H., Mattisson, T. and Lyngfelt, A., “Solid Fuels in Chemical-Looping Combustion”, International

Journal of Greenhouse Gas Control 2 (2008): 180-93.

[23] Harrison, Douglas P. "Sorption-Enhanced Hydrogen Production: A Review." Industrial & Engineering

Chemistry Research 47.17 (2008): 6486-501.

[24] Satrio, Justinus A., Brent H. Shanks, and Thomas D. Wheelock. "A Combined Catalyst and Sorbent for

Enhancing Hydrogen Production from Coal or Biomass." Energy & Fuels 21.1 (2007): 322-26.

[25] Harrison, Douglas P., and Zhiyong Peng. "Low-Carbon Monoxide Hydrogen by Sorption-Enhanced

Reaction." International Journal of Chemical Reactor Engineering 1.1 (2003): 1-9.

[26] Grasa, Gemma S., and J. Carlos Abanades. "CO2 Capture Capacity of CaO in Long Series of

Carbonation/Calcination Cycles." Industrial & Engineering Chemistry Research 45 (2006): 8846-851.

[27] Lopez Ortiz, Alejandro, and Douglas P. Harrison. "Hydrogen Production Using Sorption-Enhanced

Reaction." Industrial & Engineering Chemistry Research 40.23 (2001): 5102-109.

[28] Yi, Kwang Bok, and Douglas P. Harrison. "Low-Pressure Sorption-Enhanced Hydrogen

Production." Industrial & Engineering Chemistry Research 44.6 (2005): 1665-669.

[29] Pfeifer, C., B. Puchner, and H. Hofbauer. "Comparison of Dual Fluidized Bed Steam Gasification of Biomass

with and without Selective Transport of CO2." Chemical Engineering Science 64.23 (2009): 5073-083.

[30] Koppatz, Stefan, Christoph Pfeifer, Reinhard Rauch, Hermann Hofbauer, Tonja Marquard-Moellenstedt, and

Michael Specht. "H2 Rich Product Gas by Steam Gasification of Biomass with in Situ CO2 Absorption in a

Dual Fluidized Bed System of 8 MW Fuel Input." Fuel Processing Technology 90.7-8 (2009): 914-21.

[31] Florin, N., and A. Harris. "Hydrogen Production from Biomass Coupled with Carbon Dioxide Capture: The

Implications of Thermodynamic Equilibrium." International Journal of Hydrogen Energy 32.17 (2007):

4119-134.

36

[32] Yan, Chang-Feng, En-Yuan Hu, and Chi-Liu Cai. "Hydrogen Production from Bio-oil Aqueous Fraction with

in Situ Carbon Dioxide Capture." International Journal of Hydrogen Energy 35 (2010): 2612-616.

[33] Lin, Shiying, Michiaki Harada, Yoshizo Suzuki, and Hiroyuki Hatano. "Hydrogen Production from Coal by

Separating Carbon Dioxide during Gasification." Fuel 81 (2002): 2079-085.

[34] Hanaoka, T., Takahiro Yoshida, Shinji Fujimoto, Kenji Kamei, Michiaki Harada, Yoshizo Suzuki, Hiroyuki

Hatano, Shin-ya Yokoyama, and Tomoaki Minowa. "Hydrogen Production from Woody Biomass by Steam

Gasification Using a CO2 Sorbent."Biomass and Bioenergy 28.1 (2005): 63-68.

[35] Guoxin, Hu, and Huang Hao. "Hydrogen Rich Fuel Gas Production by Gasification of Wet Biomass Using a

CO2 Sorbent." Biomass and Bioenergy 33.5 (2009): 899-906.

[36] Mahishi, M., and D. Goswami. "An Experimental Study of Hydrogen Production by Gasification of Biomass

in the Presence of a CO2 Sorbent." International Journal of Hydrogen Energy 32.14 (2007): 2803-808.

[37] Acharya, Bishnu, Animesh Dutta, and Prabir Basu. "An Investigation into Steam Gasification of Biomass for

Hydrogen Enriched Gas Production in Presence of CaO."International Journal of Hydrogen Energy 35.4

(2010): 1582-589.

[38] Wang, Xiaodong, Maoshuai Li, Shuirong Li, Hao Wang, Shengping Wang, and Xinbin Ma. "Hydrogen

Production by Glycerol Steam Reforming With/without Calcium Oxide Sorbent: A Comparative Study of

Thermodynamic and Experimental Work." Fuel Processing Technology 91.12 (2010): 1812-818.

[39] Kinoshita, C. M., and S. Q. Turn. "Production of Hydrogen from Bio-oil Using CaO as a CO2

Sorbent." International Journal of Hydrogen Energy 28 (2003): 1065-071.

[40] Assabumrungrat, S., P. Sonthisanga, W. Kiatkittipong, N. Laosiripojana, A. Arpornwichanop, A.

Soottitantawat, W. Wiyaratn, and P. Praserthdam. "Thermodynamic Analysis of Calcium Oxide Assisted

Hydrogen Production from Biogas." Journal of Industrial and Engineering Chemistry 16.5 (2010): 785-89.

[41] Comas, Jose, Miguel Laborde, and Norma Amadeo. "Thermodynamic Analysis of Hydrogen Production

from Ethanol Using CaO as a CO2 Sorbent." Journal of Power Sources 138.1-2 (2004): 61-67.

[42] Mahishi, Madhukar R., M. S. Sadrameli, Sanjay Vijayaraghavan, and D. Y. Goswami. "A Novel Approach

to Enhance the Hydrogen Yield of Biomass Gasification Using CO2 Sorbent." Journal of Engineering for

Gas Turbines and Power 130 (2008): 011501-1 to 11501-8. Print.

[43] Florin, N., and A. Harris. "Enhanced Hydrogen Production from Biomass with in Situ Carbon Dioxide

Capture Using Calcium Oxide Sorbents." Chemical Engineering Science 63.2 (2008): 287-316.

37

Chapter III

Solid Oxide Fuel Cell Literature Review – A Review of the

Effect of Fuel Type and Composition on Solid Oxide Fuel Cell

Performance


Solid oxide fuel cells (SOFCs) are electrochemical devices which directly convert chemical energy

into electrical energy at high temperatures (typically 600 to 1000 °C). Similar to other fuel cell

types, SOFCs consist of an anode, electrolyte, cathode, and interconnect. However, due to the

higher operating temperatures, no liquid components are utilized and all portions of the cell

assembly are solid-state. The anode material is usually a nickel-yttria-stabilized zirconia (Ni-YSZ)

cermet (ceramic-metal), the electrolyte is YSZ ceramic, and the cathode material usually consists

of strontium (Sr)-doped lanthanum manganate (LaMnO3). Moreover, due to the high operating

temperature, SOFCs can operate on a multitude of fuel types and compositions. SOFC operation

on hydrogen (H2) is illustrated in Figure 3.1. The electrochemical reactions and consequent overall

reactions for H2, carbon monoxide (CO), and methane (CH4) fuels are displayed in Table 3.1.

Fig. 3.1. SOFC operation on H2 fuel.

38

Table 3.1

Chemical reactions for SOFC operation on various fuels.

Fuel Anodic Oxidation Cathodic Reduction Overall Reaction a

H2 H2 + O= → H2O + 2e- ½ O2 + 2e- → O= H2 + ½ O2 → H2O + We + Q

CO CO + O= → CO2 + 2e- ½ O2 + 2e- → O= CO + ½ O2 → CO2 + We + Q

CH4 CH4 + 4 O= → CO2 + 2 H2O + 8e- ½ O2 + 2e- → O= CH4 + 2 O2 → CO2 + 2 H2O + We + Q a We = electrical work [J mol-1] and Q = thermal energy [J mol-1].

Oxygen (O2) reduction at the cathode occurs independent of inlet fuel type, and thus the

electrochemical reaction is the same in each case. The resultant oxide ions (O=) migrate across the

YSZ electrolyte to react with the inlet fuel at the Ni-YSZ anode, hence initiating electrochemical

anodic oxidation of the fuel. Overall, the system involves the reaction of the fuel with O2, the

production of desired electrical work (i.e. electricity), and the generation of heat, H2O, and/or CO2

by-products, depending on the carbon content of the fuel. Specifically, the generation of CO2

during SOFC operation becomes unavoidable with the use of carbonaceous fuels like CO, CH4,

and higher-order hydrocarbons. Hydrocarbon operation will be further discussed in the following

sections of this review.

Table 3.1 suggests that direct oxidation of CO and CH4 is possible for SOFC operation.

However, another mechanism of utilization exists for each of the two fuels. Given that the SOFC

operates on CO, the water-gas shift reaction may proceed when the gas is humidified. This reaction

is illustrated by the following:

Water-gas shift: CO + H2O → H2 + CO2

Following the water-gas shift reaction, the product H2 continues on to be oxidized at the anode in

the same manner as if the SOFC were operating on H2. Furthermore, SOFC operation on CH4 may

proceed via the steam reforming of methane reaction, shown by the following:

Steam reforming: CH4 + H2O → CO + 3 H2

The resulting H2 undergoes anodic oxidation while the generated CO undergoes the water-gas shift

reaction to produce more H2.

Subsequently, the major advantages and disadvantages of SOFCs stem mainly from the

high operating temperatures and overall geometric design of the cell. Further, some disadvantages

arise due to the electrochemistry of the involved fuels, electrolyte, electrodes, and other

components of the cell assembly. For example, the main advantages of SOFC use include [1,2]:

39

High versatility – not limited to one type of fuel,

No precious metal catalyst requirements for operation (e.g. platinum),

Faster electrochemical kinetics due to high operating temperatures,

Solid-state electrolyte – allows for both planar and tubular geometric designs,

High efficiency and long life time.

Conversely, some disadvantages for SOFCs include [1,2]:

Difficult to find materials that can withstand such high temperatures for extended periods

of time,

Constant reduction-oxidation (redox) cycling implies that nickel is constantly being

converted from Ni to NiO – leads to volume changes and potential cracking of electrodes,

Graphite formation by the Ni-catalyzed cracking of hydrocarbons leads to carbon

deposition at the anode and ultimate reduction in cell performance,

Susceptible to sulfur poisoning.

Although many challenges are presented by the listed disadvantages, extensive research is

being conducted in an attempt to address the concerns. For example, much research focuses on

SOFC operating temperature reduction while avoiding the necessity for precious metals to catalyze

anodic fuel oxidation. Also, materials research is being undertaken to find alternative materials

which can withstand the high temperatures. Many studies also look at volume changes during

operation upon redox cycling and the subsequent effect on the structural integrity of the electrodes.

Finally, the subject of this report, much research has been conducted on the fuel flexibility of

SOFCs and the effects of fuel type and composition on overall cell performance.

3.2 Literature Review

Extensive research has been conducted and presented in the literature regarding SOFC

performance with the utilization of fuels other than H2. These include CO, CH4, and other

hydrocarbon fuels. Furthermore, experiments have been conducted using various fuel mixtures

including coal syngas, biomass syngas, and hydrocarbon mixtures.

The purpose of this review is to determine the effect of fuel type or composition on SOFC

performance. Additionally, the performance characteristics for the various fuel types and

compositions are compared and contrasted so as to highlight the optimal type and/or advantages

40

and disadvantages of the utilization of each. Performance characteristics under consideration

include voltage and power output for a given range of input current densities and cell lifetime over

an extended time period. Therefore, the following is a review of the aforementioned works and

research activities and summary of the main results and conclusions.

3.2.1 Performance Comparison of H2 and Hydrocarbons

Hydrocarbons present a plausible fuel for SOFCs since high temperature operation allows for their

direct oxidation in the cell. Further, Ni acts as an electrocatalyst for the direct oxidation of

hydrocarbons thereby enhancing their performance and potential usability in SOFCs.

For instance, Barnett et al. [2] operated an anode-supported Ni-YSZ SOFC on CH4 and

compared the results to that of H2 under the same conditions for various temperatures. The voltage

and power characteristic output curves obtained can be seen in Figure 3.2 below. The maximum

power output was greater for H2 at 1.44 W cm-2 than for CH4 at 1.27 W cm-2. Furthermore, higher

temperatures favoured greater performance in both cases. CH4, however, did produce higher open

circuit voltages (OCV) than H2 at the higher temperature measurements, especially that of 800 °C.

It was suggested that the slightly poorer performance of CH4 could be attributed to coking and

carbon deposition on the Ni-YSZ anode during cell operation.

Fig. 3.2. Voltage and power output curve comparison for (a) H2 and (b) CH4 [2].

Subsequently, Eguchi et al. [3] also studied the effect of CH4 fuel on SOFC performance.

However, mixtures of CH4, steam, and N2 in varying proportions were utilized as test fuel, and the

output voltages measured and recorded. The results can be seen in Figure 3.3. The results showed

41

that increasing CH4 fuel dilution with N2 led to reduced performance at all current densities. The

OCV also decreased with increasing N2 content. Furthermore, ethane (C2H6) and ethene (C2H4)

were also utilized as fuels and compared with CH4 [3]. The results can be seen in Figure 3.4. The

voltage characteristic curves for both ethane and ethene decrease more slowly with increasing

current density than CH4 shown in Figure 3.3. However, the increased carbon content of the former

two fuels accelerated carbon deposition in the anode and accordingly reduced the overall life of

the cell [3].

Fig. 3.3. Voltage output curves for various CH4-H2O-N2 mixtures [3].

Fig. 3.4. Voltage output curves for (a) ethane and (b) ethene [3].

42

In addition to CH4 and ethane, higher-order hydrocarbons such as propane (C3H8) and n-

butane (C4H10) were compared to H2 by Barnett and Madsen [4], using a La0.8Sr0.2Cr0.98V0.02O3

(LSCV) anode also containing Ce0.9Gd0.1O1.98 (gadolinium-doped ceria, or GDC) and 5 wt.-% NiO

for catalytic hydrocarbon oxidation. The SOFC performance using propane can be seen in contrast

to H2 in Figure 3.5. The performance for both fuels improved with increasing temperature.

However, H2 outperformed propane at every temperature and for all current densities tested.

Additionally, H2 achieved a maximum power output of 140 mW cm-2 at 750 °C while propane

reached the slightly lower value of 130 mW cm-2. Again, the weaker performance of the

hydrocarbon may be attributed to carbon deposition issues which ultimately reduce the

performance of the cell [4]. A summary of the results obtained by Barnett and Madsen for direct

utilization of hydrocarbon fuels can be seen in Table 3.2 below, including those for n-butane. The

table shows that H2 has the highest power output with the exception of propane, which is higher

for direct utilization of the fuels in the SOFC cell. Furthermore, all OCVs were roughly 800 mV,

and n-butane had the lowest maximum power density.

Table 3.2

Summary of performance data at 700 °C for H2 and various hydrocarbons. Adapted from [4].

Fuel Type Open circuit voltage

(mV)

Maximum current

density (mA cm-2)

Maximum power

density (mW cm-2)

Elapsed time in test

(h)

H2 785 508 114 362

CH4 770 490 107 266

C2H6 763 488 105 265

C3H8 801 430 128 195

C4H10 812 338 85 408

The performance characteristics of n-butane were further explored by Park et al. [5] using

a copper-ceria (Cu-CeO2) anode, as opposed to the conventional Ni-YSZ anodes. Unlike Ni, Cu

does not catalyze hydrocarbon oxidation and therefore avoids the issue of graphite formation and

subsequent carbon deposition in the anode. Regardless of this fact, the performance of n-butane

still did not match up to that of H2, as illustrated in Figure 3.6. In the figure, filled in shapes

represent data at 700 °C and open shapes correspond to data at 800 °C. Triangles represent H2 and

circles represent n-butane. H2 outperformed the hydrocarbon in terms of voltage and power output

at all tested current densities and for both tested temperatures. At 800 °C, H2 reached a maximum

power output of 0.31 W cm-2, whereas n-butane was less than two-thirds of that value at 0.18 W

cm-2. In fact, maximum power output for H2 at 700 °C was greater than that of n-butane at 800 °C.

43

Fig. 3.5. Voltage output curves for H2 and propane at different temperatures [4].

Fig. 3.6. Voltage and power output curve comparison for H2 and n-butane [5].

44

3.2.2 Performance Comparison of H2 and CO

As previously mentioned, CO is another potential SOFC fuel. SOFCs do not contain

platinum (Pt) catalysts as a result of the higher operating temperatures and consequently faster

electrochemical kinetics. Thus, Pt catalyst poisoning by CO, a major issue in polymer electrolyte

membrane fuel cells (PEMFCs), is not an issue in SOFCs. For example, Eguchi et al. [3] compared

the voltage outputs of H2-H2O and CO-CO2 binary systems at 1000 °C. H2O and CO2 were present

in fixed compositions of 0.6 and 0.5%, respectively. The study was conducted with increasing

amounts of N2 diluent, as illustrated in Figure 3.9. No major decrease in performance was observed

for either H2 or CO when dropping their respective compositions to 66.6 and 77.1% suggesting

that some fuel dilution will not significantly affect voltage output. This phenomenon has

implications in fuel costs savings, since the utilization of less fuel leads to a similar output.

Furthermore, H2 outperformed CO at all tested dilution levels at the test temperature of 1000 °C.

This occurred in spite of similar OCVs in both cases [3].

However, in contrast to the findings of Eguchi et al., Jiang and Virkar [6] determined that

dilution of H2 fuel does, in fact, significantly reduce the voltage and power output of the cell, even

at smaller dilution levels. The findings are illustrated in Figure 3.10. It is clear that even at the 15%

N2 level, the performance of the SOFC was significantly reduced when compared to pure H2. This

effect remained for all dilution levels tested. Additionally, the same effect was noticed for voltage

and power output when using helium (He), H2O, and CO2 as diluents at the same temperature [6].

Diluting CO with CO2 and operating in the same SOFC at the same temperature showed similar

results. These can be seen in Figure 3.11. Cell performance greatly reduced with each dilution

level of CO2. This is especially seen in terms of the voltage characteristic curves. Further, it is

important to note that the maximum power achieved by pure CO, roughly 0.7 W cm-2, was much

less than that of pure H2, approximately 1.8 W cm-2. The weaker performance of CO may be

attributed to greater concentration polarization at the anode and slower oxidation of CO than H2

[6].

45

Fig. 3.7. Voltage output curves comparing H2-H2O and CO-CO2 systems at 1000 °C with N2 as a diluent [3].

Fig. 3.8. Effect of adding N2 diluent to H2 fuel on voltage and power output at 800 °C [6].

46

Fig. 3.9. Effect of adding CO2 diluent to CO fuel on voltage and power output at 800 °C [6].

Finally, Jiang and Virkar [6] also tested mixtures of H2 and CO under the same

aforementioned conditions. The results are summarized in Figure 3.12 below. It can be seen that,

generally, increasing CO composition hindered cell performance in terms of both voltage and

power output, with the exception of the 14% CO-86% H2 mixture. This mixture outperformed pure

H2 in terms of power output. Thus, it was found that performance remained high given that H2

concentration remained above the approximate value of 50%, as the balance CO generates more

H2 via the water-gas shift reaction [6].

Moreover, H2-CO binary systems provide different results depending on the anode used in

the SOFC. Vohs et al. [7] compared H2 and CO performance in SOFCs for Cu-CeO2-YSZ and Ni-

YSZ anode-supported cells. Figure 3.13 demonstrates the performance characteristics of each

anode type at 700 °C. The squares represent H2, circles represent CO, and diamond shapes

represent syngas. It is evident that the OCVs of all three fuel types are higher for the Cu-CeO2-

YSZ anode. Further, the slope of the voltage curve for CO is far less steep for that anode as well.

In fact, it is very similar to that of H2, which is evidently not the case for the Ni-YSZ anode. Thus,

CO performance becomes closer to that of H2 with the use of a Cu-CeO2-YSZ anode. This is

because Ni does not catalyze CO oxidation. This effect was also readily apparent when H2 and CO

were mixed with n-butane and tested under the same conditions. 90% mixtures were made of each

47

– with the balance being n-butane – and the resulting voltage characteristic curves were only

slightly less than that of pure H2 or CO. Further, the slopes of the curves were all very similar [7].

Fig. 3.10. Voltage and power output for various H2-CO mixtures [6].

Fig. 3.11. Comparison of H2, CO, and syngas performance with (A) Cu-CeO2-YSZ and (B) Ni-YSZ anodes [7].

Correspondingly, Wang et al. [8] found that a Cu-CeO2 coating over the Ni-YSZ anode

under H2-CO conditions greatly improved cell performance. The optimal mixture composition was

found to be 65% H2-32% CO-3% H2O. The cell operating on this mixture at 750 °C ran for 1050

48

hours without any notable decrease in performance or materials degradation. Figure 3.14

summarizes the obtained data. Increasing CO content again decreased cell performance, in

accordance with the other data presented in this section.

Fig. 3.12. Effect of increasing CO content on voltage and power output using Cu-CeO2-coated Ni-YSZ anode [8].

3.2.3 Performance Comparison of H2 and Syngas

Syngas derived from coal or biomass sources can be used in SOFCs due to their relatively

high concentrations of H2, CO, CH4, and various other hydrocarbons. For example, Figure 3.8 of

Section 3.2.2 compares the voltage and power outputs for a 25 kW SOFC system operating on

biogas and coal syngas. It can be seen that biogas produces favourable performance when

compared to coal syngas [9].

Furthermore, Figure 3.13 compares syngas data to H2 and CO at 700 °C. It can be seen that

syngas underperforms when compared to H2 and CO when the Cu-CeO2-YSZ anode is utilized,

but greatly outperforms CO when using the standard Ni-YSZ anode. This result is attributed to the

fact that Ni does not catalyze the electrochemical oxidation of CO; however it will make use of

the H2 present in syngas [7].

Moreover, extensive research has been conducted in regards to the direct utilization of the

products of biomass gasification in SOFCs. Such a study was conducted by Panapoulos et al. [10]

49

wherein a 100 kWth circulating fluidized bed gasifier was used to produce heavy tar (i.e.

hydrocarbon) loadings (> 10 g Nm-3) and fed to a SOFC using a Ni-GDC (gadolinium-doped ceria)

anode. No carbon deposition was found upon inspection of the Ni-GDC anode and cell voltages

upwards of 787 mV were produced at a current density of 130 mA cm-2. This value corresponded

to a cell voltage of 811 mV under virtually tar-free conditions, resulting in a drop in performance

of less than 3%. It must also be noted that the sulfur content of the fuel, in the form of hydrogen

sulfide (H2S), was controlled via feed stream pre-treatment [10].

3.2.4 Performance in the Presence of Sulfur

The presence of sulfur in the fuel feed stream to a SOFC can poison the Ni catalyst and

significantly reduce cell voltage and power output. Sulfur blocks reaction sites on the Ni catalyst

which is thereby unable to catalyze oxidation reactions, significantly increasing concentration

polarization and ultimately worsening cell performance. Figure 3.15 [11] summarizes the effects

of the presence of sulfur at various concentrations at 1000 °C and a current density of 0.3 A cm-2.

Fig. 3.13. Concentration polarization effects from sulfur poisoning after (a) addition and (b) removal [11].

As such, Matsuzaki and Yasuda demonstrated that the greater the concentration of sulfur,

the higher the overvoltage, or voltage consumed due to the aforementioned concentration

polarization effects. Further, it is important to note that sulfur poisoning effects were reversed upon

the removal of sulfur, regardless of the concentration, and overvoltage due to concentration

polarization returned to normal levels within approximately 50 min [11].

50

3.3 Research Impact

The ability to predict the performance of a SOFC based on the fuel fed to it has implications

for the entire SOFC industry and any industries or academic settings that wish to employ SOFCs

as a major part of fulfilling their energy requirements. Different fuel types will be available

depending on where the fuel is being purchased and/or acquired from, and thus the utilized SOFC

can be tailored to suit that fuel type/composition. For example, if fuel is being acquired from

gasified biomass sources, knowledge of the performance of H2-CO, CO-CO2, CH4, and other

syngas constituent mixtures would allow for optimal SOFC selection. Accordingly, different

SOFC types and designs would be used for different fuel types. Subsequently, knowledge of the

energy requirements for the given application would aid in the designing and implementing of an

appropriate SOFC. Thus, prior knowledge of the fuel type and amounts of power that fuel can

generate would further help the process of selection.

Increased knowledge of the effects fuel type and composition play on SOFC performance

could lead to a catalogue or guidebook of commercial SOFCs that are ready to be ordered. The

consumer, whether private or public, could simply choose their desired SOFC based on their

available fuel type and energy requirements. Such a system would eliminate complications

inherent to designing SOFCs on a case-by-case basis and greatly contribute to the use of renewable

technologies. Furthermore, adapting SOFCs to accommodate for as-received fuel types would

eliminate costly and energy-intensive process steps such as fuel gas cleaning and pre-treatment.

The SOFC would then simply use the “contaminant” gases as fuel.

Furthering knowledge in this research area is likely to continue in both the near and long-

term future. Extensive research is already conducted regarding the performance of SOFCs under

various fuel types, as explained in previous sections of this review. Continuing investment into

research and development activities in this research area are required for further breakthroughs

and for further understanding of the effect of more fuel types and compositions under varying

operating conditions such as temperatures, pressures, and current densities. Creating interest

amongst researchers, both younger and older, is also of paramount importance in regards to

ensuring more advances in this research area in the near future.

51


In conclusion:

1. Increasing temperature favoured an increase in performance for all fuel types and

compositions.

2. H2 outperformed CH4 both in terms of voltage and power output.

3. H2 outperformed ethane (C2H6), ethene (C2H4), propane (C3H8), and n-butane (C4H10).

Higher-order hydrocarbons showed better performance than lower-order ones due to a

greater number of electrons released upon oxidation; however the issue of graphitic carbon

deposition at the anode became worse.

4. H2 outperformed CO under most conditions tested.

a. Increasing CO content in H2-CO mixtures generally reduced overall performance.

b. Dilution of H2 with N2, He, or H2O and CO with CO2 reduced cell performance but

presented a trade-off between fuel cost savings and performance.

5. Ni-GDC (gadolinium-doped ceria) anodes could effectively handle high tar loadings

present in fuel streams from biomass gasification. Sulfur (present in the form of H2S)

removal was required as pre-treatment.

6. The presence of sulfur greatly contributed to concentration polarization. Upwards of 290

mV of overvoltage were consumed under the presence of 15 ppm H2S at 1000 °C and 0.3

A cm-2 current density.

7. Knowledge of SOFC performance based on fuel input would allow for appropriate

selection of SOFC designs to fit given energy requirements and on-hand fuel types and

compositions.

52

References

[1] Minh, N.Q. “Solid Oxide Fuel Cell Technology – Features and Applications.” Solid State Ionics 174 (2004):

271-77.

[2] Barnett, S., Lin, Y., Z. Zhan, and J. Liu. "Direct Operation of Solid Oxide Fuel Cells with Methane

Fuel." Solid State Ionics 176.23-24 (2005): 1827-835.

[3] Eguchi, K., H. Kobo, T. Takeguchi, R. Kikuchi, and K. Sasaki. "Fuel Flexibility in Power Generation by

Solid Oxide Fuel Cells." Solid State Ionics 152-153 (2002): 411-16.

[4] Barnett, S., and Madsen, B. "Effect of Fuel Composition on the Performance of Ceramic-based Solid Oxide

Fuel Cell Anodes." Solid State Ionics 176.35-36 (2005): 2545-553.

[5] Park, S., J.M. Vohs, and R.J. Gorte. “Direct Oxidation of Hydrocarbons in a Solid-Oxide Fuel Cell.” Nature

404 (2000): 265-67.

[6] Jiang, Yi, and Anil V. Virkar. "Fuel Composition and Diluent Effect on Gas Transport and Performance of

Anode-Supported SOFCs." Journal of the Electrochemical Society 150.7 (2003): A942-951.

[7] Vohs, J.M., O. Costa-Nunes, and R.J. Gorte. “Comparison of the Performance of Cu-CeO2-YSZ and Ni-

YSZ Composite SOFC Anodes with H2, CO, and Syngas.” Journal of Power Sources 141 (2005): 241-49.

[8] Wang, S.R., Xiao-Feng Ye, J. Zhou, F.R. Zeng, H.W. Nie, and T.L. Wen. "Assessment of the Performance

of Ni-yttria-stabilized Zirconia Anodes in Anode-supported Solid Oxide Fuel Cells Operating on H2–CO

Syngas Fuels." Journal of Power Sources 195.21 (2010): 7264-267.

[9] Samuelsen, G., Yi, Y., A. Rao, and J. Brouwer. "Fuel Flexibility Study of an Integrated 25kW SOFC

Reformer System." Journal of Power Sources 144.1 (2005): 67-76.

[10] Panopoulos, K.D., Hofmann, Ph., P.V. Aravind, M. Siedlecki, A. Schweiger, J. Karl, J.P. Ouweltjes, and

E. Kakaras. "Operation of Solid Oxide Fuel Cell on Biomass Product Gas with Tar Levels 10 g

Nm−3." International Journal of Hydrogen Energy 34 (2009): 9203-212.

[11] Matsuzaki, Yoshio, and Isamu Yasuda. “The poisoning effect of sulfur-containing impurity gas on a SOFC

anode: Part I. Dependence on temperature, time and impurity concentration.” Solid State Ionics 132 (2000):

261-69.

53

Chapter IV

Chemical Looping Gasification for Hydrogen Production – A

Comparison of Two Unique Processes Simulated Using Aspen

Plus


Hydrogen (H2) has the potential to shift the global reliance on fossil fuel energy sources to cleaner,

more efficient forms of energy as it presents a viable alternative. The issues of global carbon

dioxide (CO2) and other greenhouse gas (GHG) emissions and overall atmospheric concentration

can also be addressed with the utilization of H2 technologies for the generation of energy.

Furthermore, H2 presents an advantage over other conventional alternative energies such as wind

and solar due to its energy storage and transport capabilities. Thus, H2 utilization has the potential

to become widespread in the near future.

Current methods of H2 production involve the reforming of fossil fuels; processes that

ultimately contribute to the societal and environmental issues mentioned above. Correspondingly,

contemporary H2 energy systems are neither GHG-neutral nor sustainable since these non-

renewable fossil fuels are responsible for equivalent CO2 emissions. Thus, it is essential that H2 be

produced from a renewable, carbon-neutral energy source. The thermochemical gasification-based

conversion of biomass in the presence of steam presents a viable renewable H2 source and is a

strong contender for the replacement of fossil fuel-based H2 production.

Chemical looping gasification (CLG) using biomass fuel is an example of such H2 energy

technologies as it capitalizes on renewable, environmentally friendly, and abundant sources of

energy. Biomass is an organic fuel source containing carbon, hydrogen, oxygen, nitrogen, and

sulfur, and can come in the form of agricultural wastes, municipal solid wastes, animal wastes,

saw dust, etc. [1]. The gasification of biomass involves thermochemical conversion to H2, as well

as various hydrocarbons, for subsequent use in H2 conversion technologies with the aim of energy

production. Furthermore, biomass gasification is a complex, endothermic process consisting of

many chemical reactions. These reactions depend on the gasification agent being used which is

commonly steam. The reactions are summarized in Table 4.1 [2].

54

Table 4.1

Chemical reactions in the steam gasification of biomass. Adapted from [2].

Reaction Chemical Equation ΔHo923 (kJ mol-1)

Water-gas shift CO + H2O → CO2 + H2 -35.6 (exothermic)

Methane reforming CH4 + H2O → CO + 3H2 225 (endothermic)

Water-gas (i) C + H2O → CO + H2 136 (endothermic)

Water-gas (ii) C + 2 H2O → CO2 + 2H2 100 (endothermic)

Oxidation (i) C + O2 → CO2 -394 (exothermic)

Oxidation (ii) C + ½ O2 → CO -112 (exothermic)

Boudouard C + CO2 → 2CO 171 (endothermic)

Methanation C + 2H2 → CH4 -89.0 (exothermic)

At the present time, minimal research has been conducted regarding CLG systems

development. Moreover, the major research focus involves the chemical looping combustion

(CLC) process, with emphasis on oxygen carrier development for use in CLC. Though widely

recognized by many researchers, work conducted on the CLG process developed by Fan et al. at

Ohio State University [3] requires further research and development to become a fully-

implemented, renewable H2 production technology. Furthermore, other advanced chemical

looping systems exemplar of CLG processes for the purpose of H2 production include the HyPr-

RING (Hydrogen Production by Reaction-Integrated Novel Gasification), fuel-flexible advanced

combustion-gasification, ALSTOM hybrid gasification-combustion, Advanced Gasification-

Combustion (AGC), and Zero Emission Coal Alliance (ZECA) processes [3-5].

The mitigation and subsequent elimination of sorbent/catalyst performance degradation

over time is an ongoing challenge in most chemical looping systems. Successful operation of

chemical looping systems is often jeopardized by unpredictable sorbent behaviour. Thus, research

has been conducted to identify the mechanisms contributing to sorbent losses and develop

performance improvement methods to ensure sufficient system operation and longevity.

Subsequently, sintering of the chemical looping sorbent particles as a result of high-temperature

operation and cyclic heating/cooling cycles, in conjunction with char and tar deposition,

significantly reduces sorbent capture and regeneration capabilities. The detrimental effects of these

phenomena are most notably felt in CLG systems. Another challenge presented by chemical

looping systems involves the continuous flow of solid materials between a multitude of

interconnected reactors operating at high temperatures and pressure [6].

The CLG of biomass is advantageous in that it is a clean, renewable source of H2; however

some disadvantages remain prominent. For example, product syngas streams resulting from the

55

CLG of biomass contain many impurities. Consequently, much research has been conducted

regarding syngas impurity removal, whether it be the end-of-pipe or in situ approaches. Therefore,

the research compares two different CLG processes under similar feed and operating conditions

for the purpose of H2 production using a computer simulation. The novelty of the conducted

research lies in the development of the two CLG processes using the Aspen Plus simulation

software and subsequent comparison of the results using the same biomass feedstock in both cases.

The paper also includes temperature and pressure sensitivity analyses conducted on all

relevant reactors and the results compared between the CLG 1 and CLG 2 processes. Relevant

parameters included optimal operating temperature and pressure for each reactor and product

syngas molar yield and composition. Comparison of the H2 production and purity capabilities of

the two processes is emphasized upon.

4.1.1 Simulated Processes

Two CLG mechanisms were simulated and analyzed using the Aspen Plus V7.3 software. The

simulation results were compiled for poultry litter (PL), a nonconventional biomass type.

4.1.1.1 Chemical Looping Gasification Type 1

The first CLG simulation (CLG 1) incorporated in situ product CO2 capture in the absorbing

reactor with the use of a CaO sorbent. A representative block diagram can be seen in Figure 4.1.

CO2 absorption occurs according to the following chemical reaction [7]:

CaO (s) + CO2 (g) → CaCO3 (s)

This reaction is exothermic with a heat of reaction of -178.3 kJ mol-1 [7]. In addition, near-total

CaO sorbent recovery and recycle was inherent to the simulation setup. CO2 desorption is the

reverse of CO2 absorption and is therefore an endothermic reaction. The energy required for the

reaction to occur is provided by the higher operating temperature of the desorbing reactor. The

resulting theoretical product stream is pure CO2 which can be sent for sequestration. Furthermore,

the overall reaction prior to sorbent regeneration is as follows:

CnHmOp + (2n – p) H2O + n CaO → n CaCO3 + (m/2 + 2n – p) H2

This overall reaction is endothermic with a heat of reaction of +107.5 kJ mol-1 [6]. The constants

n, m, and p represent the respective carbon, hydrogen, and oxygen contents in the biomass being

gasified.

56

Fig. 4.1. CLG 1 simulation block diagram.

Another key aspect of the CLG 1 simulation is the use of steam (H2O (g)) to address the

issue of tar and char formation. Tar and char were modelled as pure carbon and the reforming

reactions correspond to water-gas (i) and water-gas (ii) from Table 4.1 as follows:

C(s) + H2O (g) → CO (g) + H2 (g)

C(s) + 2 H2O (g) → CO2 (g) + 2 H2 (g)

These reactions were assumed to reform all tar compounds exiting the desorbing reactor.

Furthermore, the use of steam is preferable to conventional air since the thermo-oxidative

reforming of pure carbon in air results in a nitrogen (N2) stream which must be subsequently

separated from the CO2 product stream. This process is energy-intensive and therefore costly [8],

however the condensation of steam from an exit gas stream can also be a costly process. A more

practical process schematic illustrating an overview of the CLG 1 process can be seen in Figure

4.2.

57

Fig. 4.2. CLG 1 process schematic [6].

4.1.1.2 Chemical Looping Gasification Type 2

Subsequently, syngas chemical looping is another H2-production process involving the

gasification of biomass. This process separates the gasification and looping stages and produces

H2 via reduction, oxidation, and combustion cycles involving iron (Fe), hematite (Fe2O3 or iron

(III) oxide), and magnetite (Fe3O4 or iron (II,III) oxide) [6]. This process was simulated using

Aspen Plus V7.3 using poultry litter as the chosen biomass type (CLG 2). The block diagram for

the CLG 2 process can be seen in Figure 4.3.

58

Fig. 4.3. CLG 2 simulation block diagram [3].

Similarly to CLG 1, CLG 2 is comprised of many chemical reactions involving a multitude

of reactors. These reactions and corresponding reactors are outlined in Table 4.2 [3]. The reactions

persist under theoretical conditions. In actuality, tar formation in the reducer is unavoidable. Tar

reforming in the CLG 2 simulation is carried out in the oxidizer, where large amounts of steam are

introduced. Here, tar is completely reformed by the steam to CO, CO2, and H2 via the

aforementioned water-gas reactions presented in Table 4.1. Additionally, pure O2 is utilized in the

combustor as opposed to air for the same reasons explained for steam in the CLG 1 simulation [8].

Table 4.2

Chemical reactions in syngas chemical looping [3].

Reactor Chemical Equation Description

Gasifier Biomass + H2O → H2 + CO Steam gasification of biomass.

Combustor 4 Fe3O4 + O2 → 6 Fe2O3 Forms Fe2O3 for reducer.

Reducer 3 CO + Fe2O3 → 3 CO2 + Fe

3 H2 + Fe2O3 → 3 H2O + 2 Fe

Forms Fe for oxidizer and CO and

H2O to be separated.

Oxidizer 3 Fe + 4 H2O → Fe3O4 + H2 Forms product H2 and Fe3O4 for

recycle to combustor.

59

4.2 Feedstock Used

4.2.1 Composition of Biomass Types

The characteristics of three biomass types were compared to identify a suitable feedstock for

utilization in the two simulations. The types are: poultry litter, wood pellets, and oak pellets. The

biomass types vary in their chemical composition and are thus representative of a spectrum of

biomass types, with poultry litter representing a nonconventional type. Table 4.3 summarizes

ultimate analyses of poultry litter while Table 4.4 summarizes its proximate analysis.

Table 4.3

Ultimate analysis of poultry litter in both the presence and absence of sulfur and nitrogen.

Mass Composition (wt.-%)

Element Sulfur and Nitrogen Present Sulfur and Nitrogen Absent

Carbon 43.30 46.49

Hydrogen 6.62 7.11

Oxygen 5.95 6.39

Nitrogen 5.72 -

Sulfur 1.15 -

Ash 37.26 40.01

Table 4.4

Proximate analysis of poultry litter.

Parameter Mass Composition (wt.-%)

Moisture Content 20.10

Fixed Carbon 3.33

Volatile Matter 54.29

Ash 22.28

4.2.2 Chemical Equations for Gasification of Biomass Types

Each biomass type was assumed to gasify according to the following chemical formula [9]:

CnHmOp + (n – p) H2O → n CO + (m/2 + n – p) H2

The ultimate analysis of each biomass type assuming an absence of both sulfur and nitrogen was

used to calculate the n, m, and p values in the equation above. Moreover, the chemical composition

of each biomass type varied in terms of hydrogen and oxygen content, and was calculated relative

to carbon molar content. The corresponding chemical formulas for each of the biomass types were

found to be CH0.01286O0.1831, CH0.00973O0.6331, and CH0.01061O0.8667 for poultry litter, willow pellets,

and oak pellets, respectively. Poultry litter can be seen to contain the greatest hydrogen and lowest

60

oxygen content per mole of biomass. Consequently, the chemical equation for the steam

gasification of poultry litter is as follows:

CH0.01286O0.1831 (s) + 0.8169 H2O (g) → CO (g) + 0.8234 H2 (g)

It is evident that the steam gasification equations for willow and oak pellets would be of the same

form as poultry litter but rather with different molar coefficients for steam consumption and lesser

values for H2 generation. Poultry litter generated the greatest theoretical H2 yield as roughly 0.82

moles were generated per mole of biomass and was therefore chosen as the biomass type to be

used for both the CLG 1 and CLG 2 simulations.

4.3 Simulation Input Parameters and Description

The following sections outline the input data to the Aspen Plus simulation engine as well as the

chosen calculation methods. Moreover, detailed descriptions of the utilized flowsheets are

provided.

4.3.1 Setup and Calculation Methods

The flowsheet type was chosen as “Solids with metric units”, allowing for the analysis and results

presentation for solid-state input and output streams. The setup of the flowsheet involved assigning

the MIXCINC stream class to the simulation. This allowed fluid and aqueous streams (MIXED),

conventional solid streams (CISOLID), and nonconventional solid streams (NC) to be input and

analyzed during simulation runs and calculations.

Subsequently, the process type was chosen as COMMON. This allotted a generic industry

type to the simulation, as opposed to chemical, petrochemical, pharmaceutical, etc. The IDEAL

base calculation method was selected for simplicity and thus phase equilibrium calculations were

conducted using Raoult’s Law, Henry’s Law, the ideal gas law, etc.

4.3.2 Component Definition and Input

Solid biomass was modelled using a user-defined, nonconventional solid based on ultimate,

proximate, and sulfur analyses. Thus, the input for the poultry litter biomass type was based on

these parameters. Sulfur analyses – including pyritic, sulfate, and organic – were set to zero.

Furthermore, the enthalpy and density of biomass were approximated using coal properties. The

methods used by Aspen Plus for these calculations are HCOALGEN and DCOALIGT,

respectively.

61

Fluid streams were modelled using conventional components which have thermophysical

data stored in Aspen Plus databanks. Therefore, no data input were required for these components.

The components include: hydrogen (H2), water (H2O), carbon monoxide (CO), carbon dioxide

(CO2), methane (CH4), and oxygen (O2).

Additionally, solid components were modelled using conventional solids which also have

necessary thermophysical data stored in the databanks. Tar formation was approximated as solid

carbon (i.e. graphite) in the simulation. The components include: tar (C), calcium oxide (CaO),

calcium carbonate (CaCO3), iron (Fe), hematite (Fe2O3), and magnetite (Fe3O4). Calcium-based

components were exclusive to the CLG 1 simulation while iron-based components were exclusive

to the CLG 2 simulation, with fluid components being involved in both.

4.3.3 CLG 1 Flowsheet Description

The CLG 1 simulation flowsheet can be seen in Figure 4.4. The input and operating conditions for

all feed streams and block units are summarized in Tables 4.5 and 4.6, respectively.

Biomass and water at ambient conditions were fed to the gasifier after being heated to the

reactor temperature. This block gasified the biomass based on user-defined output for H2 and CO.

The output from the gasifier was then fed to the reforming reactor where further gasification

reactions occurred, resulting in H2, CO, CO2, CH4, tar, and steam formation. Subsequently, the

output from the reformer was fed to the absorbing reactor in conjunction with a CaO feed stream

used for CO2 absorption and capture. The solid and gaseous products from the absorber were then

separated with a gas-solid separator with an efficiency of 99.9%. The product gases were heated

to the WGS reactor temperature and further gasified to convert the majority of the remaining CO

to H2. Tar products contained in the WGS reactor exit stream were removed at a removal efficiency

of 99.9%, and sent to the desorbing reactor for steam reforming. The gaseous products of this

stream were subsequently condensed to remove most of the remaining steam and small amounts

of tar from the syngas, consequently increasing H2 purity in the product gas. It is important to note

that the gasification process was modelled with the combined use of the gasifying, reforming,

absorbing, and WGS reactor blocks rather than in a single reactor step. This resulted from

limitations inherent to the Aspen Plus block unit input capabilities.

62

Fig. 4.4. CLG 1 simulation flowsheet.

63

Table 4.5

Feed stream input conditions for CLG 1 simulation.

Feed Stream

Input Conditions

Temperature

(°C)

Pressure

(atm)

Flowrate

(kmol h-1) Component

BIOMASS 25 1 1a Biomass (Nonconventional)

H2O-FEED 25 1 1 H2O (Conventional)

CAO-FEED 25 1 6b CaO (Conventional Solid)

STEAM 240 1 85b H2O (Conventional) a Input as mass flowrate (kg h-1) using biomass molecular weight. b Fed in excess of required stoichiometric amount.

Table 4.6

Block unit operating conditions for CLG 1 simulation.

Block Information Operating Conditions

Name Type Temperature

(°C)

Pressure

(atm)

Other

GASIFIER RYield 750 1 Output based on set values for H2 and CO (units of kg /

kg total feed)

REFORMER RGibbs 750a 1d -

ABSORBER RGibbs 500b 1d -

WGS-RCTR RGibbs 750c 1d -

DESORBER RGibbs 650 1 -

HEATER Heater 750 1 -

HEAT-GAS Heater 750e 1 -

COOL-CAO Heater 25 1 -

COOL-GAS Heater 25 1 -

G-S-SEP1 Sep - -

Separated gaseous components (H2, CO, CO2, CH4, and

H2O) from 99.9% of solid components (C, CaO, and

CaCO3).

G-S-SEP2 Sep - - Separated DESORBER gases from 99.9% of CaO.

G-S-SEP3 Sep - - Separated syngas products from 99.9% of tar (C).

CONDENSE Flash2 20 5 - a Represents optimal operating temperature. Reformer temperature was varied from 500 to 900 °C. b Represents optimal operating temperature. Absorber temperature was varied from 300 to 800 °C. c Represents optimal operating temperature. WGS reactor temperature was varied from 400 to 1000 °C. d Represents optimal operating pressure. Pressure was varied from 1 to 20 atm. e HEAT-GAS heater temperature was set to the WGS reactor temperature being simulated.

Furthermore, the solids stream exiting the initial gas-solid separator containing CaCO3, tar,

and unused CaO was fed to the desorbing reactor, along with tar products from the WGS reactor

exit. Here, CaCO3 was thermally dissociated into CaO and CO2, whereas all of the present tar was

reformed to CO, CO2, and H2 with the use of steam. An additional gas-solid separator was utilized

to separate the resulting solids (CaO and remaining tar) and desorber gases. The unused and

regenerated CaO stream was cooled to ambient temperature and sent for re-use in the absorber.

64

Further, the resulting CO2-rich stream was also cooled to ambient temperature and sent for

sequestration.

4.3.4 CLG 2 Flowsheet Description

The CLG 2 simulation flowsheet can be seen in Figure 4.5. The input and operating conditions for

all feed streams and block units are summarized in Tables 4.7 and 4.8, respectively.

Biomass and water at ambient conditions were fed to the gasifier after being heated to the

reactor temperature. This block gasified the biomass based on user-defined output for H2 and CO.

The output from the gasifier was then fed to the reducer. Conversion of Fe3O4 to Fe2O3 via

combustion using pure O2 was carried out to facilitate Fe-generating reduction reactions, of which

the product Fe was further oxidized using H2O for H2 production. As such, the Fe2O3 resulting

from combustion was combined with the gasification products in the reducer. Here, Fe2O3 was

reduced to Fe and the gasification products underwent the various gasification reactions, resulting

in H2, CO, CO2, CH4, tar, and steam. The solid-state and gaseous products were separated in the

cyclone at a solids removal efficiency of 99.9%. The gases were then fed to the reforming reactor

while the Fe and tar were fed to the oxidizing reactor for Fe3O4 regeneration. Furthermore, CH4

completely reformed to CO and H2 in the oxidizer. The oxidizer exit gases and solids were then

separated in a secondary cyclone, with over 99.8% regenerated Fe3O4 to be recycled to the

combustor and gases fed to the reformer.

Consequently, the reducer and oxidizer exit gases were further reformed to H2 and CO2 in

the reformer to increase the H2 yield of the system. Small amounts of CH4 were regenerated in the

reformer. The reformer exit gases were next fed to a condenser to remove most of the remaining

steam and residual solids from the resulting syngas. This increased H2 purity in the syngas. CO2

capture was not inherent to CLG 2 as was the case with CLG 1.

65

Fig. 4.5. CLG 2 simulation flowsheet.

66

Table 4.7

Feed stream input conditions for CLG 2 simulation.

Feed Stream

Input Conditions

Temperature

(°C)

Pressure

(atm)

Flowrate

(kmol h-1) Component

BIOMASS 25 1 1a Biomass (Nonconventional)

H2O-IN 25 1 1 H2O (Conventional)

FE3O4-IN 25 1 0.0370b Fe3O4 (Conventional Solid)

O2 25 1 0.10b O2 (Conventional)

STEAM 240 32 20 H2O (Conventional) a Input as mass flowrate (kg h-1) using biomass molecular weight. b Required stoichiometric amount for down-stream reducer reactions to occur.

Table 4.8

Block unit operating conditions for CLG 2 simulation.

Block Information Operating Conditions

Name Type Temperature (°C) Pressure

(atm)

Other

GASIFIER RYield 750 1 Output based on set values for H2 and CO

(units of kg / kg total feed)

COMBUST RGibbs 1250 1 -

REDUCER RGibbs 870 30 -

OXIDIZER RGibbs 720 30 -

REFORMER RGibbs 500 1 -

HEATER Heater 750 1 -

CYCLONE Sep

- - Separated 99.9% of Fe and tar (C) from

reducer exit gases (H2, CO, CO2, H2O, and

CH4).

CYCLONE2 Sep

- - Separated 99.9% of Fe3O4 from oxidizer

exit gases (H2, CO, CO2, and H2O)

CONDENSE Flash2 20 1 -

4.4 Results and Discussion

Important assumptions were initially applied to the simulation to ensure that it ran smoothly and

produced results. These included:

1. Chosen biomass types contained no nitrogen or sulfur.

2. Biomass and steam reacted completely in the first gasification reactor (i.e. the gasifier) and

the only products were H2 and CO.

3. Tar by-products formed in the various reactors were pure carbon.

4. Tar reforming using steam in the CLG 1 simulation was 100% efficient.

5. CO2 desorption from CaCO3 in the CLG 1 simulation was 100% efficient.

6. Gas-solid separators were 99.9% efficient.

67

The following sections detail the results obtained from both simulations using poultry litter.

Temperature and pressure sensitivity analyses were conducted for each of the main reactors to

determine the optimal operating conditions for those reactors. The outputs to the gasifier were

specified as part of the simulation in both cases, and thus no sensitivity analysis could be carried

out on those block units.

Furthermore, the reactor type utilized for the absorber, WGS reactor, desorber, combustor,

reducer, oxidizer, and both reformers was RGibbs. This reactor is an Aspen Plus block unit which

calculates its output using the Gibbs free energy minimization method. The calculations are based

on the chemical equilibrium reactions of the components being input to the reactor under the

specified operating conditions.

4.4.1 Determining the Optimal Operating Conditions

4.4.1.1 CLG 1 Results

The reformer in the CLG 1 simulation aimed to produce more H2 following the gasifier via CO

conversion. The aforementioned gasification equations occurred simultaneously, resulting in by-

product formation in the output stream. The results of the reformer temperature sensitivity analysis

can be seen in Figure 4.6, which shows that H2 and CO yield at the reformer exit increased with

increasing operating temperature. Further, by-product formation, aside from CO, tended to

decrease with temperature, and sharply so at the higher end of the temperature scale. These

phenomena are the result of the gasification equations proceeding in the reformer. The forward

reactions favour H2 and CO production and are endothermic. Thus, increasing temperature

favoured the formation of these species. As a result, the optimal operating temperature was subject

to debate based on the criterion of higher H2 yield versus greater H2 purity. The temperature of

750 °C was chosen as optimal since H2 yield was relatively high at this point while CO yield had

not yet reached its highest value. Moreover, by-product yield was relatively low at this

temperature. The trends observed for the reformer temperature sensitivity analysis closely

reflected those found by Mahisi et al. [8]. However, ethanol (C2H5OH) was used as the source

biomass in that case.

Next, the results of the pressure sensitivity analysis on the reformer can be seen in Figure

4.7, which illustrates that the opposite phenomenon from temperature sensitivity occurred. Both

H2 and CO yield began to decrease rapidly with increasing pressure, especially within the first five

atmospheres of pressure increase. Correspondingly, other by-product formation increased rapidly

68

within the first five atmospheres. Due to the relatively large reduction in H2 yield and rapid rise of

by-product formation at higher pressures, atmospheric pressure was determined to be optimal for

the reforming reactor. The trends observed for the pressure sensitivity analysis also closely

reflected those observed by Mahisi et al. [8].

Fig. 4.6. CLG 1 reformer temperature sensitivity analysis.

Furthermore, the main goal of the absorber was to capture generated CO2 using CaO as the

sorbent. As previously mentioned, this is an exothermic equilibrium reaction which typically

occurs at temperatures from 600 to 650 °C. Absorber exit yield, CO2 capture efficiency, and CO2

product yield (i.e. CO2 sent for sequestration) were parameters of consideration. It was noted that

H2 yield tended to decrease at the absorber exit relative to H2 in the feed to the reactor. This

phenomenon was deemed a necessary sacrifice to ensure total CO2 capture, and was compensated

for further downstream with the use of the WGS reactor.

0

0.2

0.4

0.6

0.8

1

1.2

500 550 600 650 700 750 800 850 900

Yie

ld:

km

ol/

km

ol

PL

Reformer Temperature (°C)

Reformer Exit Yield vs. Reformer Temperature

H2 CO CO2 CH4 Tar H2O

69

Fig. 4.7. CLG 1 reformer pressure sensitivity analysis.

A temperature sensitivity analysis was also conducted on the absorbing reactor. The

analysis showed that H2 production increased with increasing absorber temperature. Conversely,

CO2 production decreased until reaching a local minimum at 750 °C. Accordingly, CO2 capture

efficiency slowly decreased with temperature before being rendered totally ineffective at 750 °C.

Approximately 500 °C was concluded as the optimal operating temperature for the absorber, based

on H2 and CO2 production and CO2 capture efficiency. Higher CO2 production, relatively high H2

yield, and a CO2 capture efficiency of over 99% were all present at this temperature, which

occurred prior to the rapid decline in CO2 capture. Moreover, the pressure sensitivity analysis

conducted on the absorber yielded similar results to that of the reformer, and thus atmospheric

pressure was deemed optimal for the absorber. Both CO2 production and capture efficiency slightly

increased with pressure; however minimal gains, on the order of 0.01%, were observed and were

insufficiently beneficial to merit operating the absorber at higher pressures.

In addition, CO2 desorption was totally effective in the simulation. CaO sorbent recovery

was not, since 0.01% of the sorbent feed was lost due to inefficiencies in the gas-solid separators

throughout the process. In an experimental setting, however, greater amounts of sorbent would be

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

Yie

ld:

km

ol/

km

ol

PL

Reformer Pressure (atm)

Reformer Exit Yield vs. Reformer Pressure

H2 CO CO2 CH4 Tar H2O

70

rendered ineffective as a result of calcium sulfate (CaSO4) formation from any sulfur components

present in the biomass feedstock [8], which was not considered in the conducted research. Thus,

greater amounts of sorbent regeneration would be required in experimental and real-life scenarios

than what is implied by the CLG 1 simulation.

Subsequently, the purpose of the WGS reactor was to provide a final block unit for H2

production as well as to regain H2 that was lost as a result of CO2 absorption in the absorbing

reactor. The lowered temperature required for CO2 absorption to proceed was unfavourable to H2

production and retention in the absorbing reactor. Both H2 yield and content significantly increased

as a result of the gasification reactions that occurred within the WGS reactor. Furthermore, H2

content in the syngas product stream was further upgraded with condenser utilization in the

following step. This block unit effectively liquefied the majority of the steam present in the WGS

reactor exit stream, thereby altering the syngas composition in favour of H2.

The results of the temperature sensitivity analysis conducted on the WGS reactor can be

seen in Figure 4.8. Both H2 and CO yield increased significantly after 500 °C with increasing

temperature, and then began to plateau at higher temperatures. Although the CO yield increase

with temperature was pronounced, the H2 to CO yield ratio remained very high due to low

concentrations of CO. Other by-product formation generally declined with increasing temperature

and was even negligible in the case of CO2. Accordingly, the optimal operating temperature for

the WGS reactor was found to be 750 °C, which provided relatively high H2 yield while

maintaining a lower total by-product yield. This temperature is considerably higher than the 300

°C value observed by Mahisi et al. [8]. However, H2 yield increased following WGS reactor

utilization in both CLG 1 and the simulation conducted by Mahisi et al. The reactor in that

simulation accounted for an 8% increase in H2 yield, while the CLG 1 simulation exhibited an

approximate 170% increase. The substantially larger increase in CLG 1 can be attributed to the

requirement for the regeneration of reduced amounts of H2 following the CO2 absorption step.

Additionally, the pressure sensitivity analysis conducted on the WGS reactor demonstrated that

atmospheric pressure was optimal. H2 production decreased and by-product formation increased

with pressure, similarly to trends observed for the previous reactors.

Overall, the optimal reactor temperatures for the CLG 1 simulation were 750 °C, 500 °C,

and 750 °C for the reformer, absorber, and WGS reactors, respectively. Therefore, a value of 750

°C was chosen for the gasifier as a conservative estimate. Moreover, this value is in agreement

71

with some gasifier temperatures utilized in coal-type or wood-type biomass gasification studies

outlined in the literature [10-20], and is 80 °C warmer than the value of 670 °C reported by Acharya

[6].

Fig. 4.8. CLG 1 WGS reactor temperature sensitivity analysis.

4.4.1.2 CLG 2 Results

Combustor operating conditions were chosen as 1250 °C and 1 atm based on data presented by

Fan [3]. These conditions proved sufficient to fully convert Fe3O4 to Fe2O3 via thermo-oxidation

given a stoichiometric excess of pure O2.

The CLG 2 simulation reducer aimed to generate Fe by reacting Fe2O3 with the H2 and CO

gasification products. By-product formation was evident in the simulation. These phenomena are

better illustrated in the temperature sensitivity analysis of Figure 4.9. The analysis was conducted

at 30 atm rather than atmospheric pressure as this was the recommended pressure proposed by Fan

[3]. Also, Fe formation was determined to be undesirably low at atmospheric pressure. Figure 4.9

displays that minimal increases in H2 and CO yield were initially observed before slowly

decreasing at the critical point of roughly 875 °C. Conversely, at this critical point, by-product

formation (CO2, H2O, CH4, and tar) tended to increase with temperature. Fe formation was

0

0.002

0.004

0.006

0.008

0.01

0.012

0.014

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

400 500 600 700 800 900 1000

Yie

ld:

km

ol/

km

ol

PL

(C

O,

CO

2,

H2O

)

Yie

ld:

km

ol/

km

ol

PL

(H

2an

d C

H4)

WGS Reactor Temperature (°C)

Syngas Yield vs. WGS Reactor Temperature

H2 CH4 CO CO2 H2O

72

virtually unaffected by temperature variation for the tested range. Therefore, 870 °C was chosen

as the optimal reducer temperature. This value was chosen due to its slightly lower cost

implications when compared to the critical point of 875 °C which was too proximate to the

decreasing portion of the H2 yield curve. Moreover, the optimal operating pressure of 30 atm was

confirmed via a pressure sensitivity analysis on the reducer, ranging from 30 to 37 atm, conducted

at the optimal temperature of 870 °C. H2 and CO production decreased with increasing pressure

while by-product formation increased, similar to the trends of the CLG 1 pressure sensitivity

analyses.

Fig. 4.9. CLG 2 reducer temperature sensitivity analysis.

In addition, the general purpose of the CLG 2 oxidizer was both to produce H2 in greater

quantities than the gasifier and to regenerate spent Fe3O4. The temperature sensitivity analysis for

the oxidizer was again conducted at 30 atm rather than atmospheric pressure as this was the

recommended pressure proposed by Fan [3]. The analysis demonstrated that temperature variation

only had a significant effect on CO and CH4 yield. CO increased relatively quickly with

temperature and quantities of CH4 were negligible below about 680 °C. Thus, 720 °C was chosen

as the optimal oxidizer temperature as only trace amounts CH4 were present, and negligible

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

870 872 874 876 878 880 882 884 886 888 890

Yie

ld:

km

ol/

km

ol

PL

Reducer Temperature (°C)

Reducer Exit Yield vs. Temperature at 30 atm

H2 CO CO2 H2O CH4 Tar Fe

73

changes in H2 yield were observed. Oxidizer pressure was incrementally increased to 40 atm to

determine pressure effects. However, minimal changes in component yield were observed in any

case, and 30 atm was confirmed as the optimal operating pressure. Finally 20 kmol h-1 was chosen

as the design feed steam flowrate to the oxidizer after flowrate variation analysis was conducted

from 5 to 25 kmol h-1 to determine the effects on component yields. The analysis was carried out

under steam conditions of 240 °C and 32 atm. By-product CO yield tended to decrease with

increasing flowrate, with other component species virtually unaffected, and so a higher flowrate

was chosen to minimize CO yield.

Fig. 4.10. CLG 2 reformer syngas yield temperature sensitivity analysis.

Subsequently, the CLG 2 reformer was meant to further increase H2 yield and content via

reforming of the remaining by-products. The temperature sensitivity analysis was conducted and

the optimal operating temperature chosen based on H2 yield and by-product levels present in the

syngas stream exiting the condenser unit at the process termination. The observed results from the

analysis can be seen in Figure 4.10. It can be seen that H2 and CO2 yield in the syngas begin to

decrease after reformer temperatures reach roughly 500 °C. CO yield increased rapidly over the

simulated range, albeit at lower concentrations throughout. CH4 yield rapidly decreased between

0

0.01

0.02

0.03

0.04

0.05

0.06

0.07

0.08

0.09

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

1.8

400 450 500 550 600 650 700 750 800 850

Yie

ld:

km

ol/

km

ol

PL

(C

O,

H2O

, C

H4)

Yie

ld:

km

ol/

km

ol

PL

(H

2an

d C

O2)


Syngas Yield vs. Reformer Temperature

H2 CO2 CO H2O CH4

74

400 and 500 °C. H2 yield tended to peak at approximately 1.60 kmol / kmol PL. In addition, syngas

composition was closely examined under varying temperature conditions. The respective

component curves had roughly the same shape as the syngas yield curves which are illustrated in

Figure 4.10. H2 and CO2 comprised the majority of the syngas with CO, H2O, and CH4 by-products

accounting for smaller percentages. Again, H2 composition peaked in the 450 to 500 °C range and

began to decrease with further temperature increase.

Based on the aforementioned trends, 500 °C was chosen as the optimal operating

temperature for the CLG 2 reforming reactor. This is due to both higher H2 yield and content in

the resulting syngas stream, as well as lower by-product yields and compositions in proximity to

this temperature. Increasing operating pressure above atmospheric conditions tended to decrease

desirable product yield and correspondingly increased by-product yield and concentration. Thus,

a detailed pressure sensitivity analysis was not further pursued for the CLG 2 reformer.

4.4.2 Comparison of Simulation Results

The following section details a comparison between the results of the two biomass gasification

simulations. Comparisons are made based on syngas yield and composition.

4.4.2.1 Syngas Yield Comparison

A comparison of the absolute syngas yields for the CLG 1 and CLG 2 simulations can be seen in

Figure 4.11. It is evident that CLG 2 generated more syngas than CLG 1, at values of roughly 2.54

and 0.79 kmol / kmol PL, respectively. However, it is also important to note that CO2 removal was

not a focal point of CLG 2, and thus almost 0.87 kmol CO2 / kmol PL adds to the absolute CLG 2

syngas value. Furthermore, H2 can be seen to be the main constituent in both cases, with the CLG

2 syngas producing more H2 in absolute terms.

CLG 1 generated 0.73 kmol H2 / kmol PL while CLG 2 generated 1.60 kmol H2 / kmol PL.

These values are less than those reported in the literature for either case. Processes similar to CLG

1 reported H2 yields ranging from 1.6 to roughly 5.7 kmol H2 / kmol biomass [2,8], while Fan

outlined a process similar to CLG 2 capable of producing approximately 11.73 kmol H2 / kmol

coal [3]. The latter process, however, used coal as the solid fuel and greater amounts of iron-based

oxygen carriers than CLG 2.

Furthermore, the assumption that all biomass and steam were completely converted to H2

and CO in the gasifier was challenged, and its effect on syngas yield in both cases was determined.

75

This was done in terms of conversion efficiency, labelled as gasifier efficiency. The analyses

showed that the yields of all syngas components linearly decreased with a decrease in gasifier

efficiency. It is also noted that the reduction in H2 yield from 100 to 50% gasifier efficiency is

relatively drastic; 0.73 to 0.36 kmol / kmol PL.

Fig. 4.11. Comparison of simulation syngas yields.

4.4.2.2 Syngas Composition Comparison

A comparison of the syngas compositions in either simulation can be seen in Figure 4.12, which

also includes individual component compositions. It can be seen that H2 is the main component in

either syngas stream, at 92.45 mol-% and 62.94 mol-% for CLG 1 and CLG 2, respectively. The

former value is greater than the 80.94% H2 value reported by Acharya under similar conditions.

Further, the CO2 concentration of 0.01% in CLG 1 syngas is significantly lower than the value of

5.71% reported in the literature [6]. Other studies similar to CLG 1 reported lower obtained H2

concentrations, ranging from roughly 70 to 90 mol-% [2,7,8].

However, as previously mentioned, CO2 was not removed from the syngas stream in CLG

2, and thus its relatively large constituency of 34.11 mol-% is accounted for by that fact. Had CO2

removal been inherent to the design of CLG 2, and assuming removal efficiency upwards of 99%,

0.87

2.54

0

0.5

1

1.5

2

2.5

3

CLG 1 CLG 2

Yie

ld (

km

ol/

km

ol

PL

)

Syngas Yield Comparison at Optimal Operating Conditions

H2 CO CO2 CH4 H2O

76

the H2 composition of the CLG 2 syngas would be over 95 mol-%. Further, the 62.94 mol-% value

from CLG 2 is similar to the 62.1 mol-% found in the literature [3].

Fig. 4.12. Comparison of simulation syngas compositions.

4.5 Potential for Future Research

Potential for future research stems in-part from the simulation flowsheets and block unit setups

themselves. Determining a method to model biomass gasification in a single gasifying reactor, as

opposed to four separate reactors in the case of the CLG 1 process, is exemplar of this. This would

allow for biomass gasification to occur spontaneously in an equilibrium-based reactor, i.e. RGibbs,

rather than in a reactor with user-defined outputs, as is the case for both simulations. Further, this

would allow for an investigation of the sensitivity of the gasifying reactor to steam-to-biomass,

calcium-to-biomass, and equivalence ratios.

In addition, the nitrogen and sulfur components of each biomass type could be included in

calculations to further enhance the accuracy of the simulation. Consequently, this would require

modification of the flowsheet and input parameters to deal with any sulfur dioxide (SO2), hydrogen

sulfide (H2S), or nitrogen (N2) streams that may be present from gasification of biomass containing

these elements [21].

92.45%

62.94%

1.28%

0.62%

0.01%

34.11%

6.03% 0.03%0.23% 2.30%

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

CLG 1 CLG 2

Co

mp

osi

tio

n (

mo

l-%

)

Syngas Composition Comparison at Optimal Operating Conditions

H2 CO CO2 CH4 H2O

77

Moreover, more detailed design of the gas-solid separators (e.g. cyclones) utilized in the

simulation would further increase its accuracy. For example, cyclones, baghouses, or electrostatic

precipitators (ESPs) could be designed for gas-solid separation. These process units would be more

representative of an industrial-scale scenario, as opposed to the generic separator blocks used in

this simulation. Furthermore, the energy requirements of such block units could be accurately

represented as well.

Finally, a cost analysis for each block unit could be conducted to obtain an overall syngas

production cost for each simulated process type. The cost analysis could be based on energy

requirements and feed stream raw materials supply costs. Subsequently, comparing the cost

estimates would provide an idea of the feasibility of the proposed processes from a practical

standpoint.


In conclusion:

1. Poultry litter was the chosen biomass type to be simulated both because it is a

nonconventional biomass type and due to its greater H2 yield potential when compared

with willow pellets and oak pellets.

2. The optimal operating condition estimates determined for the main reactors in both

simulations were in line with those presented in the literature.

a. CLG 1 simulation:

i. Reformer: 750 °C, 1 atm.

ii. Absorber: 500 °C, 1 atm.

iii. WGS reactor: 750 °C, 1 atm.

b. CLG 2 simulation:

i. Combustor: 1250 °C, 1 atm.

ii. Reducer: 870 °C, 30 atm.

iii. Oxidizer: 720 °C, 30 atm.

iv. Reformer: 500 °C, 1 atm.

3. CLG 1 and CLG 2 syngas yields were 0.87 and 2.54 kmol / kmol PL, respectively. CLG 2

generated the most H2 in the product syngas stream, 1.60 kmol / kmol PL, based on absolute

78

yield, with CLG 1 producing only 0.73 kmol / kmol PL. H2 production was significantly

less than that outlined in the literature for both simulations.

4. CLG 1 produced purer syngas with an H2 concentration of 92.45 mol-%, while CLG 2 had

62.94 mol-% H2. The lower CLG 2 concentration was due to the presence of CO2 in that

syngas stream, as its removal was not a focus of that simulation. CLG 1 exhibited more H2

rich syngas than other studies while CLG 2 produced results similar to those found in the

literature.

5. Future research could focus on increasing the accuracy and scalability of the simulations

through assumption mitigation or removal. Examples include modelling biomass

gasification in a single gasifying reactor, detailed design of gas-solid separators, and

inclusion of nitrogen and sulfur elements in biomass ultimate analyses.

References

[1] Holladay, J., J. Hu, D. King, and Y. Wang. "An Overview of Hydrogen Production Technologies." Catal

Today 139.4 (2009): 244-60.

[2] Florin, N., and A. Harris. "Hydrogen Production from Biomass Coupled with Carbon Dioxide Capture: The

Implications of Thermodynamic Equilibrium." Int J Hydrogen Energ 32.17 (2007): 4119-134.

[3] Fan, Liang-Shih. Chemical Looping Systems for Fossil Energy Conversions. Hoboken, NJ: Wiley-AIChE,

2010. Print.

[4] Ni, M., D. Leung, M. Leung, and K. Sumathy. "An Overview of Hydrogen Production from Biomass." Fuel

Process Technol 87.5 (2006): 461-72.

[5] Kirtay, Elif. "Recent Advances in Production of Hydrogen from Biomass." Energ Convers Manage 52

(2011): 1778-789.

[6] Acharya, Bishnu. Chemical Looping Gasification of Biomass for Hydrogen-Enriched Gas Production.

Thesis. Dalhousie University, Halifax, Nova Scotia, 2011. Dalhousie University, Department of Mechanical

Engineering. Print.

[7] Acharya, B., A. Dutta, and P. Basu. "Chemical-Looping Gasification of Biomass for Hydrogen-Enriched Gas

Production with In-Process Carbon Dioxide Capture." Energ Fuels 23 (2009): 5077-083.

[8] Mahishi, Madhukar R., M. S. Sadrameli, Sanjay Vijayaraghavan, and D. Y. Goswami. "A Novel Approach

to Enhance the Hydrogen Yield of Biomass Gasification Using CO2 Sorbent." J Eng Gas Turb Power 130

(2008): 011501-1 to 11501-8. Print.

[9] Moghtaderi, B. "Effects of Controlling Parameters on Production of Hydrogen by Catalytic Steam

Gasification of Biomass at Low Temperatures." Fuel 86.15 (2007): 2422-430.

79

[10] Abuadala, A., I. Dincer, and G.F. Naterer. "Exergy Analysis of Hydrogen Production from Biomass

Gasification." Int J Hydrogen Energ 35 (2010): 4981-990.

[11] Gonzalez, J., S. Roman, D. Bragado, and M. Calderon. "Investigation on the Reactions Influencing Biomass

Air and Air/steam Gasification for Hydrogen Production." Fuel Process Technol 89.8 (2008): 764-72.

[12] Shen, L., Y. Gao, and J. Xiao. "Simulation of Hydrogen Production from Biomass Gasification in

Interconnected Fluidized Beds." Biomass Bioenerg 32.2 (2008): 120-27.

[13] Franco, C., F. Pinto, I. Gulyurtlu, and I. Cabrita. "The Study of Reactions Influencing the Biomass Steam

Gasification Process." Fuel 82.7 (2003): 835-42.

[14] Pfeifer, C., B. Puchner, and H. Hofbauer. "Comparison of Dual Fluidized Bed Steam Gasification of Biomass

with and without Selective Transport of CO2." Chem Eng Sci 64.23 (2009): 5073-083.

[15] Lin, Shiying, Michiaki Harada, Yoshizo Suzuki, and Hiroyuki Hatano. "Hydrogen Production from Coal by

Separating Carbon Dioxide during Gasification." Fuel 81 (2002): 2079-085.

[16] Hanaoka, T., Takahiro Yoshida, Shinji Fujimoto, Kenji Kamei, Michiaki Harada, Yoshizo Suzuki, Hiroyuki

Hatano, Shin-ya Yokoyama, and Tomoaki Minowa. "Hydrogen Production from Woody Biomass by Steam

Gasification Using a CO2 Sorbent." Biomass Bioenerg 28.1 (2005): 63-68.

[17] Guoxin, Hu, and Huang Hao. "Hydrogen Rich Fuel Gas Production by Gasification of Wet Biomass Using a

CO2 Sorbent." Biomass Bioenerg 33.5 (2009): 899-906.

[18] Mahishi, M., and D. Goswami. "An Experimental Study of Hydrogen Production by Gasification of Biomass

in the Presence of a CO2 Sorbent." Int J Hydrogen Energ 32.14 (2007): 2803-808.

[19] Acharya, Bishnu, Animesh Dutta, and Prabir Basu. "An Investigation into Steam Gasification of Biomass for

Hydrogen Enriched Gas Production in Presence of CaO." Int J Hydrogen Energ 35.4 (2010): 1582-589.

[20] Faaij, A., R. van Ree, L. Waldheim, E. Olsson, A. Oudhuis, A. van Wijk, et al. "Gasification of Biomass

Wastes and Residues for Electricity Production." Biomass Bioenerg 12.6 (1997): 387-407.

[21] Piroonlerkgul, P., W. Wiyaratn, A. Soottitantawat, W. Kiatkittipong, A. Arpornwichanop, N. Laosiripojana,

et al. "Operation Viability and Performance of Solid Oxide Fuel Cell Fuelled by Different Feeds." Chem Eng

J 155 (2009): 411-18.

80

Chapter V

Tubular Solid Oxide Fuel Cell Operation on Syngas from Two

Unique Biomass Chemical Looping Gasification Processes – A

Performance Comparison Simulated Using Aspen Plus

Nomenclature

Latin Alphabet

Ac Active cell area [m2] – Eqn. (17)

F Faraday’s constant [96 485 C mol-1] – Eqn. (15)

i Current density [mA cm-2] – Eqns. (12) and (17)

n Number of electrons transferred [mol e- per mol] – Eqn. (15)

Nc Number of cells – Eqn. (18)

�̇�𝑗 Molar flowrate of species j [kmol h-1] – Eqn. (16)

P Pressure [atm = 1.01325 bar] – Eqns. (10), (11), (13), and (14)

pc Cell power [W cell-1] – Eqns. (17) and (18)

ptot Total power output [kW = 1000 W] – Eqn. (18)

Q Thermal energy [kJ mol-1 = 1000 J mol-1] – Eqns. (3), (5), and (7)

T Temperature [°C] – Eqns. (10) and (12)

Uf Fuel utilization factor [-] – Eqn. (15)

V Voltage [V = 1000 mV] – Eqns. (10), (15), and (17)

W Work [kJ mol-1 = 1000 J mol-1] – Eqns. (3), (5), and (7)

Greek Alphabet

ΓH2 Fuel equivalent H2 content [kmol h-1] – Eqns. (15) and (16)

ΔV Voltage difference [V = 1000 mV] – Eqns. (11) to (14)

η Efficiency [-] – Eqn. (15)

Subscripts

an Anode

c Cell

cath Cathode

e Electric

op Operating conditions

P Pressure

ref Reference conditions

T Temperature

tot Total

Acronyms

CLG Chemical looping gasification

LHV Fuel lower heating value [kJ mol-1 = 1000 J mol-1] – Eqn. (15)

NC No CO2

SOFC Solid oxide fuel cell

WC With CO2

81


The use of hydrogen (H2) as an alternative to fossil fuel energy sources has the potential to shift

the global search for cleaner, renewable forms of energy towards more efficient green technologies

by mitigating reliance on carbonaceous fuels. The major issue of global climate change resulting

from rapidly increasing atmospheric concentrations of carbon dioxide (CO2) and other greenhouse

gases (GHGs) can be addressed via the utilization of various H2 production and utilization

technologies. H2 is also beneficial when compared to conventional alternative energies such as

wind and solar power due to its effective capabilities as a medium for both energy storage and

transport; however both of these beneficial aspects are still considerable challenges. Consequently,

the widespread use of H2 as an energy source in the near future is highly probable.

Contemporary H2 production methods ultimately contribute to the consumption and

subsequent generation of GHGs [1]. Further, these methods are neither GHG-neutral nor

sustainable since their designs are based on the use of non-renewable fossil fuels. Therefore, the

production of H2 from a renewable, carbon-neutral energy source is of paramount importance.

Chemical looping gasification (CLG) is exemplar of such technologies and is a thermochemical

conversion process that is a feasible candidate for the replacement of fossil fuel-based H2

production. It involves the conversion of biomass fuel to synthesis gas (or syngas), mainly H2,

carbon monoxide (CO), and various hydrocarbons, for subsequent use in H2 conversion

technologies for the purpose of energy generation.

Poultry litter is an example of such biomass fuels and is both globally abundant and readily

available. The global poultry population in 2012 exceeded twenty-one billion chickens, or more

than three chickens per person on Earth [2]. The utilization of poultry excrement wastes in biomass

CLG systems for the production of high purity H2 syngas would provide a plentiful fuel source for

small-to-medium scale power generation via SOFC operation. It would also provide a critical link

between the green energy fields of biomass gasification and fuel cell power.

5.1.1 Solid Oxide Fuel Cells

Solid oxide fuel cells (SOFCs), high-temperature electrochemical devices which directly convert

chemical energy into electrical energy, are one such H2 utilization technology. Typical operating

temperatures range from 600 to 1000 °C. The cell assembly for SOFCs generally consists of an

anode, cathode, electrolyte, and interconnect, with all portions of the cell existing as solid-state,

82

thus requiring the higher operating temperature. Generally, the anode material is constructed of a

nickel and yttria-stabilized zirconia (Ni-YSZ) cermet (ceramic-metal), the cathode is strontium

(Sr)-doped lanthanum manganate (LaMnO3), and the electrolyte is YSZ, however these materials

may be altered or switched out entirely depending on the SOFC application [3-7].

A schematic of typical SOFC operation on H2 and CO can be seen in Figure 5.1 [8].

Electrochemical oxidation of the active fuel occurs at the cell anode and the reduction of oxygen

(O2) occurs at the cathode. Equation (1) shows the former for H2 fuel and Equation (2) shows the

latter. The overall reaction for the process can be seen in Equation (3).

Anodic H2 oxidation: H2 + O= → H2O + 2e- (1)

Cathodic O2 reduction: ½ O2 + 2e- → O= (2)

Overall reaction (H2): H2 + ½ O2 → H2O + We + Q (3)

Fig. 5.1. SOFC operation on H2 and CO fuels [8].

SOFCs may also operating using CO and methane (CH4) directly. The electrochemical reaction

for the cathodic reduction of O2 remains unchanged from H2 operation in either case, except for

multiples of both sides of the equation in the case of CH4. The direct oxidation of CO at the anode

and its corresponding overall reaction can be seen in Equations (4) and (5), respectively.

Anodic CO oxidation: CO + O= → CO2 + 2e- (4)

Overall reaction (CO): CO + ½ O2 → CO2 + We + Q (5)

83

In addition, the direct anodic oxidation of CH4 and its respective overall reaction can be seen in

Equations (6) and (7) [3].

Anodic CH4 oxidation: CH4 + 4 O= → CO2 + 2 H2O + 8e- (6)

Overall reaction (CH4): CH4 + 2 O2 → CO2 + 2 H2O + We + Q (7)

The advantages of SOFCs over other fuel cell technologies stem mainly from the geometric design

of the cell assembly and the higher operating temperatures. These advantages include the use of

non-precious metal catalysts, higher efficiency and longer cell life than other fuel cell types, high

fuel flexibility, and the use of a solid-state electrolyte. The electrochemical kinetics of the half-

cell reactions are faster than other fuel cell types as a result of the higher operating temperatures

[9,10]. This allows Ni to be used as a catalyst for anodic oxidation rather than precious metal

catalysts.

SOFC operation is possible using H2, CO, and CH4, as well as higher-order hydrocarbons

such as ethane (C2H6) [11,12], propane (C3H8), n-butane (C4H10) [11], etc. Generally, CO and CH4

are reformed to H2 and CO2 prior to reaching the SOFC anode and undergoing direct oxidation.

Pre-reforming of CO occurs via the water-gas shift reaction which can be seen in Equation (8).

Water-gas shift reaction: CO + H2O → H2 + CO2 (8)

The H2 generated in Equation (8) is ultimately oxidized at the anode and undergoes the process

outlined by Equations (1) to (3). Similarly, pre-reforming of CH4 occurs via the endothermic steam

reforming of methane reaction and can be seen in Equation (9).

Steam reforming of CH4: CH4 + H2O → CO + 3 H2 (9)

The CO generated in Equation (9) undergoes the water-gas shift reaction illustrated in Equation

(8), while the H2 produced proceeds to be oxidized as shown in Equations (1) to (3).

Multicomponent inlet fuel streams consisting of active fuels diluted with inert or inactive gases,

i.e. nitrogen (N2), CO2, and steam (H2O), can also be used in SOFCs, however overall performance

generally decreases as diluent gas concentration increases [11,13].

The use of a non-fluid electrolyte allows for multiple geometries for the SOFC electrolyte

and overall cell design. These include the planar (bipolar or flat plate) and tubular configurations;

both of which are depicted in Figure 5.2 [8]. The flat plate configuration is relatively simple

84

conceptually, and consists of the cell components in a series connection, however issues of

improper gas sealing and fabrication of thin layer structures arise. The flexibility of the solid-state

electrolyte allows for the tubular configuration, which was developed by Siemens-Westinghouse

[14,15]. This SOFC geometry simplifies gas sealing and individual cells are easily connected by

attachment to a common support tube [8]. The performance and operation of the tubular design is

well studied and will be the SOFC configuration used in this paper [16-21].

Fig. 5.2. Tubular and flat plate SOFC configurations [8].

In contrast, some drawbacks of SOFCs remain prominent and include material durability

issues and inlet fuel stream type or composition. Degradation of the solid-state cell assembly

results from prolonged exposure to the higher operating temperatures. Difficulties arise in seeking

out alternative materials that can withstand such high temperatures for extended periods of time.

Furthermore, the possible reduction-oxidation (redox) cycling during operation suggests that Ni is

constantly often between Ni and nickel oxide (NiO). These conversion cycles induce anode

volume changes during operation which have a detrimental effect on the structural integrity of the

anode, and consequently to performance degradation overall [9,10,22-24].

Reduction in cell performance also stems from the Ni-catalyzed thermal cracking of

hydrocarbons present in the inlet fuel stream. By-product graphite formation is generated as a

result and leads to carbon deposition at the anode, thereby occupying anode active sites and

inhibiting H2 adsorption [9,10]. Additionally, SOFCs are highly sensitive to sulfur poisoning in

the form of hydrogen sulfide (H2S), which significantly reduces cell performance. Successful

operation of SOFCs is often jeopardized by even small, parts-per-million (ppm) levels of H2S.

Corresponding research has therefore been conducted to identify the mechanisms contributing to

performance degradation in the presence of H2S, and to develop gas cleaning techniques for sulfur

removal to ensure sufficient operational longevity and performance [9,10,25].

85

5.1.2 Syngas from Biomass CLG and SOFC Operation

Direct operation of SOFCs on product syngas from the CLG of biomass is advantageous since the

inlet fuel stream can be fed directly to a SOFC with minimal pre-operation gas cleaning

requirements. Therefore, the research compares the operation of a tubular SOFC for two different

syngas streams, each resulting from unique biomass CLG processes for the purpose of H2

production. The first process (CLG-1) produced H2-rich syngas using a calcium oxide (CaO)

sorbent for CO2 capture with total sorbent recovery. The second process (CLG-2) used iron (Fe)-

based oxygen carriers to produce majority-H2 syngas. Poultry litter was used as the biomass

feedstock in both cases. More information regarding the two syngas production processes can be

found in a previous study conducted by the authors [26].

Performance comparisons are measured in terms of cell voltage, electrical efficiency, and

total power output for each syngas type. The effects of varying anode operating temperature and

pressure, inlet fuel utilization factor, and applied current density on overall cell performance are

investigated. The originality of the conducted research lies in the investigation of the effects of CO

and CO2 syngas feed composition on tubular SOFC performance and in the direct performance

comparison of multiple biomass CLG syngas types modelled in a SOFC under similar feed and

operating conditions simulated using the Aspen Plus software. The SOFC model developed in

ASPEN Plus was adapted from Zhang et al. [16].

5.2 Simulation Description and Input Parameters

The following sections provide a detailed description of the simulated process. In addition, the

input data to the Aspen Plus simulation engine are outlined, as well as the chosen and utilized

calculation methods and equations.

5.2.1 Simulated Process

A tubular SOFC was simulated and analyzed using Aspen Plus V8.0 by following the design

presented by Zhang et al. [16] as a guide. A conceptual block diagram of the process can be seen

in Figure 5.3 and the simulation flowsheet can be seen in Figure 5.4. The input and operating

conditions for all feed streams and block units can be found in Tables 5.1 and 5.2, respectively.

CLG-1 incorporated CO2 removal while CLG-2 did not. Modified versions of both

processes were run in the simulation to determine the effect of CO2 removal from the syngas stream

on SOFC performance. Modified CLG-1 syngas was similar to regular CLG-1 syngas, however

86

CO2 removal from the stream was not implemented in the former. Modified CLG-2 syngas was

similar to regular CLG-2 syngas, though CO2 removal was applied to the stream. The following

labels were used to differentiate between the processes: regular CLG-1 was denoted as CLG-1-

NC, modified CLG-1 was labelled CLG-1-WC, regular CLG-2 was CLG-2-WC, and modified

CLG-2 was CLG-2-NC. Note that NC = “No CO2” and WC = “With CO2”.

The effect of sulfur poisoning on SOFC performance was not studied since the designed

model did not accurately simulate the mechanism by which contaminant H2S inhibits anodic H2

oxidation. This is a result of the simulation engine bypassing Equation (1) to effectively model

Equation (3). Furthermore, no sulfur compounds were present in any tested syngas type [26].

Fig. 5.3. SOFC simulation block diagram.

87

Fig. 5.4. SOFC simulation flowsheet.

Table 5.1

Feed stream input conditions.

Condition Air Syngas

(CLG-1-NC)

Syngas

(CLG-1-WC)

Syngas

(CLG-2-WC)

Syngas

(CLG-2-NC)

Temperature (°C) 630 20 20 20 20

Pressure (atm) 1 10 10 10 10

Flowrate (kmol h-1) 40 a 0.87 1.68 2.54 1.66

H2 (mol-%) - 92.46 45.78 62.49 97.49

H2O (mol-%) - 0.23 0.23 2.30 2.31

CO (mol-%) - 1.28 45.16 0.62 0.00

CO2 (mol-%) - 0.01 7.44 34.11 0.19

CH4 (mol-%) - 6.03 1.40 0.03 0.00

N2 (mol-%) 79 - - - -

O2 (mol-%) 21 - - - - a Fed in excess of required stoichiometric amount.

88

Table 5.2

Block unit operating conditions.

Block

Name

Block

Type

Temperature

(°C)

Pressure

(atm) Other

Anode RGibbs 950 a 1 -

Cathode Sep - - Separated a specified molar fraction of O2 from heated

inlet air stream (CATH-IN) as feed to ANODE. b

Burner RStoich 950 1

Output based on defined reactions governed by

stoichiometric input, i.e.:

CO + ½ O2 CO2

H2 + ½ O2 H2O

Heat-1 Heater 950 1 -

Heat-2 Heater 910 1 -

Heat-3 Heater 1012.35 1 -

Heat-X HeatX - -

Flow direction: Countercurrent. Type: Design.

Specification: Hot/cold outlet temperature approach.

Value: 10 delta-C. a Represents design operating temperature. Anode temperature was varied from 900 to 1000 °C. b Molar split fraction for O2 was based on roughly half of the theoretical inlet H2 molar flowrate in the syngas feed,

and differed based on utilized syngas type. A split fraction of 6 mol-% O2 was used for CLG-1-NC, while 9.5 mol-%

O2 was used for the other syngas types.

5.2.2 Setup and Component Definition

The Aspen Plus flowsheet type was chosen as “General with metric units”, allowing for the

analysis and results presentation for fluid and aqueous (MIXED) input and output material streams.

The process type was chosen as COMMON, allotting a generic industry type the simulation. The

IDEAL base calculation method was selected for simplicity and ease of operation. Thus, phase

equilibrium calculations were conducted using Raoult’s Law, Henry’s Law, the Ideal Gas Law,

etc.

Moreover, fluid streams were modelled using conventional components which have

thermophysical data stored in Aspen Plus databanks. Thus, no physical property data input were

required for fluid components. These components include: H2, H2O, CO, CO2, CH4, O2, and N2.

5.2.3 Flowsheet Description

Syngas from a biomass CLG process was heated to 950 °C in the HEAT-1 block unit. The Heater

is an Aspen Plus temperature changing module [16]. The heated syngas stream (GAS-WARM)

was then fed to the ANODE block. Simultaneously, hot air (AIR) was heated by being fed to the

cold fluid side of the HeatX (labelled as HEAT-X) heat exchanger module prior to being fed to the

CATHODE block unit, a Sep separation module which allowed for the direct splitting of inlet

streams into multiple outlet streams. Equation (2) did not proceed in the simulated cathode since

89

Aspen Plus does not easily model electrochemical reactions Thus, the CATHODE separated a

defined molar fraction of O2 from the heated air stream as feed to the ANODE block. This process

is equivalent to oxide ion (O=) migration from the cathode to the anode. The CATHODE sent

sufficient O2 to the ANODE for Equation (3) to proceed. The O2-depleted air (AIR-DEPL) was

then heated in the HEAT-2 unit before being sent to the BURNER block unit.

The ANODE block is an RGibbs reactor module which calculates its output using the Gibbs

free energy minimization method, dependent on the defined operating temperature and pressure.

The calculations were based on the chemical equilibrium reactions of the components being input

to the reactor under the pre-defined conditions. Here, all reacted H2 and O2 underwent Equation

(3), the overall SOFC equation when operating on H2. Equations (8) and (9) were assumed to have

much faster kinetics than Equations (4) and (6), respectively, and thus all reacted CO and CH4

were presumed to undergo conversion to H2 and CO2 [16-19,27].

The ANODE products (AN-PROD) were also sent to the BURNER block unit for

combustion at under sufficient O2 conditions provided by the heated BURN-IN stream. The

combustion unit is an isothermal RStoich reactor module which has outputs based on user-defined

chemical reactions. Residual CH4 combustion was not considered since its molar flowrate in the

AN-PROD stream was negligible for all tested syngas types. The combustion products (BURN-

OUT) were heated in the HEAT-3 block unit to represent the temperature increase following

combustion and ultimately sent to the hot fluid side of the HEAT-X unit to provide heat for the

inlet AIR stream. The combustion products then exit the HEAT-X block as the EXHAUST stream.

5.2.4 Cell Performance Calculation Methods

5.2.4.1 Cell Voltage

Semi-empirical equations were used to convert the inlet and outlet anodic molar flowrates to

voltages, and were based on reference operating conditions. The reference conditions can be found

in Table 5.3. The expression for cell voltage is shown by Equation (10) [14,16,28].

Cell voltage: Vc = Vref + ΔVP + ΔVT + ΔVcath + ΔVan (10)

The reference voltage is a function of current density [14] and accounts for detriments to SOFC

performance such as ohmic, activation, and concentration polarization losses. Table 5.4 outlines

reference voltage values for different current densities. The remainder of the terms in Equation

(10) account for contributions to the overall cell voltage by the simulated operating conditions.

90

Table 5.3

Reference conditions used in voltage calculations [14].

Reference Parameter Symbol Condition

Temperature Tref 1000 °C

Pressure Pref 1 bar

O2 Partial Pressure (PO2)ref 0.164 bar a

Ratio of H2 to H2O Partial Pressures (PH2/PH2O)ref 0.15 a

a From Zhang et al. [16].

Table 5.4

Reference voltage as a function of current density [14].

ic (mA cm-2) Vref (mV)

0 720

180 a 650 a

200 640

400 560

600 460

720 400 a Standard condition.

The effect of operating pressure, and its corresponding contribution to Vc, is summarized

in Equation (11) [14,16].

Operating pressure: ΔVP [mV] = 76 (P/Pref) (11)

Subsequently, the effect of operating temperature can be seen in Equation (12) [8,14,16].

Operating temperature: ΔVT [mV] = 0.008 (Top – Tref) ic (12)

Current density has units of milliamps per square centimetre (mA cm-2) in Equation (12).

Furthermore, the effect of deviation from the reference cathodic O2 partial pressure is illustrated

in Equation (13) [14,16].

Cathode composition: ΔVcath [mV] = 92 log[PO2/(PO2

)ref] (13)

Finally, the expression for the effect of deviation from the reference anodic H2-to-H2O partial

pressure ratio is shown in Equation (14) [14,16].

Anode composition: ΔVan [mV] = 172 log[(PH2/PH2O)/(PH2

/PH2O)ref] (14)

H2 and H2O partial pressures were calculated as averages of anode inlet and outlet (SYNGAS and

AN-PROD streams, respectively).

91

5.2.4.2 Electrical Efficiency

The electrical efficiency of the cell was calculated using Equation (15) [14,16].

Electrical efficiency: ηe = We/LHV = (nFVcUf ΓH2)/LHV (15)

The quantity ΓH2 represents the equivalent H2 flowrate in the inlet syngas stream, and is determined

according to Equation (16) [14,16]. LHV values can be found in Table 5.5 [29].

Equivalent H2 flowrate: ΓH2 [kmol h-1] = �̇�H2

+ �̇�CO + 4�̇�CH4 (16)

The coefficients in Equation (16) correspond to the contribution of each species to the equivalent

H2 flowrate in the syngas stream. For example, each mole of CO contributes 1 mol H2 via Equation

(8) and each mole of CH4 contributes 4 mol H2 via Equation (9).

Table 5.5

Lower heating value of fuels [29].

Fuel LHV (kJ mol-1) a

H2 242.57

CO 284.58

CH4 800.14 a Converted from averaged values in units of [MJ Nm-3].

5.2.4.3 Power Output

The calculation of power output per cell and total power output were based on the geometry of the

simulated SOFC. The corresponding formulae are outlined in Equations (17) and (18).

Power output per cell: pc = ic Ac Vc (17)

Total power output: ptot = pc Nc (18)

Thus, a design cell geometry was assumed with an active area of 96.1 m2 and a total of 1152 cells

based on SOFC performance data [16,17,20,21].

92

5.3 Results and Discussion

The following sections detail the data obtained from the simulation and corresponding analysis

and discussion. Table 5.6 outlines the standard operating conditions utilized for the data collection

process of the simulation. Simulated SOFC performance results were obtained and compared for

the effects of the presence of CO, the presence of CO2, temperature, pressure, utilization factor,

and current density.

5.3.1 Syngas Performance Comparison

The two syngas types (CLGs 1-NC and 2-WC) and their respective modified versions (CLGs 1-

WC and 2-NC) were run in the SOFC simulation to determine the differences in the resulting cell

performance. The cell voltage and electrical efficiency comparison can be seen in Figure 5.5 and

the total power output comparison is illustrated in Figure 5.6. Although certain syngas types

outperformed others, the range of values was not immense (< 100 mV for voltage, < 5% for

efficiency, and < 15 kW for total power output).

Table 5.6

Standard operating conditions.

Operating Parameter Symbol Condition

Temperature Top 950 °C

Pressure Pop 1 atm (1.01325 bar)

Fuel Utilization Factor Uf 85%

Current Density ic 180 mA cm-2

Reference Voltage Vref 650 mA

CLG-1-WC exhibited the greatest SOFC performance in terms of cell voltage at roughly

0.77 V, while CLG-1-NC displayed the greatest performance in terms electrical efficiency at about

51%. However, the measured efficiencies were similar in magnitude. In addition to cell voltage,

CLG-1-WC showed superior performance in terms of total power output and cell power (i.e. power

per cell), with corresponding values of 133 kW and 116 W cell-1. CLG-2-WC exhibited the lowest

performance for all measured parameters.

The simulation results are compared to the literature [16,17,20,21] in Table 5.7. The

obtained cell voltages for all syngas types were greater than or similar to those in research

conducted by Zhang et al. [16], Doherty et al. [17], Calí et al. [20], and Verda and Calí [21] under

similar operating current density conditions. This was also the case for electrical efficiency and

total power output. The stack exhaust compositions obtained for each syngas type were similar to

93

those reported in the literature with CLGs 1-NC and 2-NC having significantly smaller molar

concentrations of CO2. Thus, SOFC operation on these syngas types inherently reduce

contributions to atmospheric CO2 levels.

Fig. 5.5. Cell voltage and electrical efficiency comparison (Top = 950 °C, Pop = 1 atm, Uf = 85%, ic = 180 mA cm-2,

Vref = 650 mV, NC = No CO2, WC = With CO2).

Fig. 5.6. Total power output comparison (Top = 950 °C, Pop = 1 atm, Uf = 85%, ic = 180 mA cm-2, Vref = 650 mV,

NC = No CO2, WC = With CO2).

0.4

0.45

0.5

0.55

0.6

0.65

0.7

0.75

0.8

CLG-1-NC CLG-1-WC CLG-2-WC CLG-2-NC

Vo

ltag

e (V

) an

d E

ffic

iency

Cell Voltage Electrical Efficiency

116

118

120

122

124

126

128

130

132

134


Po

wer

(k

W)

94

Table 5.7

Comparison of simulation results under standard conditions to the literature [16,17,20,21].

Parameter Simulation Results Literature Values

CLG-1-NC CLG-1-WC CLG-2-WC CLG-2-NC [16] [17] [20,21]

Cell Voltage (V) 0.73 0.77 0.71 0.72 0.70 0.683 0.661

Current Density

(mA cm-2) 180 180 180 180 178 182.86 200.6

Anode Inlet Composition

(mol-%)

H2 (92.46),

H2O (0.23),

CO (1.28),

CO2 (0.01),

CH4 (6.03),

N2 (0.0)

H2 (45.78),

H2O (0.23),

CO (45.16),

CO2 (7.44),

CH4 (1.40),

N2 (0.0)

H2 (62.49),

H2O (2.30),

CO (0.62),

CO2 (34.11),

CH4 (0.03),

N2 (0.0)

H2 (97.49),

H2O (2.31),

CO (0.00),

CO2 (0.19),

CH4 (0.00),

N2 (0.0)

H2 (67.0),

H2O (11.0),

CO (22.0),

CO2 (0.0),

CH4 (0.0),

N2 (0.0)

H2 (26.9),

H2O (27.8),

CO (5.6),

CO2 (23.1),

CH4 (10.4),

N2 (6.2)

H2 (2.9),

H2O (27.4),

CO (8.3),

CO2 (52.8),

CH4 (7.4),

N2 (1.3) a

Anode Outlet

Composition (mol-%)

H2 (1.90),

H2O (91.57),

CO (0.19),

CO2 (6.34),

N2 (0.0)

H2 (0.62),

H2O (46.85),

CO (1.00),

CO2 (51.53),

N2 (0.0)

H2 (0.54),

H2O (64.72),

CO (0.38),

CO2 (34.37),

N2 (0.0)

H2 (1.39),

H2O (98.41),

CO (0.0),

CO2 (0.20),

N2 (0.0)

H2 (11.6),

H2O (50.9),

CO (7.4),

CO2 (24.9),

N2 (5.1)

H2 (11.6),

H2O (50.9),

CO (7.4),

CO2 (24.9),

N2 (5.1)

H2 (1.39),

H2O (39.88),

CO (11.91),

CO2 (45.88),

N2 (0.94) a

Stack Exhaust

Temperature (°C) 830.57 832.74 835.02 831.80 834 833.7 279

Stack Exhaust

Composition (mol-%)

N2 (78.09),

O2 (19.49),

H2O (2.26),

CO2 (0.16)

N2 (77.23),

O2 (18.54),

H2O (2.01),

CO2 (2.22)

N2 (75.72),

O2 (18.19),

H2O (3.98),

CO2 (2.12)

N2 (77.35),

O2 (18.58),

H2O (4.06),

CO2 (0.01)

N2 (77.3),

O2 (15.9),

H2O (4.5),

CO2 (2.3)

N2 (77.3),

O2 (15.9),

H2O (4.5),

CO2 (2.3)

N2 (75.62),

O2 (17.38),

H2O (3.14),

CO2 (3.87) a

Electrical Efficiency

(%, LHV-based) 50.90 48.57 47.89 48.56 52 49.15 48 b

Cell Power (W cell-1) 109.20 115.51 106.49 107.82 103.94 c 104.16 d 110.56 d

Total Power Output (kW) 125.79 133.07 122.68 124.21 120 120 127.4 a All composition values reported in units of [wt.-%], not [mol-%]. b Reported as gross AC efficiency using the fuel LHV. All other efficiency values are based on net AC efficiency using the fuel LHV. c Calculated using the reported active area of 96.1 m2 with 1152 cells and reported cell voltage and current density values. d Calculated using the reported active area of 96.0768 m2 with 1152 cells and reported cell voltage and current density values.

95

The differences between the simulation results and those reported in the literature may be

explained by differences in calculation methods and initial assumptions. For example, a power

output of 120 kW was initially assumed in the study conducted by Zhang et al. [16] and the

resultant current density of 178 mA cm-2 was calculated based on the obtained cell voltage. In

contrast, a similar current density of 180 mA cm-2 was initially assumed for all syngas types in this

paper and the total power output was then calculated based on the obtained cell voltages.

5.3.1.1 Effect of Syngas CO Composition

CO composition levels were varied from 0.01 to 60 mol-% while maintaining the balance syngas

at fixed proportions of the other constituents. This was done for both of the unmodified syngas

types (CLGs 1-NC and 2-WC) to determine the effects of CO composition on their respective

performance in the simulated SOFC. The results for cell voltage and electrical efficiency can be

found in Figure 5.7 while those for total power output are illustrated in Figure 5.8.

Fig. 5.7. Effect of CO syngas composition on cell voltage and electrical efficiency.

Increasing CO concentrations under standard conditions significantly increased

performance for voltage and power, which both raised by 13.83% and 20.23% over the tested

range for CLGs 1 and 2, respectively. Both performance metrics also continually increased in a

linear fashion, achieving values higher than that of standard CLG-1-WC at the higher end of the

CO composition scale. Changes in efficiency with increasing CO were not as prominent. CLG-1

efficiency rose to a local maximum at 5 mol-% CO, began to decrease, and gradually increasing

0.47

0.48

0.49

0.5

0.51

0.52

0.53

0.54

0.55

0.6

0.65

0.7

0.75

0.8

0.85

0.9

0% 10% 20% 30% 40% 50% 60%

Eff

icie

ncy

Vo

ltag

e (V

)

CO Composition (mol-%)

Voltage (CLG-1) Voltage (CLG-2)

Efficiency (CLG-1) Efficiency (CLG-2)

96

for an overall percent increase of 1.49%. In contrast, CLG-2 efficiency continually increased in a

non-linear manner, achieving an increase from 47.90 to 51.33% (percent difference of 7.16%).

Fig. 5.8. Effect of CO syngas composition on total power output.

The simulated results are in disagreement with experimental findings due to the mechanism

by which CO is utilized in the simulated SOFC. Literature data suggest that increasing CO content

decreases SOFC performance [9-11]. The inconsistency is attributed to the assumption of complete

CO water-gas shift reforming to H2 via Equation (8). Experimental trials conducted in the

literature, however, included both reforming and direct anodic oxidation of CO via Equation (4).

The oxidized CO was unable to undergo conversion to H2 and ultimately be utilized in the cell.

The active CO fuel in the simulation was sufficient to enhance cell performance and

simultaneously counterbalance the detrimental effects of any inactive or inert species, as is the

case with the relatively high prevalence of CO2 in CLG-1-WC syngas.

5.3.1.2 Effect of Syngas CO2 Composition

CLGs 1-NC and 2-WC differed in-part due to their inherent CO2 removal aspects, of which only

the former of the two utilized. Their respective modified versions (CLGs 1-WC and 2-NC),

however, were considered to determine any benefits or drawbacks related to altering the designed

processes to SOFC cell performance under standard conditions, and thus correspondingly excluded

and included CO2 removal steps. The effects of varying CO2 composition in the syngas feed stream

120

125

130

135

140

145

150

155

0% 10% 20% 30% 40% 50% 60%

Po

wer

(kW

)

CO Composition (mol-%)

Power (CLG-1) Power (CLG-2)

97

on voltage, efficiency, and total power output were also considered for the unmodified syngas

types and can be seen in Figures 5.9 and 5.10, respectively.

Fig. 5.9. Effect of CO2 syngas composition on cell voltage and electrical efficiency.

Fig. 5.10. Effect of CO2 syngas composition on total power output.

CO2 composition levels from 0.01 to 60 mol-% were simulated for both of the unmodified

syngas types while maintaining the balance syngas at fixed proportions of the other constituents.

0.46

0.47

0.48

0.49

0.5

0.51

0.52

0.53

0.54

0.6

0.65

0.7

0.75

0.8

0.85

0.9

0% 10% 20% 30% 40% 50% 60%

Eff

icie

ncy

Vo

ltag

e (V

)

CO2 Composition (mol-%)

Voltage (CLG-1) Voltage (CLG-2)

Efficiency (CLG-1) Efficiency (CLG-2)

115

117

119

121

123

125

127

129

131

133

0% 10% 20% 30% 40% 50% 60%

Po

wer

(kW

)

CO2 Composition (mol-%)

Power (CLG-1) Power (CLG-2)

98

Increasing CO2 concentrations at standard conditions decreased performance for voltage,

efficiency, and total power output, all of which decreased with an approximate linear trend. The

performance metrics underwent a reduction of 5.78% for CLG-1 and only about 2.82% for CLG-

2 over the tested range. CLG-1 syngas exhibited relatively higher performance than that of CLG-

2 over the range tested as well.

Excluding the CO2 removal step from unmodified CLG-1 syngas improved SOFC cell

voltage and total power output by an absolute percent difference of 5.78%, while efficiency was

found to decrease by 4.58%. This is evident since CLG-1-WC outperformed the other syngas types

despite its elevated levels of CO2. An explanation for this phenomenon is that CLG-1-WC

exhibited much higher levels of CO than its unmodified counterpart (45.16 and 1.28 mol-%,

respectively) which accounted for its greater relative performance. Further, the decrease in

efficiency was likely attributed to the fact that ηe is inversely proportional to the fuel LHV, as

outlined by Equation (15), and CLG-1-NC had a much smaller inlet flowrate of CO, even while

maintaining slightly greater flowrates of both H2 and CH4. The LHV of CO is greater than that of

H2 and roughly one third of the value for CH4, as shown in Table 5.5 [29].

The presence of CO2 in CLG-2 syngas had the opposite effect of CLG-1, with voltage and

total power output all decreasing by a percent difference of 1.23%. However, the same trend as

CLG-1-NC was followed for efficiency which decreased by 1.37% in the presence of CO2.

Nevertheless, CO2 removal had very little effect on performance for CLG-2-WC. This was

expected since CO2 variation within the CO2 composition range for CLG-2 (0.01 to 35 mol-%)

showed a voltage decrease of only about 8 mV, as shown in Figure 5.9. The decrease in efficiency

may be explained by the coupled effects of CLGs 2-WC and 2-NC having similar fuel LHV and

ΓH2 values, with CLG-2-NC exhibiting slightly greater performance in terms of Vc, which ηe is

directly proportional to.

The results obtained regarding the presence of CO2 in feed syngas are in line with those

reported in the literature. Jiang and Virkar [13] demonstrated that increasing concentrations of CO2

in binary H2-CO2 systems significantly reduced SOFC performance in terms of cell voltage and

power output. Feed streams containing CO were found to behave similarly to H2 given that H2

composition remained above 50%, and water-gas shift played a significant role in SOFC operation

on CO [13]. This finding provides further evidence for CO being the main factor contributing to

the observed higher performance of CLG-1-WC syngas regardless of its high CO2 content.

99

5.3.2 Anode Temperature Sensitivity Analysis

The anode operating temperature for each syngas type was varied from 900 to 1000 °C to

determine the effects on cell performance under standard conditions. The results of the analysis

are displayed in Figure 5.11. An increase in cell voltage, electrical efficiency, and total power

output was observed with increasing temperature for all syngas types. Voltage increased by

roughly 144 mV for the four cases over the range tested, and efficiency and power exhibited

respective increases of 9% and 25 kW.

Fig. 5.11. Effect of anode operating temperature on syngas performance (Pop = 1 atm, Uf = 85%, ic = 180 mA cm-2,

Vref = 650 mV, NC = No CO2, WC = With CO2).

SOFC thermodynamics theoretically predict that performance generally declines with

increasing temperature [30]. However, in practical systems, SOFC performance tends to increase

with increasing temperature due to the mitigation of kinetic barriers posed by the involved

electrochemical reactions [10,12]. Thus, the observed trends in regards to anode temperature are

in line with the literature.

5.3.3 Anode Pressure Sensitivity Analysis

An anode operating pressure sensitivity analysis was also conducted for each syngas type under

standard conditions. However, simulated pressure variation had negligible effects on cell voltage,

electrical efficiency, and total power output. This was the case for all syngas types. This trend is

100

105

110

115

120

125

130

135

140

145

150

0.4

0.45

0.5

0.55

0.6

0.65

0.7

0.75

0.8

0.85

0.9

900 920 940 960 980 1000

Po

wer

(kW

)

Vo

ltag

e (V

) an

d E

ffic

iency

Anode Temperature (°C)

Volt. (1-NC) Volt. (1-WC) Volt. (2-WC) Volt. (2-NC)

Eff. (1-NC) Eff. (1-WC) Eff. (2-WC) Eff. (2-NC)

Pwr. (1-NC) Pwr. (1-WC) Pwr. (2-WC) Pwr. (2-NC)

100

contrary to the theoretical thermodynamic effects of pressure on overall SOFC cell operation since

increasing pressure should increase cell voltage [30].

The discrepancy between the results of the simulation and real-world operation can be

explained by the limitations of the Aspen Plus RGibbs reactor type chosen to model the cell anode.

The software calculation engine determined that pressure effects on Equation (3) were negligible

in comparison to the effects of the high operating temperature at 950 °C, which remained constant

throughout pressure variation. No variation in the performance parameters was apparent as a result.

5.3.4 Fuel Utilization Factor Sensitivity Analysis

The fuel utilization factor was varied from 50 to 90% for each syngas type to determine the effects

on SOFC cell voltage, electrical efficiency, and total power output. The results for voltage and

efficiency can be seen in Figure 5.12 and those for total power output are shown in Figure 5.13.

Fig. 5.12. Effect of utilization factor on cell voltage and electrical efficiency (Top = 950 °C, Pop = 1 atm, ic = 180 mA

cm-2, Vref = 650 mV, NC = No CO2, WC = With CO2).

Cell voltage underwent a slight parabolic increase for all syngas types over the range of

utilization factors tested, with CLG-1-WC again exhibiting the greatest cell voltage overall. Cell

voltage for all syngas types continually increased with Uf over the tested range. Efficiency

increased linearly with increasing Uf for all syngas types. CLG-1-NC once more exhibited the

highest efficiency, averaging about 2.20% higher than the other syngas types over the tested range.

0.2

0.25

0.3

0.35

0.4

0.45

0.5

0.55

0.6

0.68

0.69

0.7

0.71

0.72

0.73

0.74

0.75

0.76

0.77

0.78

0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9

Eff

icie

ncy

Vo

ltag

e (V

)

Utilization Factor



101

Moreover, Figure 5.13 shows that total power output increased with Uf. This trend occurred as

expected due to the utilized calculation method for each performance metric; i.e. Equation (18) is

directly proportional to Vc and thus the power measurements differed from cell voltage only by

constant coefficients.

Fig. 5.13. Effect of utilization factor on total power output (Top = 950 °C, Pop = 1 atm, ic = 180 mA cm-2, Vref = 650

mV, NC = No CO2, WC = With CO2).

Upon comparison of the simulated results to the literature, it was found that the slight

parabolic trend observed for the effect of Uf on cell voltage was in line with that of the sensitivity

analysis conducted by Doherty et al. [17], yet decreasing cell voltage was observed over the entire

tested Uf range in the analysis conducted by Zhang et al. [16]. The linear increase in electrical

efficiency was also observed by both Doherty et al. [17] and Zhang et al. [16] over the 50 to 90%

tested Uf range.

5.3.5 Current Density Sensitivity Analysis

The current density applied to the simulated SOFC stack was varied from 0 to 720 mA cm-2 to

determine the effects on cell voltage, electrical efficiency, and total power output. The results for

voltage and efficiency are illustrated in Figure 5.14 and those for total power output are in Figure

5.15.

Cell voltage decreased linearly by about 0.61 V over the tested range of current densities

for all syngas types. CLG-1-WC once more achieved the overall highest voltages. Electrical

118

120

122

124

126

128

130

132

134

0.5 0.55 0.6 0.65 0.7 0.75 0.8 0.85 0.9

Po

wer

(kW

)

Utilization Factor


102

efficiency also decreased linearly, dropping by roughly 40% for all syngas types (over 66% in

terms of percent difference). CLG-1-NC, again having the highest efficiency, was no exception to

this trend. Total power output increased over the tested range until reaching a local maximum

value upwards of about 200 kW at 550 mA cm-2 prior to decreasing at higher values of ic. This

trend occurred despite the notable reduction in cell voltage.

Fig. 5.14. Effect of current density on cell voltage and electrical efficiency (Top = 950 °C, Pop = 1 atm, Uf = 85%,

NC = No CO2, WC = With CO2).

Fig. 5.15. Effect of current density on total power output (Top = 950 °C, Pop = 1 atm, Uf = 85%, NC = No CO2,

WC = With CO2).

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 100 200 300 400 500 600 700 800

Eff

icie

ncy

Vo

ltag

e (V

)

Current Density (mA/cm2)



0

50

100

150

200

250

0 100 200 300 400 500 600 700 800

Po

wer

(kW

)

Current Density (mA/cm2)


103

The resulting trends from the simulation were comparable to those outlined in the literature.

Both Zhang et al. [16] and Doherty et al. [17] observed decreasing cell voltage and electrical

efficiency over their respective tested ranges of current density. The decrease was linear or

approximately so in both cases, as was the case for the simulated results presented in this paper.

However, the decrease in efficiency found in reference [17] was far more pronounced than either

the simulated results or reference [16] in terms of percent difference, decreasing by more than 80%

and over a narrower range of current densities (50 to 450 mA cm-2). Though, the absolute drop of

about 40% closely matched with the results obtained in this paper, as efficiency declined from

about 60% to roughly 20% over the tested range of current densities. The curvature of the power

curves in the simulated results agreed with the shape of the power density curve obtained by

Doherty et al. [17], although the latter curve rose to a local maximum of about 130 kW at 325 mA

cm-2. For comparison, the tested syngas types valued power outputs upward of 175 kW at the same

current density. Zhang et al. [16] determined a linear increase in power, however that study varied

ic over a smaller range (160 to 240 mA cm-2). Near-linearity was observed for that portion of the

power curves in this paper, and thus the results are in agreement with reference [16].

The minor discrepancies between the simulated results and those reported by Doherty et

al. [17] may again be attributed to differences in calculation methods and simulation inputs and

assumptions. For example, reference [17] used different equations than the ones used in this paper

to calculate certain performance metrics such as electrical efficiency and voltage loss due to ohmic,

activation, and concentration polarization.

5.4 Potential for Future Research

The potential for future research stems in-part from the calculation method used to convert anode

outlet composition and flowrates to electrochemical SOFC performance measurements.

Determining a method to utilize external programming software linked to Aspen Plus via user-

defined subroutines to allow for modelling of the electrochemical half-cell reactions is an example

of this. Such action would mitigate some of the limitations posed by the model, such as the inability

to detect anode pressure variation and polarization effects associated with the presence of sulfur

in the feed syngas stream.

Future research may focus on a subsequent user-defined subroutine detailing the

mechanism by which H2S occupies Ni-YSZ active sites and inhibits anodic H2 oxidation. Linking

104

this code to the SOFC model would allow for more effective prediction of the effects of sulfur

poisoning and therefore any requirements for pre-operation gas cleaning and sulfur removal.

A thermal analysis of the SOFC model would allow for a better understanding of the energy

requirements and the balance of plant for the system. Such an analysis could be based on both the

process as a whole and on the individual Aspen Plus block units in the flowsheet, and has

implications for studying the effects of varying inlet feed stream temperatures.

Finally, a cost analysis could be conducted on the system, both as a whole and on the

individual block units. The analysis could be based on the ascertained energy requirements and

feed stream raw materials supply costs. Consequently, knowledge of cost estimates would provide

an idea of the feasibility of scaling up the designed tubular SOFC system, from a practical

standpoint, for real-life implementation.


In conclusion:

1. Syngas from the two biomass CLG processes operated well in the simulated SOFC and

their respective performance values were comparable to those reported in the literature.

CLG-1 syngas exhibited higher performance and thus that process is favourable to CLG-2

for power generation via SOFC operation.

2. CLG-1-WC syngas exhibited the greatest SOFC performance in terms of cell voltage, total

power output, and cell power, achieving respective values of 0.77 V, 133.01 kW, and

115.51 W cell-1. CLG-2-WC exhibited the lowest performance.

3. Increasing CO composition significantly improved cell voltage and total power output, but

had less of an effect on efficiency for both CLGs 1 and 2.

4. Increasing CO2 composition reduced performance overall. While almost no effect was

observed for CLG-2 syngas in the presence of CO2, large amounts of CO in CLG-1 syngas

allowed its performance to increase despite the presence of CO2. CO2 removal from the

CLG-1 process is not favourable for power generation via SOFC operation.

5. Increasing anode operating temperature increased performance parameters for each syngas

type, as was the case in the literature. Anode operating pressure variation had negligible

effects on performance due to limitations of the simulation software.

105

6. Cell voltage and total power output increased with increasing fuel utilization factor for all

syngas types, with efficiency increasing linearly. Voltage and efficiency decreased with

increasing current density, while total power output increased regardless of the drop in cell

voltage. The simulated utilization factor and current density sensitivities were comparable

to published trends.

7. Future research could focus on the development of user-defined subroutines linked to

Aspen Plus to model electrode half-cell reactions and sulfur poisoning mechanisms at the

anode. Thermal and economic analyses of the SOFC system would allow for improved

simulation-to-real-world predictability.

References

[1] Levin, David B., and Richard Chahine. "Challenges for Renewable Hydrogen Production from Biomass." Int J

Hydrogen Energ 35 (2009): 4962-969.

[2] Food and Agriculture Organization of the United Nations. “FAOSTAT – Production – Live Animals.” FAOSTAT.

7 Feb. 2014. Web. 26 Mar. 2014. <http://faostat.fao.org/site/573/DesktopDefault.aspx?PageID=573#ancor>.

[3] Park, S., J.M. Vohs, and R.J. Gorte. “Direct Oxidation of Hydrocarbons in a Solid-Oxide Fuel Cell.” Nature 404

(2000): 265-67.

[4] Vohs, J.M., R.J. Gorte, and S. McIntosh. "Role of Hydrocarbon Deposits in the Enhanced Performance of Direct-

Oxidation SOFCs." J Electrochem Soc 150.4 (2003): A470-476.

[5] Vohs, J.M., O. Costa-Nunes, and R.J. Gorte. “Comparison of the Performance of Cu-CeO2-YSZ and Ni-YSZ

Composite SOFC Anodes with H2, CO, and Syngas.” J Power Sources 141 (2005): 241-49.

[6] Wang, S.R., Xiao-Feng Ye, J. Zhou, F.R. Zeng, H.W. Nie, and T.L. Wen. "Assessment of the Performance of Ni-

yttria-stabilized Zirconia Anodes in Anode-supported Solid Oxide Fuel Cells Operating on H2–CO Syngas

Fuels." J Power Sources 195.21 (2010): 7264-267.

[7] Panopoulos, K.D., Hofmann, Ph., P.V. Aravind, M. Siedlecki, A. Schweiger, J. Karl, J.P. Ouweltjes, and E.

Kakaras. "Operation of Solid Oxide Fuel Cell on Biomass Product Gas with Tar Levels 10 g Nm−3." Int J

Hydrogen Energ 34 (2009): 9203-212.

[8] EG&G Services Parson, Inc. “Solid Oxide Fuel Cell.” Fuel Cell Handbook. 5th ed. Morgantown, WV: U.S. Dept.

of Energy, Office of Fossil Energy, National Energy Technology Laboratory, 2000. Pp. 8-1 to 8-22.

[9] Minh, N.Q. “Solid Oxide Fuel Cell Technology – Features and Applications.” Solid State Ionics 174 (2004): 271-

77.

[10] Barnett, S., Lin, Y., Z. Zhan, and J. Liu. "Direct Operation of Solid Oxide Fuel Cells with Methane Fuel." Solid

State Ionics 176.23-24 (2005): 1827-835.

[11] Eguchi, K., H. Kobo, T. Takeguchi, R. Kikuchi, and K. Sasaki. "Fuel Flexibility in Power Generation by Solid

Oxide Fuel Cells." Solid State Ionics 152-153 (2002): 411-16.

[12] Barnett, S., and Madsen, B. "Effect of Fuel Composition on the Performance of Ceramic-based Solid Oxide Fuel

Cell Anodes." Solid State Ionics 176.35-36 (2005): 2545-553.

106

[13] Jiang, Yi, and Anil V. Virkar. "Fuel Composition and Diluent Effect on Gas Transport and Performance of Anode-

Supported SOFCs." J Electrochem Soc 150.7 (2003): A942-951.

[14] Campanari, S. “Thermodynamic Model and Parametric Analysis of a Tubular SOFC Module.” J Power Sources

92 (2001): 26-34.

[15] Singhal, S.C. “Advances in Solid Oxide Fuel Cell Technology.” Proceedings of the 1998 Fuel Cell Seminar.

Courtesy Associates, November 1998.

[16] Zhang, W., E. Croiset, P.L. Douglas, M.W. Fowler, and E. Entchev. “Simulation of a Tubular Solid Oxide Fuel

Cell Stack using AspenPlusTM Unit Operation Models.” Energ Convers Manage 46 (2005): 181-196.

[17] Doherty, W., A. Reynolds, and D. Kennedy. “Computer Simulation of a Biomass Gasification-Solid Oxide Fuel

Cell Power System Using Aspen Plus.” Energy 35 (2010): 4545-4555.

[18] Panopoulos, K.D., L.E. Fryda, J. Karl, S. Poulou, and E. Kakaras. “High Temperature Solid Oxide Fuel Cell

Integrated with Novel Allothermal Biomass Gasification: Part I: Modelling and Feasibility Study.” J Power

Sources 159 (2006): 570-585.

[19] Hofmann, P., K.D. Panopoulos, L.E. Fryda, and E. Karakas. “Comparison Between Two Methane Reforming

Models Applied to a Quasi-Two-Dimensional Planar Solid Oxide Fuel Cell Model.” Energy 34 (2009): 2151-

2157.

[20] Calí, M., M.G.L. Santarelli, and P. Leone. “Design of Experiments for Fitting Regression Models on the Tubular

SOFC CHP 100 kWe: Screening Test, Response Surface Analysis and Optimization.” Int J Hydrogen Energ 32

(2007): 343-358.

[21] Verda, V. and M. Calí Quaglia. “Solid Oxide Fuel Cell Systems for Distributed Power Generation and

Cogeneration.” Int J Hydrogen Energ 33 (2008): 2087-2096.

[22] He, Hongpeng, and Josephine M. Hill. “Carbon Deposition on Ni/YSZ Composites Exposed to Humidified

Methane.” Appl Catal A-Gen 317 (2007): 284-292.

[23] Alzate-Restrepo, Vanesa, and Josephine M. Hill. “Effect of Anodic Polarization on Carbon Deposition on Ni/YSZ

Anodes Exposed to Methane.” Appl Catal A-Gen 342 (2008): 49-55.

[24] Alzate-Restrepo, Vanesa, and Josephine M. Hill. “Carbon Deposition on Ni/YSZ Anodes Exposed to CO/H2

Feeds.” J Power Sources 195 (2010): 1344-1351.

[25] Matsuzaki, Yoshio, and Isamu Yasuda. “The Poisoning Effect of Sulfur-Containing Impurity Gas on a SOFC

Anode: Part I. Dependence on Temperature, Time and Impurity Concentration.” Solid State Ionics 132 (2000):

261-69.

[26] Gopaul S.G., A. Dutta, and R. Clemmer. “Chemical Looping Gasification for Hydrogen Production: A

Comparison of Two Unique Processes Simulated Using ASPEN Plus.” Int J Hydrogen Energ 39 (2014): 5804-

5817.

[27] Costamagna, P., L. Magistri, and A.F. Massardo. “Design and Part-Load Performance of a Hybrid System Based

on a Solid Oxide Fuel Cell Reactor and a Micro Gas Turbine.” J Power Sources 96 (2001): 353-368.

[28] Kuchonthara, P., S. Bhattacharya, and A. Tsutsumi. “Energy Recuperation in Solid Oxide Fuel Cell (SOFC) and

Gas Turbine (GT) Combined System.” J Power Sources 117 (2003): 7-13.

[29] Waldheim, Lars, and Torbjörn Nilsson. “Heating Value of Gases from Biomass Gasification.” IEA Bioenergy

Agreement, Task 20 – Thermal Gasification of Biomass, TPS Termiska Processor AB, May 2001.

[30] Li, X. Principles of Fuel Cells. New York: Taylor & Francis, 2006. Print.

107

Chapter VI

Integrated CLG-SOFC System

Integration of the biomass CLG and SOFC systems presents overall benefits to the combined

system. Designing the CLG process such that no syngas cleaning or pre-treatment is required prior

to being sent to the SOFC permits the direct integration of the two systems, and removes the

necessity for intermittent syngas post-processing. CLG-SOFC integration offers in situ separation

of syngas diluents such as CO2, N2, and H2O, in situ removal and capture of SOFC performance

inhibitors such as H2S, and in situ tar and char reforming processes.

The regeneration of CaO sorbent via CaCO3 dissociation in CLG 1 occurs at the relatively

high temperature of 850 °C. Large quantities of heat are required to both achieve and maintain the

necessary temperature for maximum sorbent regeneration, thereby increasing energy costs. For

example, Figure 4.2 of Chapter IV shows the use of electric heaters for CaO regeneration in CLG

1. However, the requirement for external heating can be minimized with the use of an integrated

CLG-SOFC system by recycling the high-quality waste heat of the SOFC exhaust stream. Table

5.7 of Chapter V shows that the exhaust streams exit the SOFC stack at more than 830 °C.

Therefore, the use of this heat to drive sorbent regeneration would reduce heating costs and

improve the overall efficiency of the combined system.

As an added benefit, the CO2 present in the SOFC exhaust stream could combine with the

desorbed CO2 from the sorbent regeneration reaction to increase the flowrate of the CO2-rich

product stream used for sequestration or other processes requiring CO2. Furthermore, the H2O

present in the exhaust could be separated via condensation. The N2 present in the exhaust would,

however, dilute the CO2 product stream. The issue of CO2 dilution may be addressed through the

use of indirect heating via heat exchangers applied to the sorbent regeneration step of the process.

For example, the electric heaters shown in Figure 4.2 could be replaced with a heat exchanger unit

as a heat source for sorbent regeneration. Indirect heating would therefore prevent mixing of the

product CO2 stream with the N2 present in SOFC stack exhaust.

The benefit of recycling the SOFC exhaust stream heat may not exist for CLG 2 in the

same manner that it exists for CLG 1. The heat quality of the exhaust stream is too low to provide

sufficient heat to the combustion unit, which operates at 1250 °C. External heating would be

108

required to account for the > 400 °C temperature difference between the exhaust stream and the

reactor operating temperature required for the thermo-oxidative conversion of Fe3O4 to Fe2O3.

However, this requirement for external heating is still desirable in comparison to heating the

combustion unit from ambient conditions to the combustion temperature. Moreover, recycling the

exhaust heat to another high temperature reactor, the reducer, which operates at 870 °C, would

significantly reduce the required external heating to that reactor. However, both the CO2 and N2

present in the exhaust would ultimately dilute the syngas product stream since their respective

removals are not inherent to the CLG 2 process. This leads to reduced SOFC performance and

lower power generation. The extra H2O provided by the exhaust stream to the reducer would be

separated from the syngas via condensation in downstream reactors. The same issues arise when

considering recycle of the exhaust stream heat to the oxidizer, which operates at 720 °C and would

require no external heating when using the exhaust stream as a heat source.

Although the combined CLG-SOFC system using CLG 2 is not as well matched as with

CLG 1, the integrated system is still beneficial and more efficient when altering the system to

employ the use of heat exchangers. The exchangers would make use of the majority of the high

quality heat from the SOFC exhaust stream without diluting the syngas product stream with CO2

and N2. Thus, the use of this heat as a heat source to the reactors would decrease energy expenses

and increase the overall efficiency of the combined system.

Overall, the integration of the CLG 1 and SOFC systems provides the best match in terms

of SOFC performance and system efficiency improvements. These benefits arise in the form of

near-sufficient heating from the SOFC stack exhaust stream to the CaCO3 dissociation reaction.

Further, the removal of CO2 diluent from the syngas product stream is beneficial for SOFC

operation and performance, and for applications involving CO2 sequestration. Finally, CLG 1 is a

better match than CLG 2 since the exhaust heat is not solely sufficient for the high temperature

reactors of the latter and the requirement for external heating still exists.

109

Chapter VII

Conclusions and Recommendations

6.1 Overall Conclusions

Overall, it was found that:

H2 compositions in CLGs 1 and 2 syngas were 92.45 and 62.94 mol-%, respectively, when

using poultry litter (PL) as the biomass type and Aspen Plus as the simulation software.

Therefore, CLG 1 exhibited greater high-purity H2 syngas production capabilities.

Syngas yields for CLGs 1 and 2 were 0.87 and 2.54 kmol / kmol PL, respectively. Thus,

CLG 2 demonstrated greater absolute syngas production capabilities.

Although both CLG processes performed well in the simulated tubular SOFC, CLG 1

syngas demonstrated superior performance and is therefore the favourable process for

power generation by SOFC operation.

Modifying the CLG 1 process to exclude CO2 removal improved its performance as a result

of elevated amounts of CO present in the syngas. Very little effect on SOFC performance

was observed when modifying CLG 2 to include CO2 removal. In general, increasing

syngas CO2 concentration while maintaining fixed proportions of the other constituents

reduced performance.

Increasing syngas CO concentration greatly improved SOFC cell voltage and total power

output for both CLG types. Less of an increase was observed for electrical efficiency in

both cases.

The results of both of the biomass CLG processes and the tubular SOFC simulation

compared well to published literature in terms of absolute values and operating parameter

sensitivity trends.

The reduction of heating costs and the improvement of overall efficiency can be achieved

by recycling high quality heat from the SOFC exhaust in the integrated CLG-SOFC system.

110

6.2 Limitations of Research

The results of the conducted research were limited in-part by aspects of the utilized methods for

data acquisition. The limitations mainly stemmed from the Aspen Plus simulation software, the

employed calculation methods, and the assumptions used to simplify implementation of the former

two. The following lists the research limitations in greater detail.

Aspen Plus does not contain the required thermophysical or kinetic data regarding biomass

components to allow for the occurrence of spontaneous biomass steam gasification in an

equilibrium-based reactor. User-defined gasifier output is necessary when using

nonconventional biomass components. This limitation restricted sensitivity analyses from

variation of parameters such as SBR, SCR, Ca/C, and equivalence ratio.

Aspen Plus is unable to model electrochemical reactions as it does not recognize the input

of, and interaction between, ionic components. Consequently, the overall SOFC cell

reaction for operation on H2 was modelled as opposed to the anodic and cathodic half-cell

reactions. Considering the direct anodic oxidation of any reacted CO or CH4 also was not

possible as a result.

The mechanism by which H2S inhibits H2 oxidation at the anode and subsequently reduces

SOFC performance could not be modelled in Aspen Plus. This limitation restricted the

sensitivity analysis from variation of H2S levels in the syngas feed stream.

The assumption that all reacted CO and CH4 fed to the simulated SOFC was converted to

H2 implies that direct anodic oxidation did not occur for either of the fuels. Both are

equilibrium reactions and, although the assumption is valid as an approximation, will not

go entirely to completion during pre-reforming. Thus, some direct anodic oxidation of both

carbonaceous fuels is possible under the simulated conditions. This has further implications

regarding carbon deposition at the anode and consequent performance degradation, the

mechanism of which could not be modelled in Aspen Plus.

The reference voltage used to acquire SOFC cell voltage, electrical efficiency, and total

power output values was based on a semi-empirical reference curve developed for a fuel

stream consisting of 67% H2, 22% CO, and 11% H2O at standard operating conditions of

1000 °C, 1 bar, 85% fuel utilization factor, and 25% air utilization factor. Both CLGs 1 and

111

2 had compositions different from that of the reference fuel stream and therefore the

literature reference voltage was used in approximation and for the simplicity of the model.

The simulated results were compared to literature values, however no experimentation was

conducted to further validate the obtained data. Lab-scale experiments regarding the steam

gasification of poultry litter for H2 production with subsequent SOFC operation on the

resultant syngas would provide this data for comparison to the simulated results.

6.3 Recommendations for Future Research

This thesis presents research conducted on two very broad fields in engineering and science and it

was therefore not possible to explore all aspects of the fields. The following is a list of suggested

topics that may aid in further understanding the covered concepts.

Linking an external programming software to the Aspen Plus simulation flowsheet via

user-defined subroutines would allow for the increased efficacy of both the biomass CLG

and SOFC simulated processes. This would mitigate or eliminate some of the limitations

posed by the models. Biomass gasification could occur spontaneously in an equilibrium-

based unit module rather than in a module based on user-defined outputs. The inability to

detect anode pressure variation and the mechanism by which H2S inhibits anodic H2

oxidation would also be corrected. A user-defined subroutine to allow Aspen Plus to

integrate electrochemical reactions is an example of this.

Considering the presence of sulfur and nitrogen compounds in the poultry litter feedstock

and determining their respective effects on the H2 composition and production capabilities

of CLGs 1 and 2 would further increase the real-world accuracy of the processes.

A thermal analysis of the models would provide a better understanding of the energy

requirements for both systems. This analysis could be based both on the processes as a

whole and the individual Aspen Plus unit modules.

A cost analysis of the models would provide deeper insight into the feasibility and

scalability of the processes from an economic perspective. This analysis could be based on

system energy requirements and feed stream raw material supply expenses.

112

Appendices

Appendix Table of Contents

APPENDIX LIST OF TABLES ............................................................................................... 113

APPENDIX LIST OF FIGURES.............................................................................................. 114

Appendix A: Biomass Chemical Formula Calculations ...................................................... 115

A.1 Poultry Litter Calculations ................................................................................... 115

A.2 Willow Pellet Calculations .................................................................................. 116

A.3 Oak Pellet Calculations ........................................................................................ 117

Appendix B: Raw Data for Biomass CLG Simulation ........................................................ 118

B.1 CLG 1 Raw Data .................................................................................................. 118

B.2 CLG 2 Raw Data .................................................................................................. 121

Appendix C: Sensitivity Analyses for Biomass CLG Simulation ....................................... 123

C.1 CLG 1 Sensitivity Analysis Plots......................................................................... 123

C.2 CLG 2 Sensitivity Analysis Plots......................................................................... 125

Appendix D: Raw Data for SOFC Simulation ..................................................................... 127

D.1 CLG-1-NC Raw Data .......................................................................................... 127

D.2 CLG-1-WC Raw Data ......................................................................................... 128

D.3 CLG-2-WC Raw Data ......................................................................................... 129

D.4 CLG-2-NC Raw Data .......................................................................................... 130

113

Appendix List of Tables

Appendix A

Table A.1. Poultry litter chemical formula calculations including nitrogen and sulfur ........... 115

Table A.2. Poultry litter chemical formula calculations without nitrogen or sulfur present .... 115

Table A.3. Willow pellet chemical formula calculations including nitrogen and sulfur .......... 116

Table A.4. Willow pellet chemical formula calculations without nitrogen or sulfur present ... 116

Table A.5. Oak pellet chemical formula calculations including nitrogen and sulfur ............... 117

Table A.6. Oak pellet chemical formula calculations without nitrogen or sulfur present ........ 117

Appendix B

Table B.1.1. CLG 1 simulation raw data (part 1 of 3) .............................................................. 118





Appendix D

Table D.1. CLG-1-NC simulation results ................................................................................. 127

Table D.2. CLG-1-WC simulation results ................................................................................ 128

Table D.3. CLG-2-WC simulation results ................................................................................ 129

Table D.4. CLG-2-NC simulation results ................................................................................. 130

114

Appendix List of Figures

Appendix C

Fig. C.1. CLG 1 absorber temperature sensitivity analysis ...................................................... 122

Fig. C.2. CLG 1 absorber pressure sensitivity analysis ............................................................ 122

Fig. C.3. CLG 1 WGS reactor pressure sensitivity analysis ..................................................... 123

Fig. C.4. CLG 2 reducer pressure sensitivity analysis .............................................................. 124

Fig. C.5. CLG 2 oxidizer temperature sensitivity analysis ....................................................... 124

Fig. C.6. CLG 2 oxidizer steam flowrate sensitivity analysis ................................................... 125

Fig. C.7. CLG 2 reformer syngas composition temperature sensitivity analysis ..................... 125

115

Appendix A: Biomass Chemical Formula Calculations

A.1 Poultry Litter Calculations

Table A.1

Poultry litter chemical formula calculations including nitrogen and sulfur.

Element Wt.-% Mass (kg) Moles (kmol) Molecular Formula (based on C)

Carbon 43.3 0.433 5.20033 1

Hydrogen 6.62 0.0662 0.066862 0.012857261

Nitrogen 5.72 0.0572 0.801372 0.154100221

Sulfur 1.15 0.0115 0.368805 0.070919538

Oxygen 5.95 0.0595 0.952 0.183065305

Total 62.74 0.6274 7.389369

Ash 37.26

Table A.2

Poultry litter chemical formula calculations without nitrogen or sulfur present.

Element Mass (kg) Wt.-% Mass (kg/kg biomass) Moles (kmol) Molecular Formula (C-based)

Carbon 0.433 46.494148 0.46494148 5.583947171 1

Hydrogen 0.0662 7.1083432 0.071083432 0.071794266 0.0128573

Oxygen 0.0595 6.3889187 0.063889187 1.022226995 0.1830653

Ash 0.3726 40.00859 0.400085901 - -

Total 0.9313 100 1

116

A.2 Willow Pellet Calculations

Tab

Willow pellet chemical formula calculations including nitrogen and sulfur.


Carbon 50.65 0.5065 6.083065 1

Hydrogen 5.86 0.0586 0.059186 0.009729635

Nitrogen 0.52 0.0052 0.072852 0.011976199

Sulfur 0.44 0.0044 0.141108 0.023196859

Oxygen 24.07 0.2407 3.8512 0.633101898

Total 81.54 0.8154 10.207411

Ash 18.46

Table A.4

Willow pellet chemical formula calculations without nitrogen or sulfur present.


Carbon 0.5065 51.14095315 0.511409532 6.14202847 1

Hydrogen 0.0586 5.916801292 0.059168013 0.05975969 0.0097296

Oxygen 0.2407 24.30331179 0.243033118 3.88852989 0.6331019

Ash 0.1846 18.63893376 0.186389338 - -

Total 0.9904 100 1

117

A.3 Oak Pellet Calculations

Table A.5

Oak pellet chemical formula calculations including nitrogen and sulfur.


Carbon 52.23 0.5223 6.272823 1

Hydrogen 6.59 0.0659 0.066559 0.010610693

Nitrogen 0.62 0.0062 0.086862 0.013847354

Sulfur 0.29 0.0029 0.093003 0.014826339

Oxygen 33.98 0.3398 5.4368 0.866723005

Total 93.71 0.9371 11.956047

Ash 6.29

Table A.6

Oak pellet chemical formula calculations without nitrogen or sulfur present.


Carbon 0.5223 52.70965789 0.527096579 6.33042991 1

Hydrogen 0.0659 6.65051973 0.066505197 0.06717025 0.0106107

Oxygen 0.3398 34.29205773 0.342920577 5.48672924 0.866723

Ash 0.0629 6.347764658 0.063477647 - -

Total 0.9909 100 1

118

Appendix B: Raw Data for Biomass CLG Simulation

B.1 CLG 1 Raw Data

Table B.1.1

CLG 1 simulation raw data (part 1 of 3).

ABS-

IN

ABS-

OUT

BIOMA

SS CAO

CAO-

FEED

CAO-

OUT

COND-

IN

DES-

OUT

GAS-

OUT

Temperature C 750 500 25 650 25 25 750 650 25

Pressure bar 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013 1.013

Mass VFrac 0.927 0.015 0 0 0 0 0.999 0.823 0.018

Mass SFrac 0.073 0.985 1 1 1 1 0.001 0.177 0

*** ALL

PHASES ***

Mass Flow

kg/hr

32.96

7 369.432 14.952

336.12

8 336.464 336.128 2.848

1897.8

82

1561.75

5

Volume Flow

cum/hr

149.4

27 39.09 0.013 0.102 0.102 0.102 73.487

6518.1

59 41.048

Enthalpy

Gcal/hr

-

0.027 -0.95 -0.011 -0.863 -0.91 -0.909 0.003 -5.344 -5.831

Density

kg/cum 0.221 9.451

1133.79

6

3297.6

32

3297.63

2

3297.63

2 0.039 0.291 38.047

Mass Flow

kg/hr

H2 1.556 0.601 0 0 0 0 1.63 1.998 1.998

WATER 1.762 0.226 0 0 0 0 0.055 1513.4

42

1513.44

2

CO 21.29

7 0.015 0 0 0 0 0.314 0.168 0.168

CO2 5.564 0.005 0 0 0 0 0.002 45.811 45.811

CH4 0.378 4.86 0 0 0 0 0.844 0 0

O2 0 0 0 0 0 0 0 0 0

C 2.41 3.106 0 0 0 0 0.003 0 0

CAO 0 3.06E+

02 0

336.12

8 336.464 336.128 0

336.46

4 0.336

CACO3 0 54.933 0 0 0 0 0 0 0

BIOMASS 0 0 14.952 0 0 0 0 0 0

119

Table B.1.2


GAS

ES

GASE

S-2

GASIF-

IN

H2O-

FEED

H2O-

OUT

REF-

IN

SOL-

IDS

STE-

AM

SYN-

GAS

Temperature C 500 650 750 25 20 750 500 240 20

Pressure bar 1.013 1.013 1.013 1.013 10.133 1.013 1.013 1.013 10.133

Mass VFrac 0.94 1 0.546 0 0 1 0 1 1

Mass SFrac 0.06 0 0.454 0 0.145 0 1 0 0

*** ALL

PHASES ***

Mass Flow kg/hr 6.07 1561.7

55 32.967 18.015 0.022 32.967

363.36

1

1531.2

99 2.826

Volume Flow

cum/hr

38.97

5

6518.0

57 83.969 0.018 0

1.70E+

02

1.14E-

01

3579.1

01 2.103

Enthalpy

Gcal/hr

-

0.004 -4.48 -0.055 -0.068 0 -0.019 -0.946 -4.759 -0.001

Density kg/cum 0.156 0.24 0.393 993.957 1084.94

2 0.194

3175.7

68 0.428 1.344

Mass Flow kg/hr

H2 0.601 1.998 0 0 0 1.848 0 0 1.63

WATER 0.226 1513.4

42 18.015 18.015 0.019 0 0

1531.2

99 0.036

CO 0.015 0.168 0 0 0 31.119 0 0 0.314

CO2 0.005 45.811 0 0 0 0 0 0 0.002

CH4 4.86 0 0 0 0 0 0 0 0.844

O2 0 0 0 0 0 0 0 0 0

C 0.003 0 0 0 0.003 0 3.103 0 0

CAO 0.306 0.336 0 0 0 0 305.38 0 0

CACO3 0.055 0 0 0 0 0 54.878 0 0

BIOMASS 0 0 14.952 0 0 0 0 0 0

120

Table B.1.3


TAR WGS-IN WGS-OUT

Temperature C 750 750 750

Pressure bar 1.013 1.013 1.013

Mass VFrac 0 0.94 0.469

Mass SFrac 1 0.06 0.531

*** ALL PHASES ***

Mass Flow kg/hr 3.222 6.07 6.07

Volume Flow cum/hr 0.001 51.578 73.488

Enthalpy Gcal/hr 0 -0.003 0.003

Density kg/cum 2327.14 0.118 0.083

Mass Flow kg/hr

H2 0 0.601 1.63

WATER 0 0.226 0.055

CO 0 0.015 0.314

CO2 0 0.005 0.002

CH4 0 4.86 0.844

O2 0 0 0

C 2.886 0.003 2.889

CAO 0.336 0.306 0.336

CACO3 0 0.055 0

BIOMASS 0 0 0

121

B.2 CLG 2 Raw Data

Table B.2.1


Parameter

FE+

TAR

FE2

O3

FE3O4-

EX

FE3O4

-IN

GASE

S-1

GASE

S-2

H2+

CO

H2O+

GAS O2

OXID-

OUT

Temperature C 870 1250 720 25 870 720 750 20 25 720

Pressure bar 30.398

1.01

3 30.398 1.013 30.398 30.398 1.013 1.013

1.0

13 30.398

Mass VFrac 0

0.24

7 0 0 1 1 1 0 1 0.977

Mass SFrac 1

0.75

3 1 1 0 0 0 0 0 0.023

*** ALL

PHASES ***

Mass Flow

kg/hr 7.465

11.7

75 8.558 8.576 33.981

359.21

2 29.67

350.27

7 3.2 367.77

Volume Flow

cum/hr 0.001

11.3

8 0.002 0.002 4.897 54.616

153.0

82 0.35

2.4

47 54.617

Enthalpy

Gcal/hr 0.001

-

0.00

7 -0.009 -0.01 -0.045 -1.021

-

0.017 -1.32 0 -1.03

Density

kg/cum

5449.1

98

1.03

5 5200.31

5200.3

1 6.938 6.577 0.194

1001.7

53

1.3

08 6.734

Mass Flow

kg/hr

H2 0 0 0 0 0.848 0.721 1.66 0.002 0 0.721

CO 0 0 0 0 15.14 0.036 28.01 0.004 0 0.036

CO2 0 0 0 0 12.244 4.582 0 5.099 0 4.582

H2O 0 0 0 0 4.525

353.86

5 0

345.15

5 0 353.865

CH4 0 0 0 0 1.216 0 0 0 0 0

C 1.266 0 0 0 0.001 0 0 0 0 0

O2 0

2.90

4 0 0 0 0 0 0 3.2 0

FE 6.199 0 0 0 0.006 0 0 0 0 0

FE2O3 0

8.87

2 0 0 0 0 0 0 0 0

FE3O4 0 0 8.558 8.576 0 0.009 0 0.017 0 8.567

122

Table B.2.2


Parameter REDU-OUT REF-OUT STEAM SYNGAS

Temperature C 870 500 240 20

Pressure bar 30.398 1.013 32.424 1.013

Mass VFrac 0.82 1 1 1

Mass SFrac 0.18 0 0 0

*** ALL PHASES ***

Mass Flow kg/hr 41.446 393.193 360.306 42.916

Volume Flow cum/hr 4.899 1384.416 26.317 61.244

Enthalpy Gcal/hr -0.044 -1.114 -1.12 -0.085

Density kg/cum 8.46 0.284 13.691 0.701

Mass Flow kg/hr

H2 0.848 3.234 0 3.233

CO 15.14 0.447 0 0.443

CO2 12.244 43.275 0 38.176

H2O 4.525 346.207 360.306 1.053

CH4 1.216 0.012 0 0.012

C 1.267 0 0 0

O2 0 0 0 0

FE 6.205 0 0 0

FE2O3 0 0 0 0

FE3O4 0 0.017 0 0

123

Appendix C: Sensitivity Analyses for CLG Simulation

C.1 CLG 1 Sensitivity Analysis Plots

Fig. C.1. CLG 1 absorber temperature sensitivity analysis.

Fig. C.2. CLG 1 absorber pressure sensitivity analysis.

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0

0.2

0.4

0.6

0.8

1

1.2

300 400 500 600 700 800

Yie

ld:

km

ol/

km

ol

PL

Absorber Temperature (°C)

CO2 Production and Capture Efficiency vs. Absorber Temperature

H2 Exiting Absorber CO Exiting Abs CH4 Exiting Abs

CO2 Product Stream CO2 Capture Efficiency

0%

10%

20%

30%

40%

50%

60%

70%

80%

90%

100%

0

0.2

0.4

0.6

0.8

1

1.2

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

Cap

ture

Eff

icie

ncy

(%

)

Yie

ld:

km

ol/

km

ol

PL

Absorber Pressure (atm)

CO2 Production and Capture Efficiency vs. Absorber Pressure

H2 Exiting Absorber CO Exiting Abs CH4 Exiting Abs

CO2 Product Stream CO2 Capture Efficiency

124

Fig. C.3. CLG 1 WGS reactor pressure sensitivity analysis.

0

0.002

0.004

0.006

0.008

0.01

0.012

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20

Yie

ld:

km

ol/

km

ol

PL

(C

O,

CO

2,

H2O

)

Yie

ld:

km

ol/

km

ol

PL

(H

2an

d C

H4)

WGS Reactor Pressure (atm)

Syngas Yield vs. WGS Reactor Pressure

H2 CH4 CO CO2 H2O

125

C.2 CLG 2 Sensitivity Analysis Plots

Fig. C.4. CLG 2 reducer pressure sensitivity analysis.

Fig. C.5. CLG 2 oxidizer temperature sensitivity analysis.

0

0.1

0.2

0.3

0.4

0.5

0.6

30 31 32 33 34 35 36 37

Yie

ld:

km

ol/

km

ol

PL

Reducer Pressure (atm)

Reducer Exit Yield vs. Pressure at 870°C

H2 CO CO2 H2O CH4 Tar Fe Fe3O4

0

0.0005

0.001

0.0015

0.002

0.0025

0.003

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

600 620 640 660 680 700 720

Yie

ld:

km

ol/

km

ol

(CO

and

CH

4)

Yie

ld:

km

ol/

km

ol

PL

(H

2, C

O2,

Fe 3

O4)

Oxidizer Temperature (°C)

Oxidizer Exit Yield vs. Temperature at 30 atm

H2 CO2 Fe3O4 CO CH4

126

Fig. C.6. CLG 2 oxidizer steam flowrate sensitivity analysis.

Fig. C.7. CLG 2 reformer syngas composition temperature sensitivity analysis.

0

0.005

0.01

0.015

0.02

0.025

0.03

0.035

0.04

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

5 7 9 11 13 15 17 19 21 23 25

Yie

ld:

km

ol/

km

ol

(CO

, C

H4,

Fe 3

O4)

Yie

ld:

km

ol/

km

ol

(H2

and

CO

2)

Steam Flowrate (kmol/h)

Oxidizer Exit Yield vs. Steam Flowrate at 720°C and 30 atm

H2 CO2 CO CH4 Fe3O4

0.0%

0.5%

1.0%

1.5%

2.0%

2.5%

3.0%

3.5%

0%

10%

20%

30%

40%

50%

60%

70%

400 450 500 550 600 650 700 750 800 850

Co

mp

osi

tio

n (

CO

, H

2O

, C

H4)

Co

mp

osi

tio

n (

H2

and

CO

2)


Syngas Composition vs. Reformer Temperature

H2 CO2 CO H2O CH4

127

Appendix D: Raw Data for SOFC Simulation

D.1 CLG-1-NC Raw Data

Table D.1

CLG-1-NC simulation results.

AIR AIR-

DEPL

AN-

PROD

BURN-

IN

BURN-

OUT

CATH-

IN

EXH-

AUST

GAS-

WARM

HOT-

SIDE O2

SYN-

GAS

kmol/hr

H2 0 0 0.018608 0 0 0 0 0.806931 0 0 0.806931

H2O 0 0 0.895492 0 0.9141 0 0.9141 0.001998 0.9141 0 0.001998

O2 8.4 7.896 9.22E-13 7.896 7.885758 8.4 7.885758 0 7.885758 0.504 0

N2 31.6 31.6 0 31.6 31.6 31.6 31.6 0 31.6 0 0

C 0 0 2.91E-29 0 0 0 0 0 0 0 0

CO 0 0 0.001876 0 0 0 0 0.01121 0 0 0.01121

CO2 0 0 0.061966 0 0.063841 0 0.063841 4.54E-05 0.063841 0 4.54E-05

CH4 0 0 3.84E-12 0 3.83E-12 0 3.83E-12 0.052586 3.83E-12 0 0.052586

Total Flow

kmol/hr 40 39.496 0.977941 39.496 40.4637 40 40.4637 0.87277 40.4637 0.504 0.87277

Total Flow

kg/hr 1154.016 1137.888 18.94968 1137.888 1156.838 1154.016 1156.838 2.82229 1156.838 16.1274 2.82229

Total Flow

l/min 49405.95 59076.83 1635.882 63907.57 65473.38 59830.69 61077.65 1412.209 71137.25 753.8667 34.99041

Temperature

C 630 820.5662 950 910 910 820.5662 830.566 910 1012.35 820.5662 20

Pressure bar 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 10.1325

128

D.2 CLG-1-WC Raw Data

Table D.2

CLG-1-WC simulation results.

AIR AIR-

DEPL

AN-

PROD BURN-IN

BURN-

OUT CATH-IN

EXH-

AUST

GAS-

WARM

HOT-

SIDE O2

SYN-

GAS

kmol/hr

H2 0 0 0.010795 0 0 0 0 0.770297 0 0 0.770297

H2O 0 0 0.810309 0 0.821104 0 0.821104 0.003829 0.821104 0 0.003829

O2 8.4 7.602 3.97E-12 7.602 7.587955 8.4 7.587955 0 7.587955 0.798 0

N2 31.6 31.6 0 31.6 31.6 31.6 31.6 0 31.6 0 0

C 0 0 1.72E-28 0 0 0 0 0 0 0 0

CO 0 0 0.017294 0 0 0 0 0.759836 0 0 0.759836

CO2 0 0 0.891184 0 0.908478 0 0.908478 0.125153 0.908478 0 0.125153

CH4 0 0 2.44E-12 0 0 0 0 0.023489 0 0 0.023489

Total Flow

kmol/hr 40 39.202 1.729583 39.202 40.91754 40 40.91754 1.682604 40.91754 0.798 1.682604

Total Flow

kg/hr 1154.016 1128.481 54.32496 1128.481 1182.806 1154.016 1182.806 28.78992 1182.806 25.53504 28.78992

Total Flow

l/min 49405.95 58753.6 2893.213 63431.86 66207.73 59949.59 61884.33 2722.584 71935.12 1195.994 67.45691

Temperature

C 630 822.7397 950 910 910 822.7397 832.7397 910 1012.35 822.7397 20

Pressure bar 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 10.1325

129

D.3 CLG-2-WC Raw Data

Table D.3

CLG-2-WC simulation results.

AIR AIR-

DEPL

AN-

PROD BURN-IN

BURN-

OUT CATH-IN

EXH-

AUST

GAS-

WARM

HOT-

SIDE O2

SYN-

GAS

kmol/hr

H2 0 0 0.01376 0 0 0 0 1.600495 0 0 1.600495

H2O 0 0 1.646666 0 1.660426 0 1.660426 0.058435 1.660426 0 0.058435

O2 8.4 7.602 2.83E-12 7.602 7.590349 8.4 7.590349 0 7.590349 0.798 0

N2 31.6 31.6 0 31.6 31.6 31.6 31.6 0 31.6 0 0

C 0 0 8.55E-30 0 0 0 0 0 0 0 0

CO 0 0 0.009543 0 0 0 0 0.015816 0 0 0.015816

CO2 0 0 0.87446 0 0.884003 0 0.884003 0.867439 0.884003 0 0.867439

CH4 0 0 1.35E-12 0 0 0 0 0.000748 0 0 0.000748

Total Flow

kmol/hr 40 39.202 2.544429 39.202 41.73478 40 41.73478 2.542933 41.73478 0.798 2.542933

Total Flow

kg/hr 1154.016 1128.481 68.445 1128.481 1196.926 1154.016 1196.926 42.90996 1196.926 25.53504 42.90996

Total Flow

l/min 49405.95 58876.02 4117.082 63431.86 67530.09 60074.51 63250.44 4114.662 73371.87 1198.486 1019.459

Temperature

C 630 825.0232 910 910 910 825.0232 835.0191 910 1012.35 825.0232 20

Pressure bar 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325

130

D.4 CLG-2-NC Raw Data

Table D.4

CLG-2-NC simulation results.

AIR AIR-

DEPL

AN-

PROD BURN-IN

BURN-

OUT CATH-IN

EXH-

AUST

GAS-

WARM

HOT-

SIDE O2

SYN-

GAS

kmol/hr

H2 0 0 0.023041 0 0 0 0 1.618812 0 0 1.618812

H2O 0 0 1.634285 0 1.657326 0 1.657326 0.038402 1.657326 0 0.038402

O2 8.4 7.602 3.40E-12 7.602 7.590446 8.4 7.590446 0 7.590446 0.798 0

N2 31.6 31.6 0 31.6 31.6 31.6 31.6 0 31.6 0 0

C 0 0 7.05E-31 0 0 0 0 0 0 0 0

CO 0 0 6.70E-05 0 0 0 0 7.14E-05 0 0 7.14E-05

CO2 0 0 0.003264 0 0.003331 0 0.003331 0.003204 0.003331 0 0.003204

CH4 0 0 4.95E-14 0 0 0 0 5.61E-05 0 0 5.61E-05

Total Flow

kmol/hr 40 39.202 1.660657 39.202 40.8511 40 40.8511 1.660545 40.8511 0.798 1.660545

Total Flow

kg/hr 1154.016 1128.481 29.63409 1128.481 1158.115 1154.016 1158.115 4.099048 1158.115 25.53504 4.099048

Total Flow

l/min 49405.95 58702.99 2777.916 63431.86 66100.23 59897.96 61731.33 2686.89 71818.32 1194.964 665.7045

Temperature

C 630 821.7958 950 910 910 821.7958 831.7995 910 1012.35 821.7958 20

Pressure bar 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325 1.01325

Documents

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