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1995 Simultaneous Water-Gas Shift Reaction and Carbon Dioxide Separation for Direct Hydrogen Production From Synthesis Gas. Chun Han Louisiana State University and Agricultural & Mechanical College

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A Bell & Howell Information Company 300 North Zeeb Road, Ann Arbor, Ml 48106-1346 USA 313/761-4700 800/521-0600

SIMULTANEOUS WATER-GAS SHIFT REACTION AND CARBON DIOXIDE SEPARATION FOR DIRECT HYDROGEN PRODUCTION FROM SYNTHESIS GAS

A Dissertation

Submitted to the Graduate Faculty of the Louisiana State University and Agricultural and Mechanical College in partial fulfillment of the requirements for the degree of Doctor of Philosophy

in

The Department of Chemical Engineering

by Chun Han B.S., University of Science and Technology of China, 1988 M.S. in Ch.E., Louisiana State University, 1992 May 1995 UMI Number: 9538734

UMI Microform 9538734 Copyright 1995, by UMI Company. All rights reserved.

This microform edition is protected against unauthorized copying under Title 17, United States Code.

UMI 300 North Zeeb Road Ann Arbor, MI 48103 To my parents ACKNOWLEDGEMENTS

I would like to express my sincere gratitude and deep respect for my major advisor, Dr. Douglas P. Harrison, who initiated me into this research and guided me with patience, encouragement, and understanding throughout this study.

Thanks are due to Dr. Ralph W. Pike, Dr. Geoffrey L. Price,

Dr. Gregory L. Griffin, Dr. David M. Wetzel, and Dr. Jeffrey

S. Hanor for serving as members of examining committee.

The support received from the US Department of Energy

(Contract - DE-AC21-89MC26366) is gratefully acknowledged.

Further thanks go to the Department of Chemical Enineering for providing financial support during my first semester.

I also would like to thank my fellow students, Arpaden

Silaban, Marcel Narcida, Alejandro Lopez, and Julie White who provided assistance and shared with me during this research.

The help from student workers Matt H. Schumacher, Michael

Golda, and Keith Cagnolatti are greatly appreciated.

Special thanks are given to our host family, Joyce and

George Brown. Our friendship is treasured and will be remembered for life.

I must acknowledge the debt I own to my parents and the rest of my family for their support and unconditional love.

Most of all, I thank my wife for her infinite patience, endurance and encouragement, without which this dissertation would never have been completed. TABLE OF CONTENTS

DEDICATION ...... ii

ACKNOWLEDGEMENTS ...... iii

LIST OF TABLES ...... vii

LIST OF FIGURES ...... ix

ABSTRACT ...... xvi

CHAPTER 1 INTRODUCTION ...... 1 1.1 Hydrogen and Related Technologies ... 1 1.2 Hydrogen Production Via Coal Gasification ...... 3 1.2.1 Hot Gas Cleanup ...... 4 1.2.2 C02 Removal ...... 5 1.3 Conventional Shift Process and Current Research ...... 6 1.3.1 Theoretical Basis ...... 8 1.3.2 Research Objectives ...... 11

CHAPTER 2 LITERATURE REVIEW ...... 15 2.1 Water Gas Shift (WGS) Reaction ..... 15 2.1.1 Homogeneous WGS Reaction 16 2.1.2 Catalytic WGS Reaction ...... 16 2.2 Shift Processes ...... 20 2.2.1 Conventional Shift Processes ...... 20 2.2.2 Novel Shift Concepts ...... 24 2.3 Fundamental CaO Carbonation Studies ...... 26 2.3.1 Calcination/Carbonation of CaO based Sorbents ...... 26 2.3.2 Structural Changes during Calcination/Carbonation Reactions ...... 31

CHAPTER 3 EXPERIMENTAL APPARATUS AND PROCEDURE .... 41 3.1 Fixed-bed Reactor System ...... 41 3.2 Gas Chromatography System ...... 49 3.3 Sorbent Precursor Description ...... 57 3.4 Sorbent Structural Property Measurement ...... 60 3.5 Experimental Procedure in Running the Fixed-Bed Reactor System ...... 62

CHAPTER 4 EXPERIMENTAL RESULTS AND DISCUSSION: PRELIMINARY TESTS ...... 66 iv 4.1 Reactor Temperature Response ...... 67 4-2 GC Analytical Method and Calibration...... 70 4.3 Reactor Dead Volume Evaluation — Tracer - Response Tests ...... 79 4.4 Preliminary Reaction Tests ...... 82 4.4.1 Calcination and Carbonation Reactions ...... 84 4.4.2 Simultaneous Shift and Carbonation Reactions ...... 93

CHAPTER 5 EXPERIMENTAL RESULTS AND DISCUSSION: SINGLE CYCLE STUDIES ...... 97 5.1 Reaction Parameters ...... 97 5.1.1 Silaban's Results ...... 98 5.1.2 Characteristics of the Fixed-bed Reactor ...... 100 5.1.3 The Water-Gas Shift Reaction...... 101 5.1.4 Preliminary Reaction Tests and Final Parameter Selection .... 101 5.2 Shift/Carbonation Reaction Variables ...... 104 5.3 Experimental Result Reproducibility ...... 110 5.4 Detailed Parameter Tests ...... 115 5.4.1 Comparison of Different Sorbent Precursors ...... 118 5.4.2 Effect of Calcination Temperature ...... 121 5.4.3 Effect of Calcination Flow R a t e ...... 128 5.4.4 Effect of H20 in the Calcination Sweep Gas ...... 132 5.4.5 Effect of Carbonation/Shift Temperature ...... 141 5.4.6 Effect of Carbonation/Shift Pressure ...... 146 5.4.7 Effect of Carbonation/Shift Space Velocity ...... 151 5.4.8 Effect of Feed Gas Composition ...... 157 5.4.8.1 Effect of H20/C0 Ratio ...... 157 5.4.9 Effect of Sorbent Particle Size ...... 166 5.4.10 The Nature of the Shift Reaction...... 169 5.4.11 Sorbent structural Change Along the Reactor A x i s ...... 172 5.5 Conclusions ...... 175

v 5.5 Conclusions ...... 175

CHAPTER 6 EXPERIMENTAL RESULTS AND DISCUSSION: MULTICYCLE TESTS ...... 177 6.1 Comparison of Sorbent Performance on Five-Cycle Tests ...... 177 6.2 Effect of Carbonation/Shift Temperature ...... 186 6.3 Effect of Carbonation/Shift Space Velocity ...... 196 6.4 Effect of Carbonation/Shift Gas Composition ...... 201 6.5 Effect of Calcination Conditions .... 207 6.6 Eleven-Cycle Test Results ...... 215 6.7 Ten Cycle Results Under Isobaric Conditions ...... 217 6.8 Conclusions ...... 222

CHAPTER 7 CONCLUSIONS AND RRCOMMENDATIONS ...... 226

REFERENCES ...... 233

VITA 237

vi LIST OF TABLES

Table 2-1: Surface Area and Pore Volume of CaO ... 34

Table 2-2: Structural Properties of Test Sorbents ...... 39

Table 3-1: Chemical Analysis of Dolomite (as reported by National Lime Co., Findlay, Ohio) ...... 59

Table 4-1: GC Operating Conditions ...... 75

Table 4-2: Calibration Gas Compositions (mol percent) ...... ,...... 76

Table 4-3: Standard Gas Composition Used for GC Calibration ...... 77

Table 4-4: Quadratic Constants for GC Calibration ...... 78

Table 4-5: Summary of Reaction Conditions for the Preliminary Tests ...... 85

Table 5-1: Summary of The Results from TGA Studies (Silaban, 1993) 99

Table 5-2: Reaction Parameters ...... 103

Table 5-3: Feed Gas Composition for Test 18 Using Two Calculation Methods ...... 109

Table 5-4: Comparison of Duplicate Test Results under Standard Conditions ...... 117

Table 5-5: Summary of Reaction Test Conditions for the Different Sorbent Precursors ...... 119

Table 5-6: Summary of Reaction Test Conditions for the Effect of Calcination Temperature ...... 123

Table 5-7: Summary of Reaction Test Conditions for the Effect of Calcination Gas Flow Rate ...... 130

vii Table 5-8: Summary of Reaction Test Conditions for the Effect of Calcination Gas Composition...... 135

Table 5-9: Summary of Reaction Test Conditions for the Effect of Carbonation Temperature...... 142

Table 5-10: summary of Reaction Test Conditions for the Effect of Carbonation Pressure ...... 150

Table 5-11: Summary of Reaction Test Conditions for the Effect of Carbonation Gas Flow Rate ...... 152

Table 5-12: Summary of Reaction Test Conditions for the Effect of Carbonation Gas Composition ...... 160

Table 5-13: Summary of Reaction Test Conditions • for the Effect of Sorbent Particle Size ...... 167

Table 6-1: Summary of Reaction Conditions for Multicycle Tests ...... 178

Table 6-2: Prebreakthrough Concentrations of CO and C02 in Tests 66 Through 69 .... 211

viii LIST OF FIGURES

Figure 1-1: Equilibrium CO Conversion Versus Temperature ...... 10

Figure 1-2: Proposed Fluidized-Bed Process for Hydrogen Production ...... 12

Figure 1-3: Advanced Gasification - Carbonate Fuel Cell System (from Steinfeld et al., 1991) ...... 13

Figure 2-1: The Shift Process Using Two High - Temperature Reactor With Intermediate C02 Removal (from Kohl and Riesenfeld, 1977) 22

Figure 2-2: The Shift Process Using A High - Temperature Reactor Following by A Low - Temperature Reactor (from Rase, 1977) 23

Figure 2-3: CaC03 Calcination and Recarbonation ... 28

Figure 2-4: Comparison of Carbonation Results Using Simulated Coal Gas and C02/N2 Only (from Silaban, 1993) 32

Figure 2-5: Pore Size Distribution of Reagent Calcium Carbonate at 750°C for l hour (from Narcida, 1992) 36

Figure 2-6: Pore Size Distribution of Reagent Calcium Acetate at Various Decomposition Stages (from Narcida, 1992) ...... 37

Figure 2-7: Pore Size Distribution of Dolomite Calcined at 750°C for 1 hour (from Narcida, 1992) 38

Figure 3-1: Schematic Diagram of Fixed-Bed Reactor S y s t e m ...... 42

Figure 3-2: Refined Schematic of the Fixed-Bed Reactor ...... 44

Figure 3-3: Diagram of the Original Fixed-Bed Reactor ...... 47

ix Figure 3-4: Schematic Diagram of GC Sampling Sequence ...... 50

Figure 3-5: Chromatogram for Two Gas Mixtures ..... 53

Figure 3-6: The Typical Reactor Response for a Complete Cycle Test ...... 55

Figure 3-7: TCD and FID Results of Product Gas Concentrations During Shift/Carbonation Reaction Stage ...... 56

Figure 3-8: Reactor Temperature Response During Shift/Carbonation Stage ...... 58

Figure 3-9: Pore Size Distribution from Porosimeter for Calcined Dolomite ...... 61

Figure 4-1: Reactor Temperature - Time Response in Nonreacting Experimental Test ...... 68

Figure 4-2: Temperature Deviation from Set Point in Nonreacting Experimental Test ...... 69

Figure 4-3: GC Operating System (Stage 1) 72

Figure 4-4: GC Operating System (Stage 2) 73

Figure 4-5: GC Calibration Curves ...... 80

Figure 4-6: Hydrogen Response Curve ...... 81

Figure 4-7: Reactor Time Delay as a Function of Volumetric Flow Rate ...... 83

Figure 4-8: Effect of Calcination Temperature ..... 86

Figure 4-9: Effect of Carbonation Temperature ..... 87

Figure 4-10: Prebreakthrough Partial Pressure of C02 from the Fixed-Bed Reactor ...... 89

Figure 4-11: Calcination of Different Sorbents ..... 90

Figure 4-12: Carbonation of Different Sorbents ..... 91

Figure 4-13: Product Gas Concentration During Carbonation ...... 94

Figure 5-1: Dimensionless Breakthrough and Calcium Conversion Response Curves for a Fixed-Bed Reactor ...... 107

x Figure 5-2: Breakthrough Curves for Duplicate Tests ...... Ill

Figure 5-3: Fractional C0X Removal and Sorbent Conversion for Duplicate Tests ...... 112

Figure 5-4: Temperature Deviation Within the Sorbent Bed for Test 38 ...... 113

Figure 5-5: Temperature Deviation Within the Sorbent Bed for Test 44 ...... 114

Figure 5-6: Log-plot of Breakthrough Curves for Duplicate Tests ...... 116

Figure 5-7: Fractional Removal of Carbon Oxides for Three Test Sorbents at Standard Conditions ...... 120

Figure 5-8: Comparison of the First Cycle Reactivity of Dolomite and Marble Chips on a Dimensionless Time Basis ...... 122

Figure 5-9: C02 Composition of the Calcination Product Gas as a Function of Time and Temperature ...... 125

Figure 5-10: Dimensionless Carbonation Breakthrough Curves Following Different Calcination Temperatures ...... 126

Figure 5-11: Effect of Temperature on C02 Concentration During Calcination ...... 127

Figure 5-12: Effect of Calcination Temperature on Breakthrough Curves ...... 129

Figure 5-13: Calcination Product Composition as a Function of N2 Flow Rate ...... 131

Figure 5-14: Effect of Calcination Flow Rate on Breakthrough Curves ...... 133

Figure 5-15: C02 Content of Calcination Product Gas as a Function of Temperature and Steam Content ...... 136

Figure 5-16: Pore Size Distribution of Calcined Dolomite as a Function of Calcination Atmosphere ...... 137 Figure 5-17: Dimensionless C0X Breakthrough Curves Showing the Effect of Steam During Calcination ...... 139

Figure 5-18: Dimensionless Response Curves Following Calcination at High Temp, and High Steam Content ...... 140

Figure 5-19: Prebreakthrough CO and C02 Concentrations as a Function of Carbonation/Shift Reaction Temperature ...... 143

Figure 5-20: Postbreakthrough Concentrations of CO and C02 as a Function of Shift/ Carbonation Temperature ...... 145

Figure 5-21: Change in the Shape of the Carbon Oxide Breakthrough Curves With Shift/Carbonation Temperature ...... 147

Figure 5-22: Log-plot of Breakthrough Curves Showing the Effect of Carbonation Pressure .... 148

Figure 5-23: Prebreakthrough Concentrations of CO and C02 as a Function of Space Velocity ...... 153

Figure 5-24: Postbreakthrough Concentrations of CO and C02 as a Function of Space Velocity ...... 155

Figure 5-25: Change in the Shape of the Carbon Oxide Breakthrough Curves With Space Velocity ...... 156

Figure 5-26: Reactor Response. Product Composition for Test 24 ...... 158

Figure 5-27: Reactor Response. Product Composition for Test 27-1...... 159

Figure 5-28: Postbreakthrough Concentrations of CO and C02 as a Function of H20 to CO Ratio in Feed Gas ...... 162

Figure 5-29: Reactor Response. Product Composition for Test 53 ...... 164

Figure 5-30: Postbreakthrough Concentrations of CO and C02 as a Function of H20 to CO Ratio in Feed Gas ...... 165

xii Figure 5-31: Comparison of Dimensionless Breakthrough Curves as a Function of Sorbent Particle Size ...... 168

Figure 5-32: Fractional Extent of the Shift Reaction Using Various Reactor Packing ...... 171

Figure 5-33: Pore Volume and Pore Size Distribution Curves From Axial Sections for Sorbent Calcined in N2 (test 51) and H20/N2 (test 33) 174

Figure 6-1: Comparison of CO and C02 Breakthrough Curves in the First and Fifth Cycles of Multiple-Cycle Test ...... 180

Figure 6-2: CO and C02 Breakthrough Times as a Function of Cycle Number ...... 182

Figure 6-3: Fractional Calcium Conversion at Selected Dimensionless Times as a Function of Cycle Number ...... 183

Figure 6-4: Comparison of CO and C02 Breakthrough Curves in the First and Fifth Cycles of Test 63 ...... 185

Figure 6-5: Comparison of the Multicycle Deterioration of Dolomite (Test 27) and Marble Chips (Test 63) 187

Figure 6-6: Prebreakthrough CO Concentration as a Function of Temperature and Cycle Number ...... 188

Figure 6-7: Prebreakthrough C02 Concentration as a Function of Temperature and Cycle Number ...... 189

Figure 6-8: Fractional Removal of Total Carbon Oxides as a Function of Temperature and Cycle Number ...... 191

Figure 6-9: Fractional Removal of Total Carbon Oxides During the First and Fifth Cycles of Test 56 ...... 192

Figure 6-10: CO and C02 Breakthrough Curves from the First and Fifth Cycles of Test 60 .... 193

Figure 6-11: Sorbent Deterioration in Multicycle Tests as a Function of Temperature .... 195 xiii Figure 6-12: Pore Volume and Pore Size Distribution of Axial Sorbent Samples Following Fifth Cycle of Test 56 ...... 197

Figure 6-13: CO Breakthrough Curves for the Five Cycles of Test 62 ...... 199

Figure 6-14: Comparison of the Multicycle Deterioration of Dolomite as a Function of Space Velocity ...... 200

Figure 6-15: Prebreakthrough CO and C02 Concentrations and Fractional Removal • of Total Carbon Oxides as a Function of Cycle Number: Test 64 ...... 203

Figure 6-16: Fractional COx Removal and Fractional Sorbent Calcium Conversion as a Function of Dimensionless Time in First and Fifth Cycles of Test 64 ... 204

Figure 6-17: Prebreakthrough CO and C02 Concentrations and Fractional Removal of Total Carbon Oxides as a Function of Cycle Number: Test 66 ...... 206

Figure 6-18: Fractional COx Removal and Fractional Sorbent Calcium Conversion as a Function of Dimensionless Time in First and Fifth Cycles of Test 66 ... 208

Figure 6-19: CO and C02 Breakthrough Curves from the First and Fifth Cycles of Test 67 .... 209

Figure 6-20: Comparison of the Decrease in Sorbent Reactivity for CO as a Function of Cycle Number for Tests 66 through 69 ...... 212

Figure 6-21: Comparison of the Decrease in Sorbent Reactivity for C02 as a Function of Cycle Number for Tests 66 through 69 ...... 213

Figure 6-22: Prebreakthrough CO and C02 Concentrations and Fractional Removal of Total Carbon Oxides as a Function of Cycle Number: Test 65 ...... 216

xiv Figure 6-23j Dimensionless Time Required for the CO Concentration to Reach 100 ppm and for the C02 Concentration to reach 500 ppm as a Function of Cycle Number: Test 65...... 218

Figure 6-24: CO, C02, and H2 Breakthrough Curves for the First, Sixth, and Tenth Cycles of Test 72 ...... 220

Figure 6-25: Fractional COx Removal and Fractional Sorbent Calcium Conversion for the First, Sixth, and Tenth Cycles of Test 72 ...... 221

Figure 6-26: Sorbent Performance Deterioration Through the Ten Cycles of Test 72 .... 223

Figure 6-27: Comparison of Sorbent Deterioration Rates of Tests 65 and 72 ...... 224

xv ABSTRACT

The simultaneous water-gas shift and CaO-carbonation reactions for the direct hydrogen production from synthesis gas have been studied at high temperature and high pressure using a laboratory-scale fixed-bed reactor. Experiments were conducted to evaluate the effect of process parameters - temperature, pressure, gas composition, and space velocity - on system performance during both calcination and shift/carbonation reaction stages. Commercial dolomite was chosen as the primary sorbent while two limestones were tested for comparison. Multicycle tests were conducted to evaluate sorbent durability. Sorbent structure and structural changes associated with reaction were measured using mercury porosimeter.

The experiments proved the technical feasibility of combining the shift and carbonation reactions for hydrogen production. Calcination can be conducted over a temperature range of 750°C to 900°C under either N2 or a mixture of N, and steam. Equilibrium for the combined shift and carbonation reactions can be closely approached at 15 atm, in the temperature range of 500°C to 650°C, and at space velocity as high as 3400 hr'1 (STP) using simulated coal gas feed having a H20 to CO ratio as low as 2:1. Fractional carbon oxide removal at these conditions exceeded 0.99, with a maximum of

0.999 at 500°C and 15 atm. The corresponding total carbon

xvi oxide concentration during the prebreakthrough period was below 300 ppm.

The CO and C02 concentrations during the prebreakthrough period did increase in multicycle tests. However, sorbent reactivity loss was evident by a decrease in the duration of the prebreakthrough period. The deterioration rate for dolomite was about 3% to 5% per cycle at favorable reaction conditions. When shift/carbonation pressure decreased from 15 to 1 atm, the fractional carbon oxide removal during the early stage of the reaction decreased to about 0.95, but the sorbent durability did not suffer.

xvii CHAPTER 1

INTRODUCTION

Hydrogen is one of the most important materials in many areas of commerce and science. It has been used extensively in chemical manufacturing, petroleum refining and as a direct fuel. As a chemical feedstock, hydrogen is largely consumed in ammonia production and methanol synthesis. Heavy petroleum fractions are upgraded by a number of processes such as hydrorefining and hydrocracking. As an alternative fuel, hydrogen, taking the form of an odorless, nontoxic, colorless gas, has been successfully demonstrated as a vehicular fuel and a possible household resource replacing natural gas with lower pollution levels and higher efficiency (Billings, 1983) .

In advanced electric power generation systems such as fuel cells, hydrogen is a key to maintaining catalyst integrity and achieving efficient operation.

1.1. Hydrogen and Related Technologies

Many methods of hydrogen production are available. The current technologies can be divided into four categories which are electrochemical, thermochemical, photolytic and chemical processes (Williamson et al., 1986). Electrochemical processes produce hydrogen through the electrolysis of water.

Although this has been a commercial process, it is generally considered inefficient and expensive. Thermochemical hydrogen

l production uses high temperature nuclear or solar heat to

decompose water, while photolytic approaches involve the

direct splitting of water using a catalyst and solar

radiation. The latter two methods are still in the early

stages of research. Hence chemical processes based on

conventional fossil fuels are and will continue to be a primary method of producing hydrogen.

Fossil fuels such as natural gas, oil, or coal can be

converted to hydrogen using one of two techniques: steam

reforming or partial oxidation. In the steam reforming

reaction, light hydrocarbons, including natural gas, are

reacted with steam to produce hydrogen, partly derived from

the hydrocarbon itself and partly from the steam. In the

second method, heavier feedstocks, such as coal or residual oil, are partially oxidized and reacted with steam (or a mixture of steam and oxygen) to produce hydrogen.

The direct output from a reformer or a gasifier consists mainly of a mixture of hydrogen, steam, carbon monoxide and carbon dioxide, called synthesis gas. Although carbon monoxide is sometimes the only desired product (Vannby et al.,

1992), in most plants, synthesis gas is either completely converted to hydrogen for ammonia production or adjusted to certain H2/CO ratios for methanol synthesis. 1.2. Hydrogen Production Via Coal Gasification

Because they are inexpensive and readily available,

petroleum and natural gas are still the primary sources for

hydrogen. However, coal is an order of magnitude more

abundant than fluid hydrocarbons. As world reserves of

natural gas and oil are depleted, it will become necessary to

turn to coal as a major source of hydrogen.

Realizing that coal utilization should be both

economically attractive and environmentally acceptable, the

U.S. Department of Energy sponsors the development of various

coal conversion technologies. Coal gasification promises to be one of the most important technologies. Coal-derived gas

has numerous potential diverse applications, such as electric power generation using an integrated gasification combined-

cycle (IGCC) and/or coal gasification molten carbonate fuel

cells (CGMCFC), and the production of synthesis gas, hydrogen

and other chemicals. While improved energy efficiency can be

achieved with coal gasification, reducing or removing the

adverse environmental impact associated with coal utilization has also been a major concern.

Coal gases from the gasifier usually contain particulates

and trace gas impurities such as H2S, COS and N0X, and have a

typical temperature range of 650°C - 1250°C. In order for the processes mentioned above to function properly, these contaminants must be removed. Conventional removal of these

trace contaminants is accomplished by wet scrubbing operations which require that the hot coal gas be cooled to near ambient temperature before treatment. Considering the fact that the downstream processes in IGCC, CGMCFC or NH3 production operate at high temperature, the low-temperature purification causes a loss in thermal efficiency. The cooling and heating steps complicate the process and add capital cost for heat exchangers. Hence hot gas clean-up and acid gas removal are favored in order to increase the overall efficiency.

1.2.1. Hot Gas Cleanup

Hot cyclones and ceramic filters have been used to remove particulates at a temperature close to that of coal gasifier outlet gases. Various approaches for removing H2S at high temperature have been investigated during the past three decades. The most promising is to react the H2S with an appropriate metal oxide to form the corresponding metal sulfide.

Coal gases, after the removal of particulates and trace contaminants, consist of H2, H20, CO, C02, a small amount of

CH4, and N2. Additional hydrogen can be produced by reacting the CO with H20 according to the water gas shift reaction

CO(gr) + HzO(g) + C02(g) + Hz(g) (l-l)

Carbon dioxide must be removed in order to produce high purity of hydrogen. According to the above reasoning, high temperature removal of C02 is favored for improving the overall efficiency. The U.S. Department of Energy, Morgantown Energy Technology Center has identified the need for bulk separation processes which would operate with a temperature range of 100°C - 700°C.

1.2.2. C02 Removal

Based on this objective, several research groups are investigating different approaches. Two major methods being developed are high temperature membrane separation and a calcium oxide based sorbent process.

Membrane Process

Instead of trying to remove C02 from the coal gas, a membrane process aims at separating pure H2. Because of the permeability difference between H2 and other components in the coal gas, only hydrogen passes the membrane. Gavalas (1992) studied high temperature separation of hydrogen by ceramic membranes over the temperature range of 500°C - 750°C.

Membrane preparation was carried out by chemical vapor deposition of Si02 on porous Vycor support tubes with various pore sizes. At 500°C, the H2 selectivity of the membrane is more than four orders of magnitude greater than 02 and N,.

While the efficiency is very promising, the membrane stability under high temperature is still a major concern. Similar studies on gas separation using membranes were conducted by

Edlund (1992), Liu et al. (1992) and Uemiya et al. (1991). 6

CO. Sorbent Process

Silaban (1993) studied the high temperature removal of

C02 using CaO-based sorbents. A TGA reactor was used to study the kinetics of calcination and carbonation of several different sorbents, including dolomite, calcium carbonate and calcium acetate over the temperature range of 500°C - 800°C.

Multicycle runs were made to test the durability of the sorbents. They concluded that CaO-based sorbent (especially dolomite) was capable of removing C02 at conditions close to those of the coal gas.

1.3. Conventional Shift Process and Current Research

The investigation of current hydrogen production processes, coal gasification technology, and favorable results from previous research inspired the concept of a new process which will possibly improve hydrogen production efficiency.

In the conventional hydrogen production process, synthesis gas from a reformer or coal gasifier undergoes shift reaction to convert more CO to H2. A detailed review of the conventional shift process will be presented in Chapter 2

(Literature Review), while only one example is discussed at this point. A typical industrial process involves multiple catalyzed reaction steps with intermediate C02 removal using low temperature scrubbing, and associated gas-to-gas heat exchange between the shift reaction and C02 removal steps

(Kohl and Riesenfeld, 1979). If high purity hydrogen is 7 needed, methanation or pressure swing adsorption (PSA) is usually employed following the shift reactors (Ahn and Fischer

1985). Several disadvantages exist in this process. First, the process is complicated by multiple reactors, scrubbers, and solvent regenerators, which represent a large capital cost. The relatively high temperature of the catalytic reaction and the low temperature required for wet scrubbing of

C02 require many heat exchange steps. In addition, the large amount of excess steam needed to drive the reversible shift reaction contributes a significant fraction to the hydrogen production cost. When a low-temperature shift reactor is used, more stringent sulfur removal is required because the

Cu-Zn catalyst is extremely sensitive to sulfur poisoning.

Finally, although PSA can achieve high purity of hydrogen

(99.9% or higher, Balthasar et al., 1980), the "Polybed PSA unit" working at ambient temperature will create an additional energy penalty when hot hydrogen is needed, as for example, in

NH3 production.

Based on the above analysis and the favorable results from the previous research on high-temperature high-pressure

C02 removal at LSU, a one-step process for the direct production of hydrogen from coal gas based upon combining the water gas shift (WGS) reaction with C02 removal has been studied. The general concept was first proposed by Gluud et al. (1931) and later revived by Squires (1967). However, the idea was not pursued further because of the availability of reliable and low-cost methods of C02 removal near ambient temperature at that time. It is believed that the one-step approach is technically applicable to current commercial processes and could improve the efficiency of hydrogen production.

1.3.1. Theoretical Basis

The proposed one-step process for the production of hydrogen can be represented by the following two reactions:

CO + HzO C02 + H2 (1-2)

CaO + C02 ** CaC03 (1-3)

The water gas shift reaction (1-2) is reversible and exothermic. Although a heterogenous catalyst is usually used to enhance the reaction rate, previous studies (see Chapter 2) show that WGS can occur homogeneously at sufficiently high temperature. In their recent research on the bulk separation of C02 using a calcium oxide sorbent, Silaban (1993) found indirect evidence that the above two reactions take place simultaneously at temperatures around 750°C.

The potential of the one-step process for hydrogen manufacture may be illustrated by the following thermodynamic analysis. The equilibrium constant for water-gas shift reaction can be expressed as By using Barin & Knacke's thermodynamic data (1973), we can generate Kp2 as a function of temperature:

log10 = a + bT + cT2 + dr3 + gT4 • (1-5) where a = 10.56, b = -2.90E-2, c = 3.06E-5, d = -1.41E-8, and e = 2.07E-12. Temperature is in Kelvin.

Consider a typical coal gas consisting of 22%(mol) H2,

30% CO, 8% C02 and 40% H20 (Han and Harrison, 1994) . The fractional CO conversion at equilibrium as a function of temperature is shown as curve A in Figure 1-1. At the typical

720.K outlet temperature from a high-temperature shift reactor, about 70% of CO may be converted. If a low-temperature shift reactor at 550K is employed after intermediate cooling, only about 90% CO conversion is achievable. Now, suppose the CaO carbonation reaction also takes place. The equilibrium constant for reaction (1-3) is

KPz = ~ ~ (1-6 ) *0 0 , Kp2 can be expressed as:

log10Kp2 - A + BT + CT2 + DT3 + ET* (1-7) where A = 47.69, B = -0.13, C = 0.14E-3, D = -0.75E-7, and E

= 0.15E-10 (using Barin and Knacke's data). The valid temperature range for equations (1-5) and (1-7) is from 500°K Equilibrium Fractional CO Conversion 1.1 1.0 0.7 0.8 0.9 0.4 0.5 0.6 0.2 0.3 0.1 0.0

0 50 0 70 0 900 800 700 600 500 400

Figure 1-1. Equilibrium CO Conversion Versus Temperature. Versus Conversion CO Equilibrium 1-1. Figure Feed Gas Composition Gas Feed 2 H 22% A: Shift reaction only reaction A:Shift C: Shift reaction with carbonation at 25atm at carbonation with reaction C:Shift 1 at atm carbonation with reaction B: Shift % CO 8% 2 0 CO 30% 2 0 H 40% Temperature, °C Temperature, 2 O

11 to 1100°K. An expression for the overall equilibrium for the combined reactions is

Ke = KPlKPz = — (1-8) ^CCTHtO

The combined equilibrium is pressure dependent since removing C02 causes a decrease in the number of moles in the gas phase. The fractional CO conversion at equilibrium for the combined reactions (using the same feed gas composition) at 1 atm and 25 atm is shown as curves B and C, respectively, in Figure 1-1. From the diagram, we see that equilibrium CO conversion approaches 100% at 1 atm for temperatures below

800K, and at 25 atm, essentially complete CO conversion is feasible to about 900K, temperatures which are close to the temperatures of the outlet gas from a gasifier.

1.3.2. Research Objectives

A commercial process for the one-step production of hydrogen would probably be carried out in a dual fluidized-bed reactor system with continuous circulation of CaO from calciner to carbonation reactor such as shown in Figure 1-2.

One possible application of this process is in a CGMCFC system, where H2 and C02 are primary reactants in the anode and the cathode, respectively. A schematic diagram showing the proposed fluidized-bed process incorporated in CGMCFC is given in Figure 1-3 (Steinfeld et al., 1991). HYDROGEN PRODUCTION

Inert, CO2 -k Contlenser

Sorbent Shift Reaction Regeneration CO2 Capture C0 + H20<--> CO2 + H2 CaCOa --> CaO + CO 2 CaO + CO 2 —> CaC03

CaO

GasIner Sulfur-free Coal Gas

Figure 1-2. Proposed Fluidized-Bed Process for Hydrogen Production. Air Steam

Gasifier C02

CO HRSGI H2, CO Conversion H2 (CO) Heat Coal Carbonate Bottoming C02 C 02 Fuel Cell Cycle Steam Separation

Fuel Cell Exhaust D.C. Power A.C. Power

Figure 1-3. Advanced Gasification-Carbonate Fuel Cell System (from Steinfeld et al., 1991). 14

In this research, the experimental effort was carried out in a laboratory scale fixed-bed reactor system, with the objective of developing an experimental data base, defining process parameters, and determining the properties of CaO sorbents. The research, together with the work of Silaban

(1993) and Narcida (1992), was supported by the U.S.

Department of Energy under the project "A Calcium Oxide

Sorbent Process for C02 Removal".

The overall objective of this research was to determine the feasibility of a simultaneous water-gas shift, high temperature C02 removal process to produce hydrogen using a small fixed-bed reactor. Experiments were conducted to determine the effects of temperature, pressure, space velocity and inlet gas composition on system behavior during both calcination and shift/carbonation stages. Different CaO-based sorbents were tested and multicycle runs were conducted to examine sorbent durability at various reaction conditions.

Due to the continuous conversion of CaO to CaC03, the reactor operation was time dependent. A gas chromatograph system was used to perform on-line analysis of outlet gas concentration.

The breakthrough curves were used to compare the performance in terms of reactivity and durability. Optimal operating conditions suitable for future larger scale processes have been determined. Sorbent structural property changes in selected runs have been tested by using mercury porosimeter. CHAPTER 2

LITERATURE REVIEW

Three major areas are covered in the literature review.

First, the stages of development in the water gas shift

reaction are presented. Then various shift processes are

discussed. The third part concerns research on the CaO

carbonation reaction.

2.1. Water Gas Shift (WGS) Reaction

In 1870, Thaddeus Lowe first observed the reaction between carbon monoxide and water vapor (water gas) through coal gasification (Christian et al., 1949). Since then, the

WGS reaction has been subjected to many theoretical and experimental investigations due to its great industrial importance. The WGS reaction is one step in many industrial processes, such as hydrogen/ammonia production and CO detoxification. It is also used to regulate the H2 to CO ratio for methanol synthesis. The WGS is an exothermic and reversible reaction, and hence is thermodynamically favored at low temperature if higher conversion of CO is desired. In commercial processes, an operating temperature is chosen to provide the proper balance between reaction rate and CO conversion.

15 16

2.1.1. Homogeneous WGS Reaction

Although the WGS reaction is usually assisted by a

catalyst, the literature indicates that it could occur at

sufficiently high temperature without a catalyst. Gauthier

(1909) reported that "when CO is heated to about 1200°C in a porcelain tube with a variable excess of steam, hydrogen is

set free with formation of a volume of C02."

Kondrat'ev et al. (1943) studied the shift reaction in a quartz tube reactor in the temperature range of 600°C - 800°C.

They assumed the rate of reaction was first order with respect to both CO and H20, reported an activation energy of 5 - 6 kcal/mole, and proposed a mechanism for the shift reaction.

The value of activation energy is believed to be too low, probably due to the inability of the experimental system to measure the product gas concentration with sufficient accuracy.

April et al. (1969) compiled kinetic data for many reactions that occurred in the char zone of a charring ablator at temperatures up to 1250°C. For the WGS reaction, an activation energy of 30 kcal/mole and a frequency factor of l.OxlO12 cm3/cj-raole-sec were used in the reactor modeling.

2.1.2. Catalytic WGS Reaction

The vast majority of previous studies of the WGS reaction have focused on the catalytic option. Generally, shift 17

catalysts are divided into two categories: homogeneous

catalysts and heterogeneous catalysts.

Homogeneous Shift Catalysis

Homogeneous catalysts take the form of a fluid (usually

a liquid solution) and the intention is to enhance a reaction

at low temperature. Fenton (1970a, 1970b, 1973) was the first

to describe homogenous catalysis of the shift reaction in a

series of three patents which claimed that aqueous alcoholic

solution of a variety of Group VIII metals could promote the

WGS reaction.

Laine et al. (1988) presented an extensive review on this subject which focused on the chemistry of the homogeneous WGS reaction. Various mechanisms, complicated by many steps and intermediate chemical forms, have been proposed. The temperature range of interest is from room temperature to

200°C. Although increasing the reaction rate at low temperature is always exciting, the homogeneous shift catalysts are not competitive with heterogenous catalysts, which still dominate in industrial processes.

Heterogeneous Shift Catalysis

Many materials can catalyze the water-gas shift reaction, and patents relating to new shift catalysts appear frequently in the literature. Rofer-De Porter (1984) reports that a wide variety of metals including Rh, Co, Fe, Ni, Pt, Pd and Os have 18 catalytic activity. These metals may be used either in bulk or dispersed on supports such as alumina or silica.

Many metal oxides and metal sulfides will also catalyze the reaction. Most commercial catalysts are in the oxide form. For several decades, two types of oxide catalysts have been used almost exclusively in commercial reactors (Newsome,

1980) . One is the so-called high-temperature shift catalyst, consisting of chromia promoted iron oxide. The operating temperature is in the range of 350°C - 450°C. The second, a low temperature shift catalyst, is a mixture of copper and zinc oxides whose active temperature is from 200°C to 250°C.

MgO is a known shift catalyst (Gluud, 1931) but no references to CaO as a catalyst have been found.

Moe (1962) studied the kinetics of the WGS reaction using a high temperature Cr203-Fe203 catalyst. The reversible rate equation was expressed in a power law form and was first order with respect to each reactant. A 10 kcal/mol activation energy was reported.

Bohlbro (1969) pointed out that widely different and even conflicting results for the high temperature shift kinetics appear in the literature. One of the main reasons, he mentioned, was due to some kinetic data being taken under diffusion limitations. He correlated kinetic control data for a rate equation of the form;

r = \ [ co]1 [h 2o ] ^ [ co2]n [#3]g* (i -P) (2"1) 19

= tco21 [H2] (2 -2 ) 0 = — H iCtCO] [H20] where K is the equilibrium constant. The following values were obtained for the parameters: 1 = 0.9 - 1; m = 0.25 - 0.5? n = -0.6? q = 0 - -0.1. The activation energy was in the range of 14 - 30 kcal/mol.

Shchibrya et a l . (1965) considered the kinetics for a CuO

- ZnO based shift catalyst with a particle size of 0.1 to 1 mm which ensured no diffusion resistance. They found that the kinetics of the shift reaction was accurately describe by:

r _ J: (2-3)

where A is a constant and k is the rate constant with an activation energy of 27 kcal/mol.

Newsome (1980) discussed the structural and kinetic properties of these two catalysts and concluded that, for the high-temperature catalyst, reaction rate is proportional to

[H20]0,5. Consequently, a large amount of steam is needed in order to increase the reaction rate. Newsome (1980) also reported that the low-temperature Cu-Zn based catalyst is completely sulfur intolerant. Only 1 ppm of H2S will deactivate the catalyst.

Another material which is receiving increased attention as an industrial water-gas shift catalyst is sulfided cobalt oxide-molybdenum oxide on alumina (Overstreet, 1974). The 20 advantages of this catalyst are that it is insensitive to sulfur poisoning and its active temperature range covers that of both iron-based and copper-based catalysts. BASF and

Haldor-Topsoe have commercially marketed this kind of catalyst

(Ahn et al., 1985) .

2.2. Shift Processes

2.2.1. Conventional Shift Processes

There are three commercial shift processes commonly used today. All are catalytic processes and usually a high- temperature iron-based catalyst and/or low-temperature copper- based catalyst is used. When complete CO conversion is not necessary, a single stage reactor using the high-temperature catalyst may be employed. The exit gas composition approaches that corresponding to chemical equilibrium at the operating conditions. C02 may be removed by absorption, and the product is suitable, for example, for methanol synthesis. The other two processes may be employed to achieve essentially complete

CO conversion. Both begin with a first stage reactor using the high-temperature shift catalyst. In the first process, the C02 gas produced in the first-stage reactor is removed by scrubbing, and the C02-free gas is fed to a second stage reactor at conditions similar to the first stage. Due to the intermediate C02 removal, the water-gas shift equilibrium is altered and essentially complete CO conversion is reached in 21 the second stage. Kohl and Riesenfeld (1979) discussed this process ■ and a flow diagram is given in Figure 2-1.

Alternately, the product from the first stage reactor may be cooled (without C02 removal) and fed to a second reactor containing low-temperature shift catalyst. From the previous thermodynamic analysis, we know that the water-gas shift reaction is favored at low temperatures. The reduced temperature in the second stage is sufficient to achieve the desired CO conversion level without intermediate C02 removal.

Figure 2-2 (Rase, 1977) shows a diagram of this process. An external heat exchanger is used for temperature adjustment.

In both processes, in order to obtain high purity H2, a

C02 separation process is needed. In addition to C02 removal by absorption, pressure swing adsorption has been widely used for this purpose. In a pressure swing process, a zeolite molecular sieve is usually employed to adsorb C02 at high pressure, and desorption is accomplished by reducing the pressure at a constant temperature and then purging the sorbent at low pressure. Compared with the conventional wet scrubbing process, pressure swing adsorption can result in very high purity H2 (99.9%) since H2 is essentially nonadsorbed. However, the venting and low pressure purge operation cause a greater hydrogen gas loss (Kohl and

Riesenfeld, 1985; Ruthven, 1969).

Both high-temperature and low-temperature catalysts are temperature sensitive. Since the typical outlet gas Gas Out H 20 H20 Steam Solution

Cooler Cooler Solution U \t

y A / % Steam

Solution Solution To Out Regenerator Gas In Fuel 1st Stage 1st Stage 2nd Stage 2nd Stage 2nd Stage CO Converter C 02 Absorber Heater CO Converter C 02 Absorber

Figure 2-1. The Shift Process Using Two High - Temperature Reactor With Intermediate C02 Removal (from Kohl and Riesenfeld, 1977). C02

Heat recovery

Water Heat recovery Separator

CO shift C02 C02 converter absorber stripper

Figure 2-2. The Shift Process Using A High - Temperature Reactor Following by A Low - Temperature Reactor (from Rase, 1977). 24 temperature from a hydrocarbon reformer is 850°C (Vannby,

1991) , and that from a coal gasifier is about 600-900°C

(Heller, 1983), initial gas cooling is required, which is not favorable in terms of efficiency.

2.2.2. Novel Shift Concepts

Another trend in the shift process which is receiving great interest is coupling the catalytic shift reaction with

H2 separation using a membrane reactor. Uemiya et al. (1991) studied the shift reaction assisted by a palladium membrane at

400°C. It was reported that the membrane reactor provided levels of CO conversion beyond the normal equilibrium values, and that very high purity H2 was produced. The stability of the membrane and the gas loss during separation are two major concerns for the process.

Gluud et al. (1931), in their patent, reported a high temperature fixed-bed reactor process which employed calcined dolomite to achieve simultaneous shift conversion and C02 removal for H2 production. The overall reaction is

[CaO + MgO] + CO + HzO * [CaC03 + MgO] + H2 (2"4)

MgO served as a shift catalyst. Removal of C02 from the gas phase by reaction with CaO permitted the reaction to go to completion at temperatures as high as 700°C.

Squires (1967) identified a number of disadvantages associated with Gluud's fixed-bed process. First, because the 25 shift reaction is exothermic, significant temperature increases will occur in the fixed-bed and careful control of the reactor temperature is required. Gluud et al. proposed to control the temperature by interrupting the process and cooling the partially reacted sorbent. On the other hand, calcination of the spent sorbent - CaC03 - is endothermic and

Gluud et al. proposed that the required energy be supplied by direct contact fuel combustion. Large quantities of excess air were required to keep the calcination temperature below a level at which solid reactivity would suffer. Squires (1967) pointed out that these limitations contributed to the poor thermal efficiency and the cyclic nature of the fixed-bed system. He suggested a fluidized-bed process to overcome these difficulties. A two-bed process with continuous transfer of spent sorbent from the shift reactor to regenerator and recycle of the regenerated sorbent in the opposite direction would result in a steady state process.

Hot-spotting and thermal excursions would be eliminated since the fluidized-bed would be effectively isothermal. Squires also proposed calcined dolomite as the C02 sorbent and concluded that dolomite is sufficiently durable to withstand fluidization. But neither Gluud nor Squires reported any experimental data.

The C02 acceptor process (Curran et al., 1967) was based upon the gasification of bituminous coal char* and used calcined dolomite as the carbon dioxide acceptor. Half­ 26 calcined dolomite was regenerated in a separate reactor and then was transferred to the gasifier. In the gasifier, CaO also captured sulfur by reacting with H2S to form CaS. It was found that the shift reaction and the reaction of CaO with C02 occurred simultaneously at the temperature range of 813°C-

890°C.

2.3. Fundamental CaO Carbonation Studies

Both the kinetics and the structural changes in the calcination and carbonation reactions have been reviewed.

2.3.1. Calcination/Carbonation of CaO based Sorbents

Dedman and Owen (1962) studied the reaction of C02 and calcined limestone in the temperature range of 100-600°C at the pressure range of 10 to 600 mmHg. The reaction was found to be independent of C02 pressure since the pressure range was far from the equilibrium pressure of the Ca0-C02 reaction.

The reaction was initially very rapid followed by a much slower reaction. The slower second stage reaction was attributed to the diffusion of gases in pores. An activation energy of 9.5 ± 2 kcal/mol between 200 and 600°C was found for the second slower stage.

Barker (1973) examined the reversibility of the reaction between CaO and C02 at 866°C by using a thermogravimetric analyzer. The calcium carbonate particle size was in the range of 2 to 20 /an. Multiple cycle runs (up to 40) were 27

conducted using limestone as a sorbent and typical results are

given in Figure 2-3. Calcination was always complete. Each

recarbonation started with a very rapid initial period

followed by an extremely slow rate stage. The transition from

the rapid to the slow reaction stage occurred at progressively

lower fractional carbonation values with increasing cycle number. After about 24 hours in a single cycle, the

carbonation, as shown in the figure, was almost complete.

Several researchers have provided explanations for such

reaction behavior (Barker, 1973, Silaban, 1993). Natural limestone is essentially nonporous. During calcination, pores are created within the solid as C02 is driven off. Upon

carbonation these pores are refilled. Carbonation occurs preferentially at the particle exterior and the end of the rapid reaction stage corresponds to total surface pore closure. The remaining unreacted CaO within the particle is carbonated very slowly since further reaction is controlled by solid state diffusion of C02 through the nonporous carbonate layer.

In order to determine the critical thickness of the CaC03 layer which was formed before diffusion became controlling,

Barker (1974) chose a calcium carbonate particle size of 20 nm to study the recarbonation reaction. Calcination was performed at 629°C under N2 and recarbonation at 577°C under one atmosphere of C02. At these conditions, 93% V/sighl of sample (V.) 100 90 00 70 Figure 2-3. CaC03 Calcination and Recarbonation and CaC032-3.Calcination Figure 60 50 . a 2 r recarbonation (a)hr. 24 b mlil hr yls (Barker, 1973). cycles short (b) multiple 29 recarbonation was achieved during the rapid reaction phase and this value remained essentially constant through 30 cycles.

However, the particle size of about 20 nm is impractical for commercial use.

Oakeson and Cutler (1979) concentrated on the diffusion- controlled reaction between CaO and C02 using a microbalance over a temperature range of 853 to 1044°C under C02 pressures between 2.35 and 24.89 atm. They found that the rate of reaction followed the parabolic correlation with time, especially at longer time. The pressure dependence of the reaction rate was related to a Langmuir-type adsorption isotherm. The diffusion activation energy was reported to be

29 ± 6 kcal/mol.

Bhatia and Perlmutter (1983) investigated the kinetics of the C02-lime reaction in a TGA under a mixture of C02 and N2 at one atmosphere over temperature range of 400 to 725°C. Using a random pore model (Bhatia and Perlmutter, 1980 and 1981), they found an average value of the rate constant of 0.0595 ±

0.0018cm3/gmol. s for the initial rapid reaction and an activation energy for the diffusion controlled reaction of

88.9 ± 3.6 kj/mol below 515°C and 179.2 ± 7 kj/mol above 515°C.

The temperature of 515°C was close to the Tamman temperature of CaC03 (about 523°C) suggesting a change in the solid state diffusion process at that temperature.

Dhupe and Gokarn (1987 and 1990) mixed dolomite with metallurgical grade silicon powder to study the effect of 30 inert solid on the carbonation reaction. They reported that there was an optimum inert concentration in a sorbent pellet which gave a maximum C02 uptake. However, they gave no clear explanation of the result. While the extent of reaction could be improved, the ultimate C02 capacity was decreased by adding the inert material to CaC03.

Mess (1989) used a TGA to investigate product layer diffusion in the carbonation of CaO under C02 pressures up to

12 atm at temperatures from 550°C to 1050°C by using nonporous

CaO particles. His experimental results showed that at temperatures of 600°C or above, the CaC03 product layer on the

CaO particles contained crystalline grains whose sizes varied from less than l /xm to as large as a whole particle (20 /an) .

At T > 900°C, the reaction rate was first order in C02 pressure. The activation energy in the steady state diffusion controlling region was 56.9 kcal/mol.

In his recent dissertation, Silaban (1993) studied the kinetics of the calcination and carbonation reactions as a function of temperature, pressure, C02 concentration, and background gas composition. Three sorbent precursors, calcium carbonate, hydrated calcium acetate and dolomite, were selected for kinetic studies. Multicycle runs were conducted to test sorbent durability. It was concluded that dolomite possessed better reactivity and durability than calcium carbonate, and that calcium acetate had comparable-reactivity to dolomite but poorer durability. A set of optimum operating 31 conditions, including a calcination temperature of 750°C and a carbonation temperature range of 650 - 750°C, were also reported. When results using a sulfur-free simulated coal gas were compared to results from a C02/N2 mixture, improved sorbent reactivity was observed, as shown in Figure 2-4. One possible explanation is that more C02 was produced by the water gas shift reaction so that the carbonation rate was enhanced.

2.3.2. Structural Changes during Calcination/Carbonation Reactions

As mentioned earlier, structural changes which occur within the reacting solid during the calcination-carbonation cycles are very important in understanding the behavior of the

Ca0-C02 reaction. In addition to chemical reaction, high temperature sintering can also cause changes in the structural properties (Caillet and Harrison, 1982). Sintering is the phenomenon in which fine particles in contact with each other agglomerate when heated to a suitable temperature.

Glasson (1958) measured the specific surface areas of CaO produced from Ca(0H)2, CaC03 and CaC204.H20 calcined at different temperatures. The surface areas of the precursors were 4.9, 2.5 and 1.9 m2/g , respectively. The corresponding maximum surface areas of the three calcines were 85.4 m2/g at

400°C, 43.2 m2/g at 750°C and 46.6 m2/g at 700°C. Sintering effects were also observed with time and temperature. Dimensionless Weight, W/Wo 1.1 1.0 0.8 0.9 0.7 0.6 0.5 e r u g i F 0 . Sorbent: Calcium Carbonate Calcium Sorbent: . Calcination: 750°C, N 750°C, Calcination: abnto:7QC atm 1 75Q°C, Carbonation: 2-4. Comparison of Carbonation Results Using Simulated Using Results Carbonation of Comparison 2-4. 2 olGsadC2N ny fo iaa, 1993). C02/N2 (fromand Silaban,Gas Coal Only 4 %C25%C/7 28 H20/N2 H2/8% C0/27% C02/55% 5% 6 2 atm 1 , 8 Time, minute Time, 0 2 4 6 8 20 18 16 14 12 10

NJ OJ Borgwardt et al. (1984) investigated surface area changes during calcination and carbonation of Fredonia White

Limestone. Originally, the limestone, containing 95% CaC03 and 1.3% MgC03, had a porosity of 8%, a BET surface area of

2.2 m2/g and a grain size of 2 /xm. After calcining at 700°C for 90 seconds, a BET surface area of 79 m2/g was obtained.

Sintering was observed at higher calcination temperatures and/or longer calcination times. At 700°C, a sintering period of 20 minutes reduced the surface area to 40 m2/g with a further reduction to 32 m2/g after 60 minutes. With a 30 minute sintering period, temperatures of 850°C and 950°C yielded 27 and 6 m2/g, respectively. In another paper,

Borgwardt (1985) measured the BET surface areas of two calcined limestones from different sources. Surface areas in the range of 50-60 m2/g were reported when calcination occurred between 850 and 1075°C.

Beruto et al. (1988) measured surface areas and porosities of decomposed reagent-grade CaC03, Ca(0H)2,

CaC204.H20, and Ca(CH3COO)2 powders. Table 2-1 lists the decomposition temperatures, specific surface areas and pore volumes of the CaO obtained. Carbonation of the resulting CaO was carried out at 685°C and 3.8 kPa. It was found that oxides from the acetate and oxalate samples reacted almost completely to form CaC03, and the high reactivity was attributed to the large pore volumes. Table 2-1

Surface Area and Pore Volume of CaO from Different Starting Materials*

Starting Decomposition Specific Pore Fractional Material Temperature Surface Volume of CaO (°K) Area (cm3/g) reacted (m2/g) after 4 hours

Ca(OH)2 581 137 0.248 0.65

CaC03 878 87 0.266 0.72

Ca(C2H302)2*H20 878 90 0.376 0.97

CaC204*H20 878 130 0.451 0.96

from Beruto et al., 1988. 35

Delucia (1985) studied the multicycle behavior of the calcination-carbonation reaction using C02 at atmosphere pressure over the temperature range of 50 to 800°C. The CaO particles were observed to undergo a volume decrease of 10-25% during calcination, which was used to explain the decrease in reactivity of CaO by 10-25% per cycle.

Narcida (1992) measured the structural properties

(surface area, pore volume and pore size distribution) of the three calcium-based sorbent precursors studied by Silaban

(1993) -- reagent grade calcium carbonate, reagent grade calcium acetate hydrate and dolomite -- at various stages through two complete calcination/carbonation cycles.

Calcination was conducted at temperatures of 900°C and 750°C in a N2 atmosphere. Carbonation was conducted at temperatures of 550°C and 750°C in 15% C02/85% N2 at 1 atmosphere.

Sintering of the calcined sorbents was tested at temperatures of 750°C, 900°C and 1000°C and times of 10 minutes, 1 hour and

6 hours. Typical pore size distributions of the three sorbents at different reaction stages are shown in Figures 2-5 to 2-7. The surface area and pore volume after different stages of calcination and carbonation are given in Table 2-2.

The following conclusions were reached from Narcida's

(1992) research. Calcined sorbents have higher surface area and pore volume in comparison with precursor and carbonated sorbents. Submicron pores were created by calcination of the calcium carbonate while both submicron and micron-size pores I

2.00

As Received

1.60 — First Calcination p

a

— •> 0.00 rrtit 0.001 0.01 0.1 1.0 6.0 DIAMETER (MICRON)

Figure 2-5. Pore Size Distribution of Reagent Calcium Carbonate at 750°C for 1 hour (from Narcida, 1992) . U> o> A: As Received B: Calcined at 300C C: Calcined at 550C D: Calcined at 750C a

0.001 0.01 0.1 1.0 6.0 DIAMETER (MICRON)

Figure 2-6. Pore Size Distribution of Reagent Calcium Acetate at Various Decomposition Stages (from Narcida, 1992). OJ -j I iI

•• As RtceWed

First Calcination

0.01 0.10.001 1.0 6.0 DIAMETER (MICRON)

Figure 2-7. Pore Size Distribution of Dolomite Calcined at 750°C for 1 hour (from Narcida, 1992). IP CO 39 Table 2-2

Structural Properties of Test Sorbents

CaC03 Ca-Acetate Dolomite

Surface Area (m2/g) Initial 0.9 3.8 1.7

First Calcination1 18.5 23 .2 14 .4

First Carbonation2 1.1 3.8 6.4

Pore Volume3 (cm3/g)

Initial 0.00 0.06 0.05

First Calcination1 0.25 0.96 0.40

First Carbonation2 0.00 0.15 0.15

Second Calcination 0.19 0.79 0.38

from Narcida, 1992

Calcined at 750°C, 1 atm in N2 for 1 hour

Carbonated at 750°C, 1 atm, 15% C02/N2 for 1 hour

Pore diameter range from 0.02 to 1.0 microns were created in the calcium acetate and the dolomite. The carbonation step produced complete pore closure for the calcium carbonate, complete closure of submicron pores and partial closure of micron-sized pores for the acetate, and partial closure of both submicron and micron-size pores for the dolomite. Longer sintering time and higher sintering temperature systematically reduced the surface area. However, pore volume was not affected by the sintering conditions. CHAPTER 3

EXPERIMENTAL APPARATUS AND PROCEDURE

This chapter discusses the experimental equipment and procedures used in this research. The first section describes the reactor system, including synthesis gas feed system and the fixed-bed reactor. Difficulties encountered in running the system and stages of modification are also covered. The second section discusses the specifics of the gas chromatography system for real time composition analysis. The major sorbents used in the experiment are presented in section three, followed by a brief description of the equipment used for sorbent structural analysis. Finally, the procedure for a complete run using the fixed-bed reactor is given in section five.

3.1 Fixed-bed Reactor System

This research is experimental, and a major challenge has been the design of an effective and operable fixed-bed reactor system, which includes synthesis gas simulation, a fixed-bed reactor and parameter control devices. A schematic representation of the system is shown in Figure 3-1. Gaseous feed components were obtained from high purity gas cylinders regulated by high pressure regulators. Porter Instruments

Model 201 high pressure mass flow controllers were used to control flow rate, A Matheson Model 6183 filter and a dryer

41 BPR - BACK PRESSURE REGULATOR C V-CHECK VALVE F-FILTER MFC - MASS FLOW CONTROLLER SL PI - PRESSURE INDICATOR CV MFC F C O N D - CONDENSER PRV-PRESSURE RELIEF VALVE D - DRYER

CTHreFTi ' JL

FURNACE

PRV jJ* wnW«il MFC F 3 -W A Y VJ T f VALVE M l y ~ C 3 j j

VALVE YRINGE PUMP COND COND (HjO)

3 -WAY VALVE TO GC

Figure 3-1. Schematic Diagram of Fixed-Bed Reactor System. 43 were placed in each gas line to prevent foreign material from contaminating the sensitive mass flow controller. Check valves and on/off valves were installed to prevent back flow due to possible pressure change during the experiment. Liquid water was added using a high pressure syringe pump (Harvard

Apparatus Model 909). Feed lines were heat traced using voltage control devices (Cole-Parmer Instruments) to insure that the water vaporized as it mixed with the permanent gases.

Reactant gases entered from the bottom of the reactor, were preheated while flowing upward in the annular space between the pressure vessel and the reactor insert, and then flowed downward to reach the sorbent. After reaction, the product gases from the reactor were directed through a condenser, a back pressure regulator (BPR) , and a manual valve to the GC system. During sorbent calcination, nitrogen flowed into the reactor and the feed gases were prepared through a side stream, which contained an identical condenser and a BPR.

The reactor consisted of three major parts: pressure vessel, reactor insert and an ungrounded thermowell containing six axially spaced thermocouples. The details including dimensions are shown in Figure 3-2. All parts were made of

316 stainless steel. Sorbent inside the reactor insert was divided into four axial sections by porous stainless steel discs. This configuration permitted each section to be removed separately, so that the extent of reaction and structural properties could be characterized as a function of rci

Sorbent

Quartz Wool

to C 9

Reactor Gas Inlet

CO

’zzzz. Pressure Vessel TC6

Insert Tubing

Porous Disk ** Reactor Gas Outlet

Thermowell

Figure 3-2. Refined Schematic of the Fixed-Bed Reactor. axial position. The additional section below the sorbent held quartz wool whose objective was to minimize the possibility of small particles of sorbent passing through the exit gas lines.

The thermowell containing six K-type thermocouples was positioned through the center of the sorbent bed. The top thermocouple measured the temperature of the preheated gases before they contacted the sorbent. The next four thermocouples were positioned at the approximate mid-point of each section of the sorbent bed, thus permitting the progress of the exothermic reaction front through the packed bed to be monitored directly. The sixth thermocouple was positioned in the product gas near the O-ring pressure seal, and was monitored to prevent overheating in that critical area. A multiple pin plug and K-type thermocouple wire were used to transfer the thermocouple signal to the recording device. An

IBM computer was used to record the temperature data every minute during a run using a Strawberry Tree hardware interface.

A single zone split-tube furnace (Applied Test System

Series 3210) equipped with a temperature controller (Model

2010) and CFE Model 2040 limit controller was used to maintain the reactor temperature. The temperature controller contains a microprocessor and is programmable with capacity of up to 8 ramp-and-soak intervals and up to 254 cycles. The limit controller is designed to shut down the furnace when the furnace temperature exceeds 1000°C. The system pressure was 46 controlled by two back pressure regulators (Tescom Series 26-

2300) , one for the reactor system and one for the feed gas preparation (side stream). Condensers were used to condense water from both feed gas and product gas lines before entering the GC sampling loop.

This fixed-bed reactor system is applicable to the study of any gas-solid reaction. The reactor has a capacity of 10 to 15 grams of solid, and has temperature and pressure limits of 1000°C and 300 psig, respectively.

The fixed-bed system underwent many modifications due to difficulties encountered during the experimental program.

Several important revisions are discussed here. The reactor system shown in Figure 3-1 is the final version which has functioned very well.

Reactor Loading/Unloading

In the original reactor design, a 316 stainless steel tube which was split axially was used inside the reactor insert to support the bottom disc (Figure 3-3) . Accessory equipment was designed to load and unload sorbent by moving the reactor insert relative to the thermowell. Due to the large resistance between the sorbent and the reactor inside wall, difficulties were encountered in collecting individual sorbent sections after high temperature, high pressure runs.

Several modifications were tried, including designing an electrically controlled mechanical loading/unloading system, 47

1.5'

rci

Sorbent Sections Quartz Wool

Reactor G as Inlet TC6

CO

7777?. Pressure Vessel

Insert Tubing Reactor Gas Outlet Support Tubing

Porous Disk iTbenmowell

Figure 3-3. Diagram of the Original Fixed-Bed Reactor. 48 but all of them failed because of the large ratio of reactor length to diameter. The reactor was modified to the design shown in Figure 3.2, and a new way of loading/unloading sorbent was developed. Since then, no problems in loading and unloading have been encountered.

Dead Volume

Unsteady state operation is an intrinsic property of any fixed-bed gas-solid noncatalytic reaction system. Hence, minimizing the dead volume is very important if the product gas analysis is to reflect the current conditions in the reactor. Using small size tubing (1/8") and replacing large condensers with small ones proved to be very effective.

Further reduction of the dead volume was realized by removing filters from the product gas lines. Some dead volume was unavoidable and there was still an appreciable time lag between adjusting the valve position to introduce reactive gases to the reactor and the time those gases reached the chromatograph sampling valve. A series of nonreacting tracer response tests were conducted to evaluate the delay time, tD, as a function of space velocity, pressure, and temperature.

Results of these tests are detailed in the next chapter.

Steam Condensation

High pressure steam used in this process . has given various problems. First, all the lines where the steam possibly appeared had to be heat traced and the temperatures had to be carefully monitored in order to prevent condensation. In the original design, only check valves were used after mass flow controllers. It was found that the check valves were not very reliable in preventing backflow of water vapor. Due to pressure variation during the experiments, water vapor flowed back to the mass flow controllers, clogged the orifices, and caused them to malfunction. New manual valves and careful flow handling methods (described in the experimental procedure) were developed to solve the problems.

3.2 Gas Chromatography System

The analytical train consisted of a Shimadzu Model GC-14A gas chromatograph equipped with an automatic sampling valve, dual columns, and both thermal conductivity (TCD) and flame ionization (FID) detectors. Ultra high purity argon was chosen as the carrier gas. A Carboxen 1000 column was used to separate H2, N2, CO, and C02; optimum operating conditions and valve timing were found by trial and error. A HayeSepN column was used to absorb the small amount of water vapor which might escape the condenser. A schematic diagram showing the steps of GC operation is given in Figure 3-4. A manually operated four-port valve was employed before the automatic ten-port sampling valve to switch between feed and product gases. As shown in Step 1, reactor product gas flowed into the sampling valve, through the sample loop, back in the sampling valve, ►i ctx r Cl

VO VO VI

C2 C3 TCD V*ri

> W A Y •WAT ,K2

HI GC OPERATION — STEP ONE GC OPERATION — STEP TWO [n o

C1, C2: Carrier G ases (Argon) VO CBX: Caboxen 1000 Column HSN: HayeSep N Column

C2 It c o TCD: Thermal Conductivity Detector FID: Flame Ionization Detector ,H2 METH: Methanator

K2 A ir GC OPERATION— STEP THREE

Figure 3-4. Schematic Diagram of GC Sampling Sequence. tJT O and to vent. The sampling loop was maintained at a constant temperature so that the same amount of gas was picked from the gas stream. Carrier gas Cl passed through the ten-port valve, and through the Carboxen column and the detectors. Carrier C2 flowed into the ten-port valve, through the HayeSepN column, back to the ten-port valve, and to vent. In Step 2, gases to be analyzed were first directed to the HayeSepN column which served as a guard bed to trap trace quantities of water vapor which escaped the condenser. When the permanent gases reached the Carboxen column, the automatic sampling valve was switched to the position shown in Step 3. While permanent gases were separated in the Carboxen column, the HayeSepN column was backflushed by argon to prevent build-up and leakage of H20 vapor into the Carboxen 1000 column.

The sample gases were separated in the Carboxen column and passed to the thermal conductivity detector. The TCD was used to detect H2, N2, and large concentrations (a 0.5%) of CO and C02. When the concentrations of CO and C02 were below the detection limit of the TCD, -column-—effluent was - directed through a ruthenium-catalyzed methanizer where CO and C02 were converted to CH«, and analysis was completed using the more sensitive FID. Ultra high purity hydrogen was fed into the methanizer and the flow rate was adjusted so that all the carbon oxides were converted into methane for FID analysis.

In order to prevent carbon deposition in the methanizer when the product contained high concentrations of CO and C02, a 52 three-way valve was placed between the TCD and the methanizer to switch off the FID when the carbon oxide concentration exceeded about 1%. This system permitted the continuous analysis of CO and C02 from low ppm levels through several percent. Some uncertainty was present near the 0.5% range where the TCD sensitivity was limited, and the quantity of CO and C02 tended to overwhelm the capacity of the methanator.

Shimadzu EZChrom Version 5.0 software was used for GC data acquisition and processing. Reactor product gases were sampled every 8.5 minutes, which proved to be frequent enough to show the component breakthroughs under most operating conditions.

Seven standard gas mixtures consisting of CO, C02, H2 and

N2 from Matheson Inc. were used for GC calibration.

Recalibrations were conducted several times when signs of shifting of the GC operating conditions were detected.

Detailed calibration results will be presented in the next chapter, while two examples are discussed here.

A typical chromatogram for a calibration gas mixture consisting 10% H2, 70% N2, 10% CO and 10% C02 is shown in

Figure 3-5 (a) and (b) . The four peaks in the TCD response clearly show that the components were separated. However only the TCD was used in this case because CO and C02 concentrations were too large for the methanator. Figure 3-5

(c) and (d) show both TCD and FID responses for a gas mixture with 3% H2, 96.8% N2, 0.1% CO and 0.1% C02. In this case, only S3

Flame Ionization Detector

25000. (a) u Gas Composition: v 0 10% H2 10% CO 1 10% C 02 70% N2 0J t s

io 2.5 5.0 7.5 Minutes Thermal Conductivity Detector 20000 . 15000. (b) u 10000. Gas Composition: v 5000. 1, W, 10% H2 10% CO 0 0 1 . CO 10% C 02 70% N2 t -5000. s -10000. ■15000. NJ VJ S' •20000. 0 2.5 ... 5.0 7.5 Minutes

Flame Ionization Detector

(e) u Gas Composition: V 3% H2 0.1% CO 0 25000. CO 0.1% C 02 96.8% N2 1 t s V s ■ ... ✓A! — 0. 0 2:5 . 5:0 7.5 Minutes Thermal Conductivity Detector 20000 15000. w u 10000. Gas Composition: V 5000. A/* 3% H2 0.1% CO 0 0 Hi 1 . 0.1% C 02 96.8% N2 t -5000. s -10000. -15000. V r -20000 2 ; 5 ... 5.0 7.5 Minutes

Figure 3-5. Chromatogram for Two Gas Mixtures. 54

CO and C02 were detected by the FID, while only H2 and N2 were shown on the TCD because of the low concentration of CO and

C02. At 3.2 minutes, the sampling valve was switched back to its normal position.

Figure 3-6 shows a typical reactor response in terms of

C02 concentration (TCD analysis) and reactor temperature history during calcination and combined reaction stages. As received dolomite was heated at a rate of 5°C/min under N2 at

3.3 atm. Calcination began when the temperature reached about

520°C. The first peak in curve B of Figure 3-6 corresponded to the calcination of MgC03, and the second peak to calcination of CaC03. The C02 material balance was checked by comparing the total C02 moles detected during calcination with the initial C02 content in the sorbent. The C02 material balance typically closed to within ±5%. When calcination was completed, the reactor was cooled to 550°C and maintained at that temperature while the feed gases were pressurized to 15 atm and analyzed by the GC, as shown by the third "peak" of curve B. At t = 354 minutes, the feed gases were switched to the reactor and the combined shift and carbonation reactions began. Figure 3-7 gives a better representation of the reactor response during this carbonation/shift reaction stage.

As shown in Figure 3-7, during the earlier time period

(prebreakthrough), CO and C02 concentrations (dry basis) were effectively constant at about 80 and 300 ppm, respectively, which were below the TCD detection limit and hence only shown C02 Mole Percent (dry basis), TCD Response . Test 54 (dolomite) 54 Test . _ P = 15 atm 15 = P _ Figure 3-6. The Typical Reactor Response for a Complete a for Response Reactor Typical The 3-6. Figure SV = 1425 hr_1 1425 = SV T = 550°C = T

Cycle Test. Cycle ** 0 300 200 1 Time, min. Time, r ■! 93 N 59.3% edGsCmoiin . Composition: Gas Feed .% O .% CO 3.1% CO 6.6% .% H 4.6% 2 2 64 H 26.4% 2 ' O 2 "

Temperature, °C ur tn Component Concentration (dry basis), ppm 105 104 103 102 101 2 2 2 2 Figure 3-7. TCD and FID Results of Product Gas Concentrations Gas Product of Results FID and TCD 3-7. Figure

Prebreakthrough 8 0 2 4 6 8 500 480 460 440 420 400 380 FIDAnalysis Analysis TCD 1 ____ 02 C 1 ____ During Shift/Carbonation Reaction Stage. Reaction Shift/Carbonation During 1 ____

1 1 j ie minuteTime, ratruh Postbreakthrough ^reakthroughj ___ 1 ______1 ______1 ______93 N 59.3% Feed Gas Composition Gas Feed .% H 4.6% .% O .% CO 3.1% CO 6.6% Test 54 (dolomite) 54 Test 1 ______-i—i i=t = i i- — i - i SV = 1425hri = SV T = 550°C = T P = 15atm P= 1 ______2 2 64 H 26.4% 1 ______

1 ______2 O 2

cn in 57 by the FID. Hydrogen concentration was at its maximum during this time. CO and C02 breakthrough began at t = 425 minutes and the concentrations increased rapidly until a second steady state was reached at about t = 455 minutes. A decrease in the

Hj concentration accompanied the CO and C02 breakthrough. The beginning of the postbreakthrough period corresponded to effectively complete conversion of CaO to CaC03, and only the shift reaction occurred during this period.

Figure 3-8 shows the movement of the temperature front through the sorbent bed in a plot of temperature deviation from the set point as a function of time. The temperature in the top section of the packed bed was approximately 10°C higher than the bottom section temperature, and the combined shift and carbonation reactions (both exothermic) produced a temperature increase of about 15°C as the reaction front passed each thermocouple position.

3.3 Sorbent Precursor Description

Dolomite (obtained from National Lime Co., Findlay, Ohio) was chosen as the primary sorbent on the basis of favorable carbonation results in previous studies (Siliban 1993, Narcida

1992) . A brief description of the sorbent composition is given in Table 3-1. Several other materials were also tested for comparison purposes. The dolomite was received in relatively large chunks. These materials were crushed and screened with different size rages retained for reactor tests. Temperature Deviation from Set Point, °C -10 -20 10 20 5 30 9 40 430 410 390 370 350 - Test 54 (dolomite) 54 Test Figure 3-8. Reactor Temperature Response During Response Temperature Reactor 3-8. Figure Shift/Carbonation Stage. Shift/Carbonation ie minuteTime, 93 N 59.3% .% O .% CO 3.1% CO 6.6% .% H 4.6% Feed Gas Composition Gas Feed 5 470 450 SV = 1425hr1 SV T = 550°C = T P = 15atm = P 2 2 64 H 26.4%

490 2 O 2 tn 00 Table 3-1

Chemical Analysis of Dolomite (as reported by National Lime Co./ Findlay, Ohio)

Component Weight %

CaC03 54.5

MgC03 45.0

Si02 0.2

Fe203 0.07

A1203 0.08

S 0.03

Other 0.12

Loss on ignition after calcination at 1800°F - 47.4% 60

For most runs, dolomite having a size range of 149-210 nm was used, although other particle size ranges were also tested.

3.4 Sorbent Structural Property Measurement

The structural property measurements focused on pore volume and pore size distribution. Only selected sorbent cases were chosen for structural tests due to the earlier detailed study of structural properties by Narcida (1992).

An Autopore II 9220 porosimeter from Micromeritics

Instrument Corporation, which used the mercury intrusion method, was employed to determine the total pore volume and pore size distribution. The porosimeter has two subsystems - low pressure/vacuum and high pressure (Narcida 1992) . ]h the low pressure/vacuum subsystem, the sample was first evacuated to 50 /imHg by a vacuum pump. Then appropriate valves were opened and mercury from the reservoir was introduced to fill the sample. The air pump was used to increase the sample pressure to 30 psia, so that inter-particle voids were filled with mercury. The mercury traps were employed to drain excess mercury for safe operation. In the high pressure subsystem, the sample filled with mercury was exposed to high pressure through the hydraulic pump and the intensifier. The raw data recorded was the intrusion volume as a function of pressure.

The porosimeter has a capability of detecting pores in a range of 0.003 to 4000 j*m. However, pores between 0.01 to 6 f m were of interest of this study. Figure 3-9 gives typical Cumulative Volume (ml/g) 1.2 1.1 1.0 0.8 0.9 0.7 0.6 0.5 Figure 3-9. Pore Size Distribution from Porosimeter for Porosimeter from Distribution Size Pore 3-9. Figure a l l i i m Calcined Dolomite. Calcined Diameter (micron)Diameter CalcinedDolomite 1.4 1.2 0.8 0.6 0.4 0.2 0.0

Log Diff. Vol (dV/dLogD) (ml/g) M 62 results for calcined dolomite, where cumulative pore volume and log differential pore volume are plotted versus pore diameter. As clearly shown in Figure 3-9, the calcined dolomite has a pore size range of 0.02 to 0.1 /xm with a peak in about 0.08 /xm.

3.5 Experimental Procedure in Running the Fixed-Bed Reactor System

Four portions of the sorbent precursor were weighed using a Sartorius balance and each portion was added to one section of the reactor (~ 2.6 cc). A specially designed device was used to assist in loading the sorbent into the reactor after a quartz wool section had been put in the bottom. The sorbent was packed as uniformly as possible to prevent any large internal spaces. Then the reactor insert, together with the thermocouple well, was connected to the pressure vessel with o-ring seals used to maintain operating pressure. N2 from a high pressure cylinder was fed at 800 seem using a mass flow controller to build up the system pressure to 15 atm after all the lines had been connected. A soap bubble detector was used to check for leakage, particularly around the reactor coupling area. If no leak was detected, the system pressure and N2 flow rate were readjusted to operating conditions for calcination. If the feed gas contained water, heating tape was wrapped around the reactor coupling area to insure that no water condensed there. After the desired pressure and flow rate were achieved, power was supplied to the furnace to initiate heating at a rate of approximately 5°C/minute. There were two stages in a typical one cycle run: calcination stage and combined carbonation and shift reaction stage. Calcination usually took 3 to 4 hours depending upon the temperature and flow rate. During calcination, reactive gases for the shift- carbonation reaction stage were prepared using the side stream. Permanent gases including CO, H2, C02 and N2 came from separate high pressure cylinders and flow rates were adjusted according to the desired operating conditions. The gas line was heated to insure vaporization of water. After the temperature along the gas line was steady, liquid water was supplied by the precalibrated high pressure syringe pump and was vaporized as it mixed with the permanent gases. The product gas during calcination was directed to the GC for composition analysis. At the end of calcination, the power supply to the furnace was shut off and the temperature controller was reprogrammed to the desired temperature for the shift-carbonation reaction stage. Both the reactor system and feed gas pressures were readjusted by using separate back pressure regulators to the same value. All heated areas were rechecked to insure water vaporization at the new pressure.

Before the start of the shift-carbonation reaction, feed gas was directed to the GC to check its composition. 64

When the system temperature, pressure, and feed gas composition reached desired levels, feed gas was directed into the reactor and product gas was directed to GC.

During the early stages, CO and C02 concentrations were closely monitored using the FID (Figure 3-7). CO and C02 concentrations were generally small during this period and were below the TCD detection limit. At breakthrough, the CO and C02 concentrations began to increase and when the concentrations were above 1%, the FID was switched offline, and the TCD was used for further analysis. 3 to 3.5 hours was the typical time span for the combined reactions.

At the end of the test, the reactive gases were switched to the side stream and a flow of N2 (500 seem) was used to purge the reactor. At the same time, the water supply was stopped, the side stream pressure was slowly reduced, and the

N2 flow rate was increased to prevent possible back flow of vapor to the mass flow controllers. When the side stream pressure reached 1 atmosphere, the flow of CO, H2, and C02 was stopped and the on/off valves in these lines were closed to isolate the mass flow controllers. After enough time (about

30 minutes) for the water vapor to be purged from the reactor, the power supply to the furnace was shut off. The reactor system was depressurized when it cooled to about 400°C, a temperature at which calcination would not occur.

After reaching room temperature, the reactor-insert was removed from the pressure vessel, and the sorbent was removed 65

section by section, and stored for possible subsequent

structural property analysis. Liquid water in the two

condensers was also collected for material balance check. The

reactor was cleaned and dried for future runs.

For some tests where the calcination gas contained

nitrogen and water, both N2 lines were used during calcination

to prevent water vapor from entering the inactive feed gas

lines. The water was introduced to the reactor only when the

reactor temperature reached the point at which the water would not condense.

The experimental procedure described above was followed

for one complete calcination and combined reaction cycle.

Many of the experimental tests were continued for several

cycles. The procedure was similar but involved additional

switching of reactive gases between the reactor and the by­ pass lines between the calcination and combined reaction cycles. Because one single cycle run usually took about 12 -

14 hours, multicycle tests had to be conducted over several days. At the end of one cycle, the sorbent stayed in the reactor overnight at 15 atm and about 500°C in a N2 atmosphere.

These conditions were set to ensure that no calcination took place between cycles, and that the system temperature was well below the temperature at which sorbent sintering could occur. CHAPTER 4

EXPERIMENTAL RESULTS AND DISCUSSION: PRELIMINARY TESTS

This chapter presents the results of the preliminary studies which include the temperature response of the fixed- bed reactor under nonreacting conditions, chromatograph analytical method development and calibration, reactor system dead volume evaluation, and a series of initial reaction tests. These studies set the basis for future runs, provided important parameters for analysis of results, and showed the capability of the experimental system under a wide range of conditions. Preliminary reaction tests focused on calcination and carbonation of calcium oxide based sorbents. Reaction conditions were chosen based on the previous research by

Silaban (1993). Commercial dolomite was used as sorbent precursor for most of the runs and marble chips were tested for comparison purposes. Finally, one test showing the feasibility of combined shift and carbonation reactions is reported.

During the preliminary tests, a great deal of effort was devoted to refining the reactor system and developing a reliable chromatograph analytical technique. At the end of the preliminary tests, a carefully organized experimental procedure had been developed and the experimental system has worked well since then. Because the experimental results of the preliminary reaction tests may involve some errors

66 67 associated with the immature experimental technique, they were used only as a guide for choosing conditions for the future runs.

4.1 Reactor Temperature Response

Both the water-gas shift reaction and the carbonation reaction are exothermic and the temperature profile along the fixed-bed reactor is important in understanding the reactor behavior. As the first step, a series of non-reaction tests were conducted to show the temperature response of the reactor under a preprogrammed heating sequence. The resulting temperature history was then compared with the actual temperature response under reaction conditions to obtain the temperature deviations due to the reactions.

Figures 4-1 and 4-2 show the temperature distributions at the various thermocouple positions in a non-reaction test where the set-point temperature was increased at 5°C/min to

660°C and held constant thereafter. 250 seem of N2 flowed through the reactor at 3.2 atm. In Figure 4-1, T£urn represents the furnace temperature, T2 represents the temperature of the preheated feed gas, and T2 through Ts represent the temperatures at four axial positions in the sorbent bed. T6 represents the temperature of the product gas after heat exchange with the feed (Figure 3-2).

During the early heating stages there is considerable lag between the set-point and the reactor temperatures. This lag Temperature, °C 100 500 700 300 400 200 600 10 0 300 200 100 0 Figure 4-1. Reactor Temperature - Time Response in Response - Time Temperature Reactor 4-1. Figure Tfurn Nonreacting Experimental Test. Experimental Nonreacting Time, minute

Temperature Deviation from Set Point, °C -20 -10 20 0 0 5 20 7 20 9 30 1 30 3 30 5 360 350 340 330 320 310 300 290 280 270 260 250 Figure 4-2. Temperature Deviation from Set Point in Point Set from Deviation Temperature 4-2. Figure AT 2 T A Nonreacting Experimental Test. Experimental Nonreacting 4 Time, minute a T3

is due to the large mass of stainless steel and is qualitatively similar to the lag observed in the earlier electrobalance tests (Silaban 1993) using the same kind of pressure vessel. Most of the lag was overcome by the time the set-point temperature was reached. After approximately 250 minutes the temperature of the preheated feed gas was within about 25°C of the set-point temperature, and the temperatures within the bed were, on the Figure 4-1 scale, approximately equal to each other and also equal to the set-point temperature. The temperature of the cooled product gas leveled out at about 40°C, well below the temperature limit for safe operation of the o-ring.

Figure 4-2 shows the temperature deviation from the set point, AT2 through AT5 during the latter stages of the test.

From this figure we see that an axial temperature gradient of approximately 6°C exists, with the temperatures ranging from

668°C near the top of the packed bed to 662°C near the bottom.

4.2 GC Analytical Method and Calibration

This research dealt with desulfurized coal gas containing

H2, CO, C02, N2, and steam. In order to study the concentration changes due to the simultaneous shift and carbonation reactions, a reliable gas analysis system was required. The following criteria were used in designing a chromatograph analytical strategy: 71

1. capable of separation and detection of H2, CO, C02, N2;

2. capable of analyzing CO and C02 down to 10 ppm level ;

3. reasonable analytical time for one sample;

4. tolerance of small amount of moisture.

Developing a chromatograph analytical technique achieving all these objectives proved to be no easy task. Much trial and error effort was required. Figure 4-3 shows the original design. Three multi-port valves, three columns, and a detector train equipped with a thermal conductivity detector

(TCD), a methanizer, and a flame ionization detector (FID) were employed for the gas analysis. This configuration was quite complicated and didn't function well due to the following reasons. First, the methanizer used nickel as a catalyst and was operated at 380°C. When high concentrations of CO and C02 (~ 5%) were analyzed, the methanizer was soon plugged by carbon deposition. Secondly, nitrogen used as the carrier gas has a thermal conductivity which is very close to that of CO. Hence CO couldn't be analyzed by the TCD.

Figure 4-4 represents the second stage development of the

GC system. Neither the methanizer nor the FID was used and argon gas was chosen as the carrier gas. The valve timing required to separate the four components was established, but the TCD did not have sufficient sensitivity to permit analysis of CO and C02 concentrations below about 1000 ppm.

The final chromatograph analytical method, .which was described in the Chapter 3, realized all the design objectives MS |—

Feed Gas

Product Ga:

Sample

Loop IflETr HS1, HS2: HayeSep Columns C1. C2: Carrier Gases MS: Molecular Sieve TCD: Thermal Conductivity Detector FID: Flame Ionization Detector METH: Methanator

Figure 4-3. GC Operating System (Stage 1). PP

HSN

V2

Vent C2 DJ. RESTR. TCD Sample Vent -/0<3cXJ u Loop

PP: Poropack Column HSN: HayeSep Columns C1, C2: Carrier Gases MS: Molecular Sieve TCD: Thermal Conductivity Detector

Figure 4-4. GC Operating System (Stage 2). 74 and avoided the problems mentioned above. Several key changes were made including:

1. ruthenium as the methanation catalyst;

2. argon as the carrier gas;

3. shortening the sample loop by half;

4. using a carboxen 1000 column and eliminating the six- port valve.

The operating conditions are summarized in Table 4-1.

Since ruthenium catalyzes the methanation reaction at a lower temperature (300°C) , the problem of catalyst fouling due to carbon deposition was significantly lessened. As shown in the

Figure 3-4, a 3-way valve was used to switch off the methanizer and FID when CO and C02 concentrations were above approximately 0.5%.

The gas chromatograph had to be calibrated before the analysis of reactor product gas. Seven standard gas mixtures obtained from Matheson were used for GC calibration. The compositions of the gas mixtures are summarized in Table 4-2.

Combinations of these standard mixtures shown in Table 4-3 were used to develop calibrations for both detectors. Higher concentrations were used for TCD calibration while lower concentrations were used for FID calibration. Both detectors were calibrated using intermediate concentrations to provide overlap. Quadratic constants relating mol fraction and detector response are summarized in Table 4-.4. Each calibration is based on a minimum of three concentration 75

Table 4-1

GC Operating Conditions

1. Carrier gas Argon

2. Temperature of VI and sample loop 55°C

3. Sample loop volume 0.156 cm;

4. Columns: HayeSep N 4 ft

Carboxen 1000 8 ft

5. Column temperature 85°C

6. TCD temperature 80°C

7. Methanizer temperature 300°C

8. FID temperature 250°C Table 4-2

Calibration Gas Compositions

(mol percent)

Standard ID CO C02 h 2 n 2 1 9.98 9.99 9.98 balance

±0.20 ±0.20 ±0.20

2 0.050 0.100 15.04 balance

±0.001 ±0.002 ±0.30

3 1.00 4.00 20.00 balance

±0.02 ±0.1 ±0.40

4 0.099 0.099 2.99 balance

±0.002 ±0.002 ±0.06

5 0.501 0.500 5.99 balance

±0.010 ±0.010 ±0.10

6 0.999 0.999 11.90 balance

±0.020 ±0.02 ±0.20

7 5.02 5.03 20.00 balance

±0.10 ±0.10 ±0.40 77

Table 4-3

Standard Gas Composition Used for GC Calibration

Compound Detector Calibration Mixture

CO FID 2, 4, 5, 6

TCD 1, 6, 7

n o FID 4, 6

to 5,

TCD 1, 5, 6, 7

h 2 TCD 2, 3, 4, 5, 6,

n 2 TCD 2, 4, 5, 6, 7 78

Table 4-4

Quadratic Constants for GC Calibration

yt = a ^ + b A 2

Yi = mole percent of i

ait bA = calibration constants

Ai = chromatograph peak area

Ri2 = correlation coefficient

Component Detector bi Ri2 CO FID 1.551x10*’ 1.658X10*1® 1.0071

TCD 1.19x10*“ -8.009X10*11 1.0264

C02 FID 1.010x10*® -1.225X10*13 0.9865

TCD 1.356x10*“ 3.371X10*11 1.0012

h 2 TCD 7.617x10*® 3 .476xl0*13 1.0075

n 2 TCD 8.390x10*® 1.008X10*11 0.9997 79

levels. Figure 4-5 shows the final calibration curves used

for concentration calculations.

Three instances of calibration overlap occurred. That

is, C02 concentrations of 0.5 mol percent and 0.999 mol

percent, and a CO concentration of 0.999 mol percent were used

to calibrate both detectors. The calibrated C02

concentrations were within 2.5% of the actual concentration in both cases using both detectors. However, there was

considerable difference between the CO mol percent calculated

from the TCD calibration and the actual CO concentration of

the calibration gas. The CO concentration calculated using

the FID calibration was quite close to the actual CO

concentration. This illustrates the importance of the more

sensitive FID at low concentrations.

4.3 Reactor Dead Volume Evaluation -- Tracer - Response Tests

A series of tests in which hydrogen was introduced as a

step function into nitrogen flowing through the reactor was carried out to measure the lag time between the introduction of reactive gases and their appearance in the reactor product gas. The response curve corresponding to the introduction of

10% H2 at a total flow rate of 100 seem at 550°C and 15 atmospheres is shown in Figure 4-6. Hydrogen flow was initiated at t„ = 4.0 minutes, and no hydrogen was detected in the first three product gas samples taken after 9, 18, and 27 minutes. After 36 minutes the hydrogen mol fraction was Area Count (x 10000) Area Count (x 10000) 700 600 500 100 400 200 300 7 9 0 8 5 6 1 3 4 2 0 0 0 7 CO (TCD) CO 7 . 04 0.8 0.4 0.0 O(I) 0 z (FID) CO : r r - - / r Mole Percent Mole lili Mole Percent Mole 4 / OS 1—i— I—i—1—i— / 8 / 12 1.2 Figure4-5. GC CalibrationCurves. o o o o o < to 0) o 500 ^ 300 5 o' o o o 4-* < 200 i S S C 1 6 2 4 0 400 100 0 0 . 04 0.8 0.4 0.0 - ■ C02 (FID) (FID) C02 ■ ■

Mole Percent Mole !_.i 4 Mole Percent Mole i— 1 / Q J I — i— 1 — i— 1 — i— 8 P 12 1.2 o 200 g 150 x 2 2 100 250 5 0 5 0 25 20 15 10 5 0 Mole Percent Mole T = 550°C F = 100 seem SP (H2) = 10%

0 10 20 30 40 50 60 70 Time, minute

Figure 4-6. Hydrogen Response Curve. 82 approximately 0.04 and the concentration continued to increase as shown in the figure. A lag time of 40 minutes was calculated using the estimated reactor volume between the feed and chromatograph sampling valves, the gas volumetric flow rate, and assuming plug flow. The fact that hydrogen was observed prior to 40 minutes along with the gradual increase in hydrogen concentration indicates that axial mixing is important and that plug flow is not a good assumption under the conditions considered.

After examining this and other response curves under different conditions, we concluded that the time required for the hydrogen concentration to reach 80% of its steady-state value was the proper definition of the reactor lag time. In the Figure 4-6 example, the resulting lag time is 52 minutes.

Figure 4-7 shows a plot of lag time versus volumetric flow from a series of tracer - response tests at 550°C and 15 atm.

As shown in the figure, the lag time decreased to 20 minutes at 250 seem, which was the standard total flow rate for most of the later runs.

4.4 Preliminary Reaction Tests

The preliminary tests examined calcination/carbonation of

CaO-based sorbents at reaction conditions similar to those used in the previous TGA study by Silaban (1993). One test showing the feasibility of the simultaneous shift reaction and

C02 separation reaction is also discussed. Time Delay, minute 60 40 50 30 20 10 0 ■- 0 100

T = 550°C P= 15atm 120 Figure 4-7. Reactor Time Delay as a Function of Function a as Delay Time Reactor 4-7. Figure 140

Volumetric Flow Rate. Flow Volumetric Flow Rate, seem 160 180 220200 240

OJ 00 84

4.4.1 Calcination and Carbonation Reactions

Five tests consisting of one complete calcination and carbonation cycle were conducted during the preliminary reaction tests. Test conditions are summarized in Table 4-5.

Effect of Temperature

Figure 4-8 compares the C02 content of the calcination product gas from dolomite as a function of time for set-point temperatures of 800°C and 750°C. Calcination started at about

100 minutes when the sorbent bed temperature was around 425°C.

The initial peak for both curves A and B occurred at about the same time at a temperature of about 650°C. Calcination was completed earlier in the 800°C test, which was expected because higher temperature corresponds to higher calcination rate.

The effect of the carbonation temperature is more interesting and Figure 4-9 shows the C02 concentration versus time during carbonation. t = 0 corresponded to the time when the carbonation gas composition (15% C02/N2) was switched to the reactor. Curve A in Figure 4-9 shows that the C02 concentration in the product quickly increased to about 15% at

3.2 atmospheres, indicating that little or no carbonation was occurring. At t = 52 minutes, the system pressure was increased and the final pressure of 15 atm was reached after

69 minutes. Soon thereafter, the concentration of C02 decreased to about 5% and remained at that level until the run 85

Table 4-5

Summary of Reaction Conditions for the Preliminary Tests

Test 02 03 04 05 06

Sorbent D* D* D* D* MC’ Initial Mass, g 18.5 20.4 19.1 19.8 19.9 Particle Size 149- 149- 149- 210- 149- Range, (xm 210 210 210 250 210

Calcination Phase Temp., °C 800 750 750 750 750 Pressure, atm 3.2 3.2 3.4 3.2 3.2 Gas Comp., %N2 100 100 100 100 100 Flow Rate, seem 250 250 500 500 500

Carbornation Phase Temp., °C 800 750 650 550 550

Pressure, atm 3.2/15 15 15 15 15

Feed Gas Comp., %C0 0 0 0 0 0 %C02 15 15 15 15 15 %h 2 0 0 0 0 0 %h 2o 0 0 0 0 0 %n 2 85 85 85 85 85

Flow Rate, seem 250 150 100 100 100

Space Velocity (STP), hr*1 1138 6 83 455 455 455

D* - dolomite

MC* - marble chips C02 Mole Percent (dry basis) 40 20 30 10 o

0 innnnaopBBHa Figure 4-8: Effect of Calcination Temperature. Calcination of Effect 4-8: Figure 100 Time, minute 200 O OB D OOBBH OOO O O Q O O Test 03 (750°C)03 Test (800°C)02 Test 300 as CO C02 Mole Percent (dry basis) q A (3.2 atm) (3.2 A - Figure4-9. E£fectof Carbonation Temperature. o □ □□□□ - q Q-Q 50 o

&OO&Q & O O O& 0 150 100 Time, minute Time, e-e-e-0 1 atm) (15 A

200 T est 03 (750°C) (750°C) 03 est T (800°C) 02 est T T est 05 (550°C) 05 est T (650°C) 04 est T A “ A “ A A A 250 i was stopped at t = 165 minutes. Curves B through D in Figure

4-9 show the reactor response at a carbonation pressure of 15

atmospheres and temperatures of 550°C, 650°C, and 750°C. It is

interesting to note that C02 breakthrough occurred in all

three cases and that the prebreakthrough concentration of C02 decreased as the carbonation temperature decreased, which

indicated that the degree of carbonation was thermodynamically limited. Figure 4-10 compares the experimental C02 partial pressure during the prebreakthrough period and the equilibrium

C02 partial pressure.

Different Sorbent Precursors

Figures 4-11 and 4-12 compare calcination and carbonation results from tests 05 and 06. Test 05 used dolomite while test 06 used marble chips. The detailed composition of marble chips is not available but it consists mainly of calcium carbonate. In Figure 4-11, we can clearly see that the first peak in Curve A corresponds to calcination of MgC03 in dolomite. The near zero concentration of C02 in the prebreakthrough period for both tests (Figure 4-12) shows again that nearly complete removal of C02 is feasible at 550°C and 15 atm. Breakthrough occurred at an earlier time using dolomite because MgO was not recarbonated at the reaction conditions. C02 Partial Pressure, atm 0.9 0.8 0.6 0.7 0.4 0.5 0.3 0.2 0.1 0.0 .0 5 60 5 70 5 800 750 700 650 600 550 Figure 4-10. Prebreakthrough Partial Pressure of Pressure Partial Prebreakthrough 4-10. Figure C02 from the Fixed-Bed Reactor. C02Fixed-Bed the from Temperature, °C Temperature, Equilibrium Experimental

C02 Mole Percent (dry basis) innnnnnnnna Figure4-11. Calcination ofSorbents.Different 100 Time, minute Time, 200 B — Test 06 (Marble Chips) (Marble 06 Test (Dolomite) — 05 B Test — « 300 VO o C02 Mole Percent (dry basis) 12 14 16 18 20 10 8 2 6 4 m

150 200 250 92

Carbon Monoxide Formation

During the preliminary reaction tests, small concentrations of carbon monoxide were detected in the calcination product gas at high temperature. The reaction

C(s) + C02(gr) - 2 CO(g) (4-1) with the reactor wall serving as the carbon source was the suspected cause of the CO. A blank run was made in which a

15%C02 -N2 mixture was fed to the empty reactor. Only C02 and

N2 were detected in the product gas until the reactor temperature reached about 450°C, when trace quantities of CO began to appear. The amount of CO increased to a maximum value of about 0.5 mol percent as the reactor temperature increased to the 750°C set point. C02 feed to the reactor was stopped and the CO concentration returned to zero soon thereafter. Results of this blank test confirmed that CO originated from the above reaction, rather than being produced during the calcination process.

The participation of the reactor wall in the reaction should not be a significant factor in this research due to the following reasons. First, the extent of reaction 4-1 was very small and tended to occur only at high temperature (~750°C) .

Secondly, the combined shift and carbonation stage was of major interest in this research, and the shift - carbonation reaction was usually carried out at a temperature much lower than 750°C (discussed in the next chapter). As a result, the reactor wall was considered inert through this research. 93

4.4.2 Simultaneous Shift and Carbonation Reactions

Close to 100% removal of C02 at 550°C and 15 atm from the preliminary carbonation tests encouraged immediate study of

the combined shift and carbonation reactions. A series of

tests with the feed gas containing either C0/H20/N2 or

C0/C02/H2/H20/N2 were conducted. Results from one test are presented at this point to show the feasibility of the simultaneous shift reaction and C02 separation.

Figure 4-13 shows the reactor response from test 15 in which the reactor feed contained only CO, H20 and N2. The dolomite had been previously calcined before the feed gas was directed to the reactor. The reaction conditions are also shown in the Figure 4-13. All components required for the water gas shift reaction were present in the reactor feed.

Since no C02 or H2 was fed, the shift reaction must occur before either H2 or C02 can appear in the reactor product.

CO, H20, and N2 were introduced into the reactor at t =

368 minutes. H2 began to appear in the reactor product about

30 minutes later, and the H2 concentration increased to a constant level of about 14.5% (dry basis) in the time period

470 to 530 minutes. Small concentrations of CO and C02 were found immediately after the valve switch. They quickly decreased and were below the detection limits of the TCD from

430 to 530 minutes. Breakthrough of CO and C02 began at 530 minutes and the H2 concentration began to decrease at the same time. After approximately 620 minutes, a new.steady-state was Mole Percent (dry basis) 16 0 50 0 70 0 900 800 700 600 500 400 Figure 4-13. Product Gas Concentration During Concentration Gas Product 4-13. Figure Carbonation. Time, minute Time, SV = 570 hr-1 570 = SV T = 550°C = T P = 15 atm 15 = P

95 reached in which the reactor product contained about 9.4% H2 and C02 and 4.1% CO (all dry basis). These concentrations were maintained until the run was terminated after 790 minutes.

The time required for the H2 concentration to increase to r the 14.5% level is attributed to the time required for the reactor feed gas to displace the inert gas in the system volume between the feed gas valve and the product sampling valve. When only H2 and N2 were detected by the TCD, both the shift and carbonation reactions proceeded essentially to completion. During this period the more sensitive FID indicated approximately 200 ppm of CO and 400 ppm of C02 in the product gas. The H2 concentration in the product gas corresponding to complete shift and carbonation reactions would be 14.8%, quite close to the 14.5% experimental value.

Breakthrough of CO and C02 corresponded to the leading edge of the carbonation reaction front just reaching the end of the sorbent bed, while the second steady-state beginning at about

620 minutes corresponded to essentially complete conversion of

CaO to CaC03. Only the gas phase shift reaction was possible after this time. The fact that the final C02 and H2 concentrations were equal was consistent with the water gas shift stoichiometry with no C02 or H2 in the feed.

Test 15 provided definite proof that the shift and carbonation reactions would occur simultaneously at reaction conditions of interest. During the prebreakthrough phase on the test, approximately 99.5% of the total carbon oxides

(C0+C02) were removed from the gas phase. More importantly, this test represented the fifth complete calcination- carbonation cycle of the dolomite sorbent, showing that sorbent durability achieved in the electrobalance reactor tests may also be achieved in the fixed-bed reactor. CHAPTER 5

EXPERIMENTAL RESULTS AND DISCUSSION: SINGLE CYCLE STUDIES

This chapter focuses on results of single cycle tests which include calcination and carbonation/shift reaction stages. The first section discusses the reaction parameters and the factors in selecting reaction conditions. The second section defines variables used for the evaluation of the reactor behavior. Then experimental data reproducibility is described. Section 5.4 details the reaction parameter studies and presents results under different reaction conditions.

Finally, conclusions based on single cycle tests are summarized in Section 5.5.

5.1 Reaction Parameters

Two kinds of reaction parameters were considered in this research. The first group involved reaction conditions and included temperature, pressure, flow rate (or space velocity) , and feed gas composition. Because each single cycle test consisted of calcination and shift-carbonation phases, these four parameters were considered separately in each phase. The second group of parameters included different sorbent precursors and sorbent particle size. Detailed study of all the ten (10) parameters would include hundreds of tests even if only two cases were considered for each parameter (210 =

97 98

1024). This is impossible and unnecessary. The following factors were used in selecting reaction parameters:

1. Silaban's electrobalance results;

2. Characteristics of the fixed-bed reactor system;

3. Kinetics of the water-gas shift reaction;

4. Results of the preliminary reaction tests.

5.1.1 Silaban's Results

As discussed in the Chapter 2, Silaban (1993) conducted kinetic studies on the calcination and carbonation of different CaO-based sorbents using a TGA reactor. After screening nine sorbent precursors, he focused on three - reagent grade calcium carbonate, commercial dolomite, and reagent grade calcium acetate -for further reaction analysis.

Table 5-1 lists major conclusions reached by his research.

From Table 5-1, we see that the optimum calcination conditions were 750°C using an inert gas. Pressure was not important during calcination. For carbonation, the optimum kinetics were achieved over a temperature range of 650°C -

750°C at 15 atm pressure, using any desulfurized coal gas. Of the three sorbents, calcium carbonate was found to posses the lowest reactivity and capacity maintenance, while dolomite exhibited the best performance in terms of durability.

These results provided a guide for the current research.

However, there were two major differences between Silaban's work and this research. First, the TGA study used about 12 mg 99

Table 5-1

Summary of The Results from TGA Studies (Silaban, 1993)

X. Optimum Reaction Conditions from TGA Study

Calcination temperature: 750°C

Calcination pressure: 1-15 atm

Calcination atmosphere: any inert gas with low C02

partial pressure

Carbonation temperature: 650 - 750°C

Carbonation pressure: 15 atm

Carbonation atmosphere: any sulfur-free or low-sulfur

coal gas

XX. Sorbent Performance Comparison

Durability (reactivity and capacity maintenance):

Dolomite > Calcium Acetate > Calcium Carbonate 100 of sorbent, while in the fixed-bed study the sorbent amount was increased to about 13 g. Both inter- and intra-particle mass transfer had to be considered in this research.

Secondly, the current research dealt with combined shift and carbonation reactions while only the carbonation reaction was investigated in the previous study. Hence additional factors specific to this research had to be considered in determining the reaction parameters.

5.1.2 Characteristics of the Fixed-bed Reactor

The fixed-bed reactor system discussed in the Chapter 3 imposed certain limitations in selecting reaction parameters.

The furnace and the pressure control system were similar to those used in the previous TGA studies. 3.3 atm reactor system pressure was the minimum which could be achieved when the back pressure regulator was used. The mass flow controllers, which were factory calibrated, had a certain capacity with respect to a specified gas. The same was true for the syringe pump used to feed water.

Another limitation was associated with the porous stainless steel disks used to separate the sorbent sections in the reactor. These disks had to prevent sorbent from entering the exit lines and also cause a small pressure drop. Disks capable of retaining particles with a size above 50 /an were chosen. 101

Among the three sorbents used in Silaban's research, reagent calcium carbonate and calcium acetate had an intrinsic particle size of -400 mesh (< 38 /xm) , which is below the limit of the porous disks. Hence these two sorbents were excluded from the experimental tests. Other CaO-based sorbents having larger particle size were chosen for comparison purposes.

5.1.3 The Water-Gas Shift Reaction

The water-gas shift reaction was another important factor in determining the parameters in the second reaction stage.

High temperature increases the rate of the reaction, but decreases the equilibrium conversion of CO. Although a temperature range from 650°C to 750°C was reported to be the optimum carbonation condition from the TGA study, lower temperatures were expected to favor the combined shift and carbonation reactions. The actual optimal temperature was established through the reaction tests.

5.1.4 Preliminary Reaction Tests and Final Parameter Selection

The preliminary reaction tests described in Chapter 4 showed the reactor response under different conditions.

Although the tests didn't cover all possible conditions, they guided the choice of reaction parameters.

As shown in the Figure 4-9, total carbon oxide removal approached 100% at 550°C and decreased dramatically when the carbonation temperature increased to 750°C. Figure 4-13 102

showed similar results for the combined shift-carbonation

reaction.

Based upon the above factors, the final reaction

parameters listed in Table 5-2 were selected. The standard

conditions were used for most of the runs and the alternative

conditions were tested for selected parameter combinations.

Since the calcination stage was , less important than the

combined reaction stage, calcination parameters were not

varied over as wide a range as carbonation parameters.

For calcination, pressure and the inert gas flow rate

were not important except that high flow rate tended to sweep

the produced C02 out of the reactor faster and thereby shorten

the lengthy calcination period. For most tests, a nitrogen

flow rate of 700 seem was used during calcination.

Calcination using a mixture of steam and N2 was studied to

show that high concentrations of C02 would be produced after

steam condensation. The ability to produce large CO,

concentrations may become more important if future regulations restrict atmospheric C02 emissions.

The most important parameters for the single cycle tests were the calcination temperature and gas composition, and the temperature, pressure, space velocity, and gas composition for the shift/carbdnation reactions. The H20 to CO ratio was used as a secondary parameter, because of its importance in determining the extent of the shift reaction. Another factor which had to be considered when choosing feed gas composition Table 5-2

Reaction Parameters

Reaction Stage Parameters Standard Conditions Alternative Conditions

Calcination: Temp., °C 750 650, 800, 850, 900

Pressure, atm 3.3 1

SV\ hr'1 (STP) 3990 1425

Gas Comp*. 100%N2 20%, 40%, and 60%H2O/N2

Shift - Temp., °C 550 400, 500, 600, 650, 750

Carbonation: Pressure, atm 15 5, 1 SV*, hr'1 (STP) 1425 885, 2280, 3420

Gas Comp**. 5.6%CO-20%H2O-N2 simulated coal gas with

varying H2, CO, C02, H20,

and N2 concentrations

*: Space velocity

**: Mole percent 103 104 was the possibility of carbon deposition. Research conducted by Lamoreaux et al. (1986) was used to guide the selection of feed gas compositions which were free of carbon formation at reaction conditions of interest.

5.2 Shift/Carbonation Reaction Variables

The performance of the sorbent under different reaction conditions was evaluated using a number of measures including:

i) prebreakthrough CO and C02 concentrations;

ii) postbreakthrough CO and C02 concentrations;

iii) fractional COx removal during the prebreakthrough

period;

iv) final fractional sorbent conversion;

v) dimensionless breakthrough curves;

vi) temperature deviation from the set-point;

vii) sorbent structural changes.

Pre/post breakthrough CO and C02 concentrations and sorbent structural data were direct experimental results while the other variables involved manipulation of the primary data.

Items i) through vi) were evaluated for each run, but only a limited number of sorbents were chosen for structural tests.

With the value of the dead volume lag time known at various reaction conditions (Chapter 4), the following dimensionless parameters were defined: 105

= f °

cox• = (f ° * * „ (5 - 2 ) (y co,+ y co)F

FCOX = 1 - COX* ' (5-3)

** = - 7 ^ s ^°(y°co2 +y°co) - y(ya?2 + yco>]Ati (5-4) ■N caO

Where:

F° = molar rate of total reactor feed

(wet basis)

F = molar rate of total reactor product

(wet basis)

y°co2» y°co = mol fraction of C02 and CO in feed gas

(wet basis)

yco2» yco = mol fraction of C02 and CO in reactor product (wet basis) as a function of time

N°Ca0 = initial mols of CaO in the reactor

t = time from initiation of the carbonation

cycle

tD = dead volume lag time

At* = time increment between chromatograph

samples

t* = dimensionless carbonation cycle time

COX* s dimensionless breakthrough concentration

of total carbon oxides 106

FCOX = fractional removal of total carbon oxides

X* = axial average fraction of CaO converted to

CaC03

The characteristics of these equations are such that X*

= t* as long as complete removal of carbon oxides is accomplished, i.e., as long as COX* = 0 (or FCOX = 1). Once carbon oxide breakthrough begins, X* < t*. If both the shift and carbonation reactions were infinitely fast and complete sorbent conversion could be achieved, the reaction would be complete at X* = t* = 1.0, and COX* would change in a step function with COX* = 0 for t* < 1.0 and COX* = 1.0 for t’ a

1.0. The idealized dimensionless response curve is shown in

Figure 5-1(a) and compared to a typical response curve in

Figure 5-1(b).

In Figure 5-1(b), the value of COX* should approach the thermodynamic equilibrium limit for small t*, which for practical purposes is zero for the shift and carbonation reactions at the conditions of interest. As t* increases, the value of COX* will increase because of the finite rate of the global reactions and the traditional S-shaped breakthrough curve will result. For large t*, COX* will approach 1.0 as the total carbon oxide concentration in the product approaches the total carbon oxide concentration in the feed. Note that COX*

= 1.0 does not imply that both reaction rates are zero. The shift reaction may still occur but if the carbonation reaction rate is zero, no carbon dioxide is removed from the gas phase FCOX or X 0 / / / / / / / ) daie Ratr Response Reactor Idealized a) / / i / / / / X/ Figure 5-1. Dimensionless Breakthrough and Calcium Conversion Calcium and Breakthrough Dimensionless 5-1. Figure / / / * iesols Tm, t Time, Dimenslonless / / / /

/ / / / / / 2 1 FCOX Response Curves for a Fixed-Bed Reactor. Fixed-Bed a for Curves Response *

* x § 0.50- 0.50- § £ o L. 0.00 1 . 00 - ) cul eco Response Reactor Actual b) iesols Tm, t Time, Dimensionless FCOX

* 108

and the total carbon oxide concentration is constant. After

carbon oxide breakthrough begins, the slope of the X* versus t* curve will decrease and the X* will asymptotically approach its final value.

In order to calculate t*, FCOX (or COX*) , and X*, the feed gas composition must be known. Multiple, and sometimes contradictory, methods of calculating feed gas composition are available. All gases were fed from high pressure cylinders through calibrated mass flow controllers. Water was added as a liquid using a high-pressure positive displacement syringe pump. Knowing the flow rates of the individual components, the feed gas composition can be directly calculated.

Alternatively, the composition of the feed gas can be determined by assuming that correct flow rates are obtained from the N2 flow meter and the H20 syringe pump, and using the chromatographic analysis of the feed to determine the CO, C02 and H2 concentrations. Table 5-3 compares the feed gas composition for run 18, in which all five feed components are present, using the two calculation methods.

All feed gas compositions reported are based upon the second method. That is, the chromatographic analysis for CO,

C02 and H2 coupled with measured flow rates for N2 and H,0.

Since the CO, C02, and H2 concentrations in the reactor product must be measured by chromatography, the selected approach means that the calculation basis for the feed and product streams is consistent. Table 5-3

Feed Gas Composition for Test 18 Using Two Calculation Methods

Using All Flow Meters Using N2 Flow Meter, and Syringe Pump Syringe Pump and Chromatograph

Component Mol Fraction SCCM Mol Fraction SCCM

CO 0.06 15.0 0.064 16.0

C02 0.06 15.0 0.056 13.8 h 2 0.05 12.5 0.047 11.7 h 2o 0.20 50.0 0.200 50.0 n 2 0.63 157.5 0.633 157.0 110

5.3 Experimental Result Reproducibility

As shown in Chapter 3, there were many control devices used in this research and each device had a range of accuracy.

Hence it is expected that a certain degree of error exists in the final results. Main errors came from the flow meters, GC integration methods, and those associated with loading and packing the sorbent into the reactor. Duplicate runs were conducted at different stages of the research.

Figure 5-2 shows the dimensionless breakthrough curves for CO, C02, and H2 (TCD results) for duplicate tests 38 and

44. Standard reaction conditions were used except the carbonation gas flow rate was 100 seem (SV = 885 hr'1) . Both the component concentrations and the slopes of the breakthrough curves for these two tests were very close. The prebreakthrough concentrations of CO and C02 were below the

TCD detection limit. The more sensitive FID showed that the prebreakthrough concentrations were 64 ppm for CO and 264 ppm for C02 for test 38, and 48 ppm for CO and 254 ppm for C02 for test 44, respectively. The comparison of the sorbent fractional conversion and the fractional COx removal are presented in Figure 5-3. The prebreakthrough COx removal of test 38 (99.68%) was very close to that of test 44 (99.62%), and the final sorbent conversions were 0.914 and 0.907 for tests 38 and 44, respectively. Figures 5-4 and 5-5 show the temperature responses for these two runs. The temperature increased at t* < 0 because feed gas reached the sorbent I ! i

Concentration, mole percent 7 8 4 5 1 6 3 2 0 Figure5-2. Breakthrough Curvesfor Duplicate Tests. 0.20.0 ■ Test 38 ■ Test * Test 44 Test * 0.4 Dimensionless Time, t* Time, Dimensionless 0.6

0.8 CO 2 1.0 CO 1.2

1.4 Ill Fractional COx Removal . 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 Figure 5-3. Fractional COx Removal and Sorbent Conversion Sorbent COxand Removal Fractional 5-3. Figure for Duplicate Tests. for Duplicate Dimensionless Time, t* Time, Dimensionless

1.0 1.1 0.6 0.7 0.8 0.9 0.5 0.3 0.1 0.2 0.4 0.0

Fractional Sorbent Conversion, X to t-* Temperature Deviation from Set Point, °C -10 - 0.1 Figure 5-4. Temperature Deviation Within the Within Deviation Temperature 5-4. Figure 0.1 0.3 Dimensionless Time, t* Time, Dimensionless Sorbent Bed for Test 38. Test for Bed Sorbent . 0.7 0.5 . 1.1 0.9 Test 38 Test

. 1.5 1.3

113 Temperature Deviation from Set Point, °C -10 T2 T - - 0.1 Figure 5-5. Temperature Deviation Within the Within Deviation Temperature 5-5. Figure 0.3 Dimensionless Time, t* Time, Dimensionless Sorbent Bed for Test 44. Test for Bed Sorbent . 07 . 1.1 0.9 0.7 0.5 Test 44 Test . 1.5 1.3

114 115 before it reached the chromatograph. For both tests, the temperature in the top section of the packed bed was approximately 10°C higher than the bottom section temperature, and the combined shift and carbonation reactions produced a temperature increase of about 7°C in both cases as the reaction front passed each thermocouple position.

Temperatures at different positions inside the reactor show approximately the same tendency for the two tests.

Duplicate runs (test 61 and test 63-1) were also conducted for another sorbent - marble chips. In Figure 5-6,

CO and C02 concentrations are plotted versus dimensionless time. The log scale permits the FID results to be included.

The prebreakthrough concentrations for these two tests were almost identical.

Table 5-4 lists values of some key variables for another set of duplicate tests using dolomite. In terms of prebreakthrough C0X removal, the values ranged from 99.61% to

99.71%. Again, the data show good reproducibility.

5.4 Detailed Parameter Tests

About 60 successful runs were conducted. Thirty-seven of these were single cycle tests. These tests were divided into different groups according to the reaction conditions. Each group studied the effect of a single parameter while all the other parameters were kept constant. The first cycle results of the multicycle tests are included in the discussion of the Concentration, ppm 101 10° 0.2 iue56 Lgpo fBekhog uvs for Curves Breakthrough of Log-plot 5-6. Figure 2 0 C 0.4 0.6 Duplicate Tests. Duplicate Dimensionless Time, t* Time, Dimensionless 0.8 - □ □ □ «-B □ □ o-e-o -o o o # . # 1.0 - Test 63-1 Test - * Test 61 Test * 1.2

1.4

9X1 117 Table 5-4

Comparison of Duplicate Test Results under Standard Conditions

Test # Prebreakthrough Cone. Prebreakthrough C0X

CO (ppm) C02 (ppm) removal, FCOX

27-1 9 190 99.71%

51-1 21 253 99.61%

65-1 29 211 99.66%

Test conditions:

Calcination Carb./Shift

Temperature: 750°C 550°C

Pressure: 3.2atm 15 atm

Gas comp.: 100%N2 5.6%CO-20%H2O-N2

Flow rate: 700 seem 250 seem

Sorbent: Dolomite

Particle Size: 149 fim - 210 jxm

Initial weight: 13.2 g 118 detailed parameter studies. Main effort was focused on calcination and combined shift and carbonation reactions, and results showing the nature of the water-gas shift reaction are also discussed.

5.4.1 Comparison of Different Sorbent Precursors

Two limestone sorbent precursors were tested to compare the performance with dolomite. The conditions for these tests are shown in Table 5-5.

The limestone sorbent precursors, which are primarily

CaC03, have a larger C02 capacity than does dolomite because of the presence of inert MgC03 in the latter. The exact composition of neither limestone was available. Calcination results suggested that the limestone used in test 57 contained significant concentrations of impurities while test 61 calcination results suggested that the CaC03 content of the marble chips was close to 100%. These conclusions were based on the measured amount of C02 liberated during calcination.

Standard shift/carbonation reaction conditions were used in each test and results are compared with dolomite results in

Figure 5-7, where the fractional removal of total carbon oxides (based on TCD analysis) is plotted as a function of time. Effectively complete carbon oxide removal was achieved during the early stages of the reaction in all three tests.

Initial breakthrough with limestone (test 57) began after about 50 minutes compared to 100 minutes with dolomite and 130 119

Table 5-5

Summary of Reaction Test Conditions for the Different Sorbent Precursors

Test 27-1 57 61

Sorbent D L* MC* Initial Mass, g 13 .2 13.2 13 .2 Particle Size 149- 149- 149- Range, /urn 210 210 210

Calcination Phase

Temp., °C 750 750 750 Pressure, atm 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 Flow Rate, seem 700 700 700

Carbornation Phase

Temp., °C 550 550 550 Pressure, atm 15 15 15 Feed Gas Comp., %C0 5.6 5.6 5.6 %C02 0 0 0 %h 2 0 0 0 %h 2o 20.0 20.0 20.0 %n 2 74.4 74.4 74.4

Flow Rate, seem 250 250 250

Space Velocity (STP) , hr’1 1425 1425 1425

L: Limestone MC: Marble Chips Fractional Removal of Carbon Oxides, FCOX 1.0 0.8 0.9 0.6 0.7 0.4 0.5 0.3 0.1 0.2 0.0 .1 iue57 Fatoa eoa fCro xds for Oxides Carbon of Removal Fractional 5-7. Figure Limestone Test 57 Test ffi i /rw H 50

Three Test Sorbents at Standard Conditions. Standard at Sorbents Test Three Reaction Time, minute Time, Reaction 100 Dolomite et5 ] 51 Test 150

Test 61 Test Marble Chips Marble

200

250 120 121 minutes with marble chips. The slope of the limestone breakthrough curve was also much lower, and reaction was not complete until 225 minutes of total elapsed time. The slopes of the breakthrough curves using dolomite and marble chips were similar, with only about 40 minutes from initial breakthrough to complete reaction. The longer time to initial breakthrough using marble chips is a reflection of the increased capacity of high calcium marble chips compared to dolomite. When the marble chips and dolomite results are compared on a dimensionless time basis (which eliminates the capacity differences) as in Figure 5-8, we see that the calcium in dolomite is more accessible to the C02. Initial breakthrough did not occur until approximately 80% of the calcium in dolomite had reacted, while initial breakthrough with marble chips occurred at about 68% calcium conversion.

The detailed reaction studies focused on dolomite because of its superior performance.

5.4.2 Effect of Calcination Temperature

Table 5-6 summarizes the reaction conditions of the tests which studied the effect of calcination temperature. Two subgroups are included in the table. Subgroup A corresponds to the simple shift-carbonation feed gas composition, while

Subgroup B corresponds to a simulated coal gas. Calcination at 650°C (test 28) was not acceptable as the reaction was still incomplete six hours after the cycle was initiated. As Fractional Removal of Carbon Oxides, FCOX 1.0 1.1 0.7 0.8 0.9 0.5 0.6 0.4 0.3 0.1 0.2 0.0 0.0 xj gn i ^ B P iiD a c n p n -g m o jp ix Figure 5-8. Comparison of the First Cycle Reactivity Cycle First the of Comparison 5-8. Figure 0.2 Dimensionless Time Basis. Time Dimensionless a on Chips Marble and Dolomite of 0.4 Marble Chips Marble Dimensionless Time, t* Time, Dimensionless Test 61 Test 0.6

0.8 Test 51 Test Dolomite 1.0

1.2

1.4

122 Table 5-6

Summary of Reaction Test Conditions for the Effect of Calcination Temperature

Subgroup A Subgroup B

Test 27-1, 28, 29 66-1, 67-1, 68-1 Sorbent D D DD DD Initial Mass, g 13.2 13.2 13.2 13.2 13.2 13.2 Particle Size 149- 149- 149- 149- 149- 149- Range, fim 210 210 210 210 210 210

Calcination Phase Temp., °C 750 650-750 800 750 800 850 Pressure, atm 3.3 3.3 3.3 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 100 100 100 Flow Rate, seem 700 700 700 700 700 700

Carbomation Phase Temp., °C 550 550 550 550 550 550 Pressure, atm 15 15 15 15 15 15 Feed Gas Comp., %C0 5.6 5.6 5.6 6.6 6.6 6.6 %C02 0 0 0 3.1 3.1 3.1 %h 2 0 0 0 4.6 . 4.6 4.6 %h 2o 20.0 20.0 20.0 26.4 26.4 26.4 %n 2 74.4 74.4 74.4 59.3 59.3 59.3

Flow Rate, seem 250 250 250 250 250 250

Space Velocity (STP) , h r ' 1 1425 1425 1425 1425 1425 1425 123 124 indicated in Table 5-6, the calcination temperature was subsequently increased to 750°C to complete the reaction.

A calcination temperature of 800°C reduced the time required for calcination and increased the C02 composition of the product gas during the sorbent decomposition period.

Figure 5.9 compares the C02 product gas compositions and reactor temperatures for tests 27-1 and 29. MgC03 decomposition occurred during the heating period and was, as expected, approximately the same in both tests, as shown in the first peak of both calcination curves. The lower peak value for test 29 was due to the large sampling time interval.

The duration of the calcination cycle was reduced by about 30 minutes and the C02 mol fraction during CaC03 decomposition increased from about 0.05 to 0.09.

Results from the previous study (Silaban, 1993) indicated that improved sorbent durability was associated with lower calcination temperature. However, there was little, if any, difference in the first cycle carbonation performance of tests

27 and 29 as shown by the dimensionless response curves in

Figure 5-10. CO and C02 concentrations were below the TCD detection limit during the prebreakthrough period of both tests. Initial breakthrough occurred at approximately the same dimensionless time and the slope of the breakthrough curves was effectively equal.

Figure 5-11 shows the response of the C02 concentration during the calcination period for tests 66-1, 67-1, 68-1 using C02 Mole Percent (dry basis) 4 2 8 0 4 2 6 0 2 4 6 8 10 2 10 6 10 0 220 200 180 160 140 120 100 80 60 40 20 0 Figure 5-9. C02 Composition of the Calcination Product Calcination the of C025-9. Composition Figure ^rtl^n w k

Jrtr Gas as a Function of Time and Temperature. and Time of Function a as Gas Time, minute Time, Test 27-1 Test Test 29 Test Test 29 Test Test 27- Test

900 700 800 600 400 500 100 300 200 0

Temperature, °C in to h> (CO +CO 2) Breakthrough, COX. 1.0 1.1 0.7 0.8 0.9 0.5 0.6 0.4 0.2 0.3 0.0 0.1

. 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0

Figure 5-10. Dimensionless Carbonation Breakthrough Curves Breakthrough Carbonation Dimensionless 5-10. Figure et27- ,750°C -1 7 2 Test et29,800°C 9 2 Test Following Different Calcination Temperatures. Calcination Different Following Dimensionless Time, t* Time, Dimensionless

Fractional Sorbent Conversion, X NJ H> C02 Mole Percent (dry basis) 12 14 10 8 4 6 2 0 0 Figure 5-11. Effect of Temperature on COa on Concentration Temperature of Effect 5-11. Figure 0250 50 et6- (750°C) 66-1 Test et6- (850°C) 68-1 Test et6- (800°C) 67-1 Test During Calcination. During 100 Time, minute Time,

150 200

127 128

calcination temperatures of 750°C, 800°C, and 850°C. All three were multicycle tests and the first cycle results are used here for comparison. High calcination temperature reduced the

time required for complete calcination and increased the C02

concentration during the CaC03 decomposition period. Figure

5-12 shows results from the subsequent carbonation/shift cycle using simulated coal gas as the feed. No deterioration was observed in the fractional C0X removal and the sorbent conversion curves in the first cycle. In multicycle test 72, calcination temperature was increased to 900°C. Although this test was conducted under different conditions and can not be directly compared with the tests in Figure 5-11, the subsequent shift/carbonation showed no significant deterioration. The multicycle test results will be covered in the next chapter.

5.4.3 Effect of Calcination Flow Rate

Reducing the inert gas flow rate during calcination provides an alternate means of increasing the C02 composition of the calcination product gas. Table 5-7 lists the tests in which the calcination gas flow rate was varied. In test 30, the N2 flow rate was reduced from the standard 700 seem to 250 seem, and the C02 content as a function of time is given in

Figure 5-13. During the MgC03 decomposition period the maximum C02 content increased from approximately 13% to 24%; a similar increase from 5% to 7% was observed during CaC03 Fractional COx Removal 1.0 0.7 0.8 0.9 0.6 0.4 0.5 0.3 0.1 0.2 0.0 . 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 Figure 5-12. E£fect of Calcination Temperature on Temperature Calcination of E£fect 5-12. Figure et67- ,800°C -1 7 6 Test ,750°C -1 6 6 Test et68- ,850°C -1 8 6 Test Breakthrough Curves. Breakthrough Dimensionless Time, t* Time, Dimensionless

1.1 1.0 0.8 0.7 0.9 0.6 0.4 0.5 0.3 0.0 0.1 0.2

Fractional Sorbent Conversion, X VO NJ H* 130

Table 5-7

Summary of Reaction Test Conditions for the Effect of Calcination Gas Flow Rate

Test 27-1 30

Sorbent D D Initial Mass, g 13.2 13.2 Particle Size 149- 149- Range, jim 210 210

Calcination Phase

Temp., °C 750 750 Pressure, atm 3.3 . 3.3 Gas Comp., %N2 100 100 Flow Rate, seem 700 250

Carbornation Phase

Temp., °C 550 550 Pressure, atm 15 15 Feed Gas Comp., VO %C0 5.6 in %C02 0 0 %h 2 0 0 %h 2o 20.0 20.0 %n 2 74.4 74.4

Flow Rate, seem 250 250

Space Velocity (STP) , hr'1 1425 1425 C02 Mole Percent (dry basis) 30 30 Q 0 0 10 0 250 200 150 100 50 ------n —•«* Figure 5-13. Calcination Product Composition as a as Composition Product Calcination 5-13. Figure »— ------—t 1 ------Function of N2of Rate. Flow Function 1 ------1 Time, minute 1 ------* ------1 ------1 —<* i I n. ta Flow Rate Steam 1 ------CalcinationGas % (seem) (%) 700 0 250 0 1 ------r

300 131 132 decomposition. An increase in the total time required for calcination of about 50 minutes also accompanied the decreased flow rate.

Figure 5-14 compares the shift - carbonation cycle of these tests. Fractional C0X removal and fractional sorbent conversion for these two tests are almost identical.

Prebreakthrough concentrations of CO and C02 for test 30 were

23 ppm and 150 ppm, compared with 10 ppm and 190 ppm for test

27-1. The differences were within experimental error, and no clear effect of calcination gas flow rate on the subsequent sorbent shift - carbonation was observed.

5.4.4 Effect of HjO in the Calcination Sweep Gas

In a commercial process, it will be desirable to produce calcination off-gas having the highest possible concentration of C02. This feature will become increasingly important if future environmental regulations limit the atmospheric discharge of C02. C02 concentration can be increased by carrying out the calcination step at high temperature and using a small sweep gas flow rate as shown in sections 5.4.2 and 5.4.3. In addition , the inclusion of a condensible component such as steam in the sweep gas will result in higher

C02 concentrations in the permanent gases following steam condensation.

The effect of the addition of up to 60% steam to the calcination sweep gas was studied at calcination temperatures Fractional COx Removal 1.1 1.0 0.8 0.9 0.7 0.5 0.6 0.3 0.4 0.1 0.2 0.0

. 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0

Figure 5-14. Effect of Calcination Flow Rate on Rate Flow Calcination of Effect 5-14. Figure et30,250sccm 0 3 Test et27- ,700sccm -1 7 2 Test Breakthrough Curves. Breakthrough Dimensionless Time, t* Time, Dimensionless

Fractional Sorbent Conversion, X u> 134

of 750°C and 800°C; calcination reaction conditions are

summarized in Table 5-8. Results are compared in Figure 5-15.

The addition of steam was effective in increasing the C02

content (dry basis) as expected, and also in causing

decomposition to occur more rapidly. In both tests 33 and 45,

C02 formed during MgC03 decomposition peaked at about 125

minutes, compared to 150 minutes in dry N2. The C02 content

at the peak increased from 24% in dry nitrogen to

approximately 43% in the steam-nitrogen mixture. The time

required for complete calcination decreased from approximately

260 minutes in dry N2 to 230 minutes in 60% H20/N2 at 75.0°C and

to 210 min in 60% H20/N2 at 800°C. The maximum C02 content of

the product gas (dry basis) during CaC03 decomposition

increased from about 7% in N2 at 750°C to 19% in 60% H20/N2 at

800°C.

While the addition of steam produced the dual benefit of

increasing the C02 concentration and decreasing the

calcination time, it also caused a significant change in the pore size distribution of the calcined sorbent, and, when

coupled with high calcination temperature, appeared to have a marginally negative effect on the kinetics of the subsequent carbonation cycle. Figure 5-16 compares the pore size distribution curves for as-received dolomite, and for dolomite calcined at 750°C in dry N2 and in 40% H20/N2. An overall pore size range of 0.005 to 0.6/m is represented in the figure.

As-received dolomite is essentially nonporous so that the 135 Table 5-8

Summary of Reaction Test Conditions for the Effect of Calcination Gas Composition

Test 30 31 33-1 45-1

Sorbent D DD D Initial Mass, g 13.2 13 .2 13 .2 13 .2 Particle Size 149- 149- 149- 149- Range, nm 210 210 210 210

Calcination Phase Temp., °C 750 750 750 800 Pressure, atm 3.3 3.3 3.3 3.3 Gas Comp., %N2 100 80 40 40 %H20 0 20 60 60

Flow Rate, seem 250 250 250 250

Carbornation Phase Temp., °C 550 550 550 550 Pressure, atm 15 15 15 15 Feed Gas Comp., %C0 5.6 5.6 5.3 5.6 %C02 0 0 0 0 %h 2 0 0 0 0 %h 2o 20.0 20.0 20.1 20.0 %n 2 74.4 74.4 74.6 74.4

Flow Rate, seem 250 250 250 250

Space Velocity (STP) , hr'1 1425 1425 1425 1425 C02 Mole Percent (dry basis) 20 0 Figure5-15. C02 Content ofCalcination Product Gas asa Function ofTemperature and Steam Content. 100 Time, minute Time, .* Test 33-1,60%H20,750°C Test 45-1,60%H2O,Test Test 30,0%H20,750°C Test r — i --- 200 BOQ

o C

300 136 Log Diff. Vol (dV/dLogD), ml/g 1.4 2.0 0.8 0.4 0.6 0.0 0.2 .8 .6 .0 .2 6 5 Figure 5-16. Pore Size Distribution of Calcined Dolomite Calcined of Distribution Size Pore 5-16. Figure AsReceived Dolomite CalcinedDolomite (750°/3.3atm/40%H20-60%N2) CalcinedDolomite (750°C/3.3atm/N2) as a Function of Calcination Atmosphere. Calcination of Function a as 2 Diameter, micron Diameter, 4 62 6 5 4 3 10-1 4 5 4 3

137 138

entire pore volume following calcination was produced by driving off C02. Although the total pore volume for both calcined samples was 0.36 cm3/g» there was almost an order of magnitude increase in the diameter of the pores formed in the

H20/N2 atmosphere. The range of pore diameters was approximately 0.08 to 0.3 pm with a peak at about of 0.2 pm following calcination in dry H20/N2, compared to a range from

0.02 to 0.1 pm with a peak at 0.04 pm when calcination was carried out in N2.

When the final calcination temperature was 750°C, the addition of steam seemed to have little effect on the subsequent first carbonation/shift reaction cycle. This is shown in Figure 5-17 where dimensionless breakthrough curves are plotted for three runs in which 0% H20, 20% H20, and 60%

H20 was added to the calcination sweep gas. However, the addition of 60% steam to the calcination gas coupled with the higher calcination temperature of 800°C had a negative effect on carbonation/shift kinetics, as shown in Figure 5-18.

Initial breakthrough occurred earlier and the slope of the breakthrough curve was smaller than at the standard calcination conditions.

The above analysis confirms that increasing the calcination temperature (up to 850°C) , reducing the sweep gas flow rate, and adding steam are effective ways of increasing the C02 concentration in the product gas. However, each of the methods of increasing C02 concentration has undesirable (CO +C O 2 ) Breakthrough, COX* 1.0 0.8 0.9 0.6 0.7 0.4 0.5 0.3 0.1 0.2 0.0 .1 0.0 Figure 5-17. Dimensionless COx Breakthrough Curves COx Breakthrough Dimensionless 5-17. Figure

0.8 1.0 1.2

1.4 139 (CO +C O 2) Breakthrough, COX* 1.0 0.9 0.7 0.8 0.5 0.6 0.4 0.3 0.2 0.1 0.0 .1 Figure 5-18. Dimensionless Response Curves Following Curves Response Dimensionless 5-18. Figure

. 0.8 0.6 1.0 1.2

1.4 140 141

consequences as well. High calcination temperature will

require additional energy to preheat the calcination sweep gas

and may produce excessive sorbent sintering. A low sweep gas

flow rate will cause the calcination rate to decrease and the

size of the calcination reactor to increase. Addition of a

condensible component such as steam may impose energy

penalties in order to generate the steam. Optimum calcination

conditions will depend upon process economics, and the impact

of calcination conditions on the sorbent durability.

5.4.5 Effect of Carbonation/Shift Temperature

Preliminary test results showed that the

carbonation/shift reaction temperature played an important

role in determining the extent to which carbon monoxide was

converted and C02 was removed from the gas. Since both the

shift and carbonation reactions are exothermic, the

equilibrium conversion is a decreasing function of

temperature. However, the global reaction rate increases with

increasing temperature, thereby increasing the chance that

equilibrium conditions will be closely approached.

Table 5-9 shows test conditions involving

carbonation/shift temperature variation. Five tests were

included with shift - carbonation temperatures ranging from

400°C (test 37) to 650°C (test 56-1) . The feed gas contained

only CO, H20 and N2. The temperature effect during the prebreakthrough period is shown in Figure 5-19. Table 5-9

Summary of Reaction Test Conditions for the Effect of Carbonation Temperature

Test 27-1 35 36 37 56-1

Sorbent D D D D D Initial Mass, g 13.2 13.2 13.2 13.2 13.2 Particle Size 149- 149- 149- 149- 149- Range, /im 210 210 210 210 210

Calcination Phase Temp., °C 750 750 750 750 750 Pressure, atm 3.3 3.3 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 100 100 Flow Rate, seem 700 700 700 700 100

Carbornation Phase Temp., °C 550 500 600 400 650 Pressure, atm 15 15 15 15 15 Feed Gas Comp., %C0 5.6 5.6 5.6 5.6 5.6 %C02 0 0 0 0 0 %h 2 0 0 0 0 0 %h 2o 20.0 20.0 20.0 20.0 20.0 %n 2 74.4 74.4 74.4 74.4 74.4

Flow Rate, seem 250 250 250 250 250

Space Velocity (STP), hr'1 1425 1425 1425 1425 1425 142 Prebreakthrough Concentration (dry basis), ppm Figure 5-19. Prebreakthrough CO and C02 Concentrations as C02and CO Concentrations Prebreakthrough 5-19. Figure 400 Temperature. Reaction Carbonation/Shift o£ Function a O / CO 5 0 550 500 450 Temperature,°C qlbim 5.6% equlibrium xeietl 20.0% experimental FEED COMPOSITION FEED SV =* SV 1425hri(STP) 600 74.4% P -1 5 atm 5 -1 P

650 CO HsP N2

Fractional Removal of Carbon Oxides, F M 144

Prebreakthrough concentrations of CO and C02 and fractional

removal of total carbon oxides are plotted versus reaction

temperature. Equilibrium concentrations of both components are included for reference. Duplicate run results at 550,

600, and 650°C are also shown in the figure to provide a measure of reproducibility. Experimental concentrations are reasonably close to the equilibrium values except at the lowest test temperature of 400°C. Consequently, carbon oxide removal is effectively equilibrium limited at temperatures of

500°C or greater, and kinetics are important only at temperatures less than 500°C. Maximum fractional carbon oxide removal of 0.999 occurred at 500°C and the fractional removal exceeded 0.99 from 400°C to 550°C. Even at 650°C, the fractional carbon oxide removal exceeded 0.93.

Results during the postbreakthrough period when only the shift reaction occurred followed the same trends. At high temperature (a 600°C) the experimental concentrations closely approached equilibrium values while at lower temperatures (s

550°C) the experimental results were relatively far from equilibrium. This behavior is shown in Figure 5-20. At the lower reaction temperatures associated with commercial shift reaction a heterogeneous catalyst is required. At the higher temperatures used in this study, it is possible that the shift reaction occurs homogeneously in the gas phase. Alternately, it may be catalyzed by MgO, CaO, CaC03, or even the reactor walls. MgO has been reported to be a shift catalyst in the Postbreakthrough Concentration (dry basis), % 7 9 8 6 500 equlibrium . experimental Figure 5-20.Figure Postbreakthrough Concentrations ofCO 2 540 520 Carbonation Temperature. and COaasaFunction Shift/ of Temperature, °C Temperature, 6 50 600 580 560 CO 2 FEED COMPOSITION FEED V=12h1 (STP) 1425hr1 = SV 5atm 1 = P 00 H 20.0% 44 N 74.4% .% CO 5.6% 2 640 620

2 2 O

145 146

literature (Gluud, 1931). Since iron is a major component of

the commercial high temperature shift catalyst, it is reasonable to suspect that the stainless steel surfaces of the reactor may catalyze the reaction. The mechanism of the shift reaction will be discussed in more detail later.

The effect of the reaction temperature on the shape of the breakthrough curve is shown in Figure 5-21 where fractional removal of total carbon oxides is plotted versus dimensionless time, t*. Test 27 was carried out at 550°C while test 37 was at the lowest experimental temperature of 400°C.

The use of dimensionless time compensated for the effect of temperature on feed gas flow rate and made direct comparison between the two tests possible. Fractional carbon oxide removal approached 100% during the early stages of both tests.

Initial breakthrough, however, occurred at t* ~ 0.5 at 400°C compared to t* ~ 0.77 at 550°C. Similarly, the carbonation reaction was effectively complete at t* ~ 1.15 at 550°C while some carbonation was still occurring at 400°C when the test was- terminated at t* ~ 1.4. The smaller slope of the breakthrough curve is another indication that the global rate of reaction is limited by kinetics at the low temperature.

5.4.6 Effect of Carbonation/Shift Pressure

The effect of pressure during the combined reaction stage is shown in Figure 5-22 where breakthrough results from tests

27-1 (15 atm) , test 34 (5 atm) , and test 70 (1 atm) are Fractional COx Removal 1.0 1.1 0.9 0.7 0.8 0.6 0.4 0.5 0.3 0.2 0.1 0.0

. 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0

Figure 5-21. Change in the Shape of the Carbon Oxide Carbon the of Shape the in Change 5-21. Figure FEED COMPOSITION FEED 00 H 20.0% 44 N 74.4% SV = 1425hr1 = SV .% CO 5.6% P = 15atm = P Breakthrough Curves With Shift/Carbonation With Curves Breakthrough Temperature. 2 2 O

Dimensionless time, t* time, Dimensionless et27,550°C 7 2 Test et37,0° . ,400°C 7 3 Test j 1 L

147 Concentration, ppm 105 3 0 1 104 101 102 100 . 02 . 06 . 1.0 0.8 0.6 0.4 0.2 0.0 Figure5-22. Log-plotof Breakthrough Showing Curves the Effect ofCarbonation Pressure. Dimensionless Time, t* Time, Dimensionless /■ et34,5atm 4 3 Test 5atm - ,1 -1 7 2 Test - CO ° CO o - Test 70,1 atm 70,1 Test -

2 . 1.4 1.2

==a==a- 148 compared. Table 5-10 presents the reaction conditions of the three tests. The ability to remove carbon oxides during the prebreakthrough portion of the cycle is of primary interest.

At the standard reaction pressure of 15 atm (test 27-1) the prebreakthrough CO concentration was at or below 10 ppm for t*

< 0.58 while the C02 concentration was at or below 200 ppm for the same time span. These concentrations correspond to approximately 99.7% removal of total carbon oxides. At 5 atm

(test 34), the driving forces for the reactions were lower and the prebreakthrough concentrations were approximately 100 ppm

CO and 900 ppm C02, values which correspond to about 98.5% carbon oxide removal. Finally, during test 70 (1 atm), the rate of combined shift/carbonation reactions was very slow and no clear breakthrough was observed. In the early portion of shift/carbonation (t* < 0.18), both the CO and C02 concentrations were about 0.5% and the corresponding COx removal was about 85%. When t* reached about 0.9, the C0X removal decreased to 44% and CO and C02 concentrations were

3.4% and 0.65%, respectively.

The slopes of the breakthrough curves are smaller at low pressure which is indicative of a reduced global rate. The decrease is due to the decrease in gas concentrations and residence time. If we assume that the rate of the shift reaction is first order with respect to CO and H20, there is a factor of 15x15 decrease in the shift reaction rate for test

70 compared with test 27-1, while the factor is 3x3 for test Table 5-10

Summary of Reaction Test Conditions for the Effect of Carbonation Pressure

Test 27-1 34 70

Sorbent D D D Initial Mass, g 13 .2 13 .2 13.2 Particle Size 149- 149- 149- Range, fim. 210 210 210

Calcination Phase

Temp., °C 750 750 750 Pressure, atm 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 Flow Rate, seem 700 700 • 700

Carbornation Phase

Temp., °C 550 550 550 Pressure, atm 15 5 1 Feed Gas Comp., %C0 5.6 5.6 5.6 %C02 0 0 0 %h 2 0 0 0 %h 2o 20.0 20.0 20.0 %n 2 74.4 74.4 74.4

Flow Rate, seem 250 250 250

Space Velocity (STP) , hr'1 1425 1425 1425 151

34. In addition, lower pressure reduces the gas residence time, which also decreases the CO conversion. As a result, the global rate of reaction for test 70 was far less than that of test 27-1. In the postbreakthrough period, where carbonation has ended and the water-gas shift is the only reaction occurring, pressure is also important. The shift reaction proceeds to a greater extent at high pressure due to the effect of pressure on concentration and, hence, kinetics.

5.4.7 Effect of Carbonation/Shift Space Velocity

Table 5-11 shows the test conditions with the space velocity as the parameter during the shift - carbonation reaction stage. An increase in the space velocity reduces the gas residence time in the reactor, thereby reducing the time available for both the shift and carbonation reactions to occur. In commercial processes, increasing space velocity means a smaller reactor and reduced capital costs. Hence the objective is to find largest possible space velocity without significantly sacrificing COx removal. Prebreakthrough CO and

C02 concentrations as a function of space velocity are shown in Figure 5-23. Except at the lowest space velocity, the concentrations of both components increase with increasing space velocity. Since temperature and feed gas composition were constant, equilibrium concentrations are independent of space velocity as indicated by the horizontal dashed lines.

This type of response is expected since residence time ! i

Table 5-11

Summary of Reaction Test Conditions for the Effect of Carbonation Gas Flow Rate (N H Test 1 38, 39, 40, 41

Sorbent D D D D D Initial Mass, g 13.2 13.2 13 .2 13 .2 13.2 Particle Size 149- 149- 149- 149- 149- Range, /xm 210 210 210 210 210

Calcination Phase Temp., °C 750 750 750 750 750 Pressure, atm 3.3 3.3 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 100 100 Flow Rate, seem 700 700 700 700 100

Carbomation Phase Temp., °C 550 500 600 400 650 Pressure, atm 15 15 15 15 15 Feed Gas Comp., %C0 5.6 5.6 5.6 5.6 5.6 %C02 0 0 0 0 0 %h 2 0 0 0 0 0 %h 2o 20.0 20.0 20.0 20.0 20.0 %n 2 74.4 74.4 74.4 74.4 74.4

Flow Rate, seem 250 150 350 450 600

Space Velocity (STP) , hr'5 1425 885 1995 2665 3420 152 Prebreakthrough Concentration (dry basis), ppm 300 100 200 00 50 00 50 3000 2500 2000 1500 1000 Figure 5-23. Prebreakthrough Concentrations of CO of Concentrations Prebreakthrough 5-23. Figure experimental equlibrium □ and C02 as a Function of Space Velocity. Space of C02and Function a as pc eoiy 1 (STP) r1h Velocity, Space t r CO 2 FEED COMPOSITION FEED 00 H 20.0% 44 N2 74.4% P = 15atm = P T = 550°C = T .% CO 5.6% 2 O

3500

153 154

decreases with increasing space velocity. The increase in CO

and C02 concentrations at the lowest space velocity was

unexpected, and, for this reason, duplicate tests were

performed with essentially the same results. It is

significant to note that even at the largest space velocity of

3420 hr*1, the fractional removal of total carbon oxides was

approximately 0.995 during the prebreakthrough period.

Postbreakthrough concentrations of CO and C02 are also

dependent on space velocity as shown in Figure 5-24.

Equilibrium conversion is again independent of space velocity

as represented by the horizontal dashed lines. As space

velocity increased, there was a downward trend in the C02

concentration and an upward trend in the CO concentration.

These effects are consistent with the reduced time which is

available for the shift reaction to occur at large space

velocity.

Increased space velocity alters the shape of the

breakthrough curve in much the same way as lower temperature

(Figure 5-25) . At the highest space velocity (3420 hr*1) ,

initial breakthrough occurred at t* ~ 0.7, and t* ~ 1.45 was

required for complete carbonation. The breakthrough curve at

the lower space velocity (885 hr*1) was steeper; initial breakthrough occurred at t* ~ 0.85 and complete carbonation at

t* ~ 1.15. At the intermediate space velocity of 1425 hr*1

(not shown in Figure 5-25) initial breakthrough occurred at t*

~ 0.77 and reaction was complete at t* ~ 1.15. (0 8 *OT CO experimental JD 7 >» equlibrium 3 C 0 2 6 c o *-» • ■ duplicate results 5 T = 550°C -f-*I— c P = 15atm CD O 4 CO2 c o O 3 JZ O) =J CO FEED COMPOSITION * o 2 5.6% CO 20.0% H2O CO 1 74.4% N2 CD CO

(0 0 O 3500 CL 1000 1500 2000 2500 3000 Space Velocity, hr1 (STP)

Figure 5-24. Postbreakthrough Concentrations of CO and C02 as a Function of Space Velocity. 155 1.1 r i » r 1.0 0.9 0.8 Test 41, SV = 3420hrA iTest 38» sv = 885hr1 j 0.7 T = 550°C 0.6 P = 15atm 0.5 FEED COMPOSITION 0.4 5.6% CO 0.3 20.0% H2O 74.4% N2 0.2 0.1 0.0 0.0 0.2 0.4 0.6 0.8 1.0 1.2 1.4 Dimensionless time, t*

Figure 5-25. Change in the Shape of the Carbon Oxide Breakthrough Curves With Space Velocity. 157

5.4.8 Effect of Feed Gas Composition

Most of the tests discussed so far used the simple feed gas composition for the combined shift and carbonation reactions. The carbonation feed gas in test 24 contained all of the major components of coal gas (CO, C02, H2, H20, and N2) so that both the shift and carbonation reactions were free to occur simultaneously. All the other reaction parameters of test 24 were the same as the base test 27-1. In test 27-1, the carbonation feed gas contained only CO, H20, and N2 so that no carbonation could occur and no hydrogen could appear in the product gas unless the shift reaction occurred first.

Results in the form of component breakthrough curves versus time for tests 24 and 27-1 are presented in Figures 5-

26 and 5-27, respectively. The results are qualitatively similar except that the final concentrations of H2 and C02 were equal in test 27-1, but not equal in test 24. In test 24

(Figure 5-26), the final concentration of C02 was lower than

H2 because of the presence of H2 in the feed.

5.4.8.1 Effect of H20/C0 Ratio

The H20 to CO ratio is an important economic parameter since the addition of excess steam will add to the process cost. Two sets of runs were conducted to evaluate the effect, one for the simple feed gas (C0/H20/N2) and the other for the simulated coal gas. Table 5-12 shows the reaction conditions for the tests involved. Mole Percent (dry basis) 0 r- 30 0 - 20 0 - 10 5 30 390 370 350 iue52. eco epne ProductComposition 5-26.Figure Reactor Response. FEED COMPOSITION FEED for Test24. V=12hi (STP) 1425hri = SV 30 H 13.0% 08 N 60.8% 74 H 17.4% T = 550°C = T P = 15atm = P .% CO 4.4% .% CO 4.4% 1 40 450 430 410 Time,minute 2 2 2

O

2

7 40 510 490 470 Test 24 Test CO

2 158 8

7 Test 27-1

CO ‘co 6 CO A >. 5

4 c

0

350 370 390 410 430 450 470 490 510 Time, minute

Figure 5-27. Reactor Response. Product Composition for Test 27-1. 159 Table 5-12

Summary of Reaction Test Conditions for the Effect of Carbonation Gas Composition

Subgroup A Subgroup B

Test 27-1 42 43 53 54 55

Sorbent D D D DDD Initial Mass, g 13.2 13.2 13.2 13 .2 13.2 13.2 Particle Size 149- 149- 149- 149- 149- 149- Range, /im 210 210 210 210 210 210

Calcination Phase Temp., °C 750 750 750 750 750 750 Pressure, atm 3.3 3.3 3.3 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 100 100 100 Flow Rate, seem 700 700 700 700 700 700

Carbomation Phase Temp., °C 550 550 750 500 600 400 Pressure, atm 15 15 15 15 15 15 Feed 161

Simple Feed Gas

Prebreakthrough CO and C02 concentrations as a function

of the HjO to CO ratio in the feed gas are shown in Figure 5-

28. Increasing the ratio caused both the CO and C02

concentrations to decrease. The decrease in CO was due to the more favorable shift equilibrium and more rapid kinetics associated with the higher steam concentration. Although the

carbonation equilibrium is independent of the H20 to CO ratio, the C02 concentration was reduced because of the faster production of C02 via the shift reaction, and, as a consequence, the increased time available for the C02 to react with CaO.

Simulated Coal Gas

In tests 53, 54, and 55, the reactor feed contained all major constituents of a coal gas but with varying ratios of

H20 to CO (Table 5-12). The base composition for each test was:

CO: 8.0%

C02: 3.8%

H2: 5.6%

H20: 10.6%

N2: 72.0%

Except for the large N2 concentration, the ratios of CO, C02,

H2 and H20 are representative of a Texaco gas. Sufficient Prebreakthrough Concentration (dry basis), ppm 300 200 100 . . 25 . 3.5 3.0 2.5 2.0 1.5 FEED COMPOSITION FEED V=12hi (STP) 1425hri = SV Figure5-28. Postbreakthrough Concentrations ofCO T = 550°C = T P = 15atm = P CO - 5.6% - CO N H 2 2 - balance - O - as indicated as - O

H Ratio inRatio Feed Gas. and C02as aFunction of HaO to CO 2 O to CO Ratio in Feed Gas inFeed Ratio CO to O

CO CO 2 2 experimental equlibrium

4.0 162 163 additional H20 was then added to provide molar ratios of H20

to CO equal to 2, 3, and 4.

Breakthrough curves for test 53 (3 mol H20 per mol CO) as determined by the TCD are shown in Figure 5-29. The total C0X molar feed rate was approximately 85% greater than the C0X feed rate associated with the standard gas composition. In spite.of the larger feed rate, a prebreakthrough steady-state with CO and C02 concentrations below the TCD detection limit was established. The H2 concentration during this period was effectively constant at 15.2%, which is reasonably close to the value of 15.9% calculated by material balance assuming t complete conversion of CO and complete removal of C02.

Similar prebreakthrough steady states were achieved in tests

54 and 55.

The prebreakthrough CO and C02 concentrations as determined by the flame ionization detector for these runs are plotted versus the H20 to CO ratio in Figure 5-30.

Equilibrium concentrations are shown for reference purposes.

CO concentration decreased as the H20 to CO ratio increased from 2 to 4. The C02 concentration was relatively constant at each of the feed gas compositions. It should be pointed out that, due to the short prebreakthrough period for test 55, the concentrations of CO and C02 shown in Figure 5-30 involve increased uncertainty. The prebreakthrough concentrations from Figure 5-30 are appreciably higher than obtained in the previous tests at the same H20 to CO ratios but with lower Mole Percent (dry basis) 4 2 8 0 4 2 6 0 8 40 2 40 6 40 0 50 540 520 500 480 460 440 420 400 380 Figure5-29. Reactor Response. Product Composition for Test53. Time, minute Time, 2 0 C CO Test 53 Test Prebreakthrough Concentration (dry basis), ppm 300 0 0 1 200 Figure 5-30.Figure Postbreakthrough Concentrations ofCO CO 2 CO CO 2 H RatioinFeed Gas. and COa asaFunction ofHaO to CO 2 O to CO Ratio in Feed Gas Feed in Ratio CO to O experimental equilibrium CO 4 3 EDGSCMOIIN . COMPOSITION GAS FEED O CO CO 2 V=1425hr1 h 5 2 4 1 = SV H T = 550°C = T 5atm 1 = P 2 H 2 O

59.3 63.4 68.2 165 166

concentrations of total C0X in the feed gas. However, the

fractional removal of total carbon oxides was effectively

constant, varying only from 0.996 to 0.997 in the six tests.

5.4.9 Effect of Sorbent Particle Size

Three different sorbent particle sizes were tested in

runs 25, 26, and 27-1. Key reaction conditions for the tests are summarized in Table 5-13. Prebreakthrough concentrations of H2, CO and C02 were similar in all cases, with CO and C02 concentrations below the detection limit of the thermal

conductivity detector. In addition, the prebreakthrough H2 concentrations were similar in each test. Fractional sorbent conversion, X’, at t* = 1 and t* = tf* (final time) were also

similar in each run. However, the smallest particles (test

26) produced a shorter prebreakthrough time and the breakthrough curve had a smaller slope. These features are emphasized in Figure 5-31 where dimensionless breakthrough curves for tests 25 (largest particle size) and 26 (smallest particle size) are compared. In addition, the postbreakthrough concentrations appear to depend on sorbent particle size. Postbreakthrough H2 concentration increased from 3.8% with the smallest particles to 4.2% with the largest particles. Table 5-13

Summary of Reaction Test Conditions for the Effect of Sorbent Particle Size

Test 25 26 27-1

Sorbent D DD Initial Mass, g 13 .2 13.2 13 .2 Particle Size 250- 74- 149- Range, pirn 300 149 210

Calcination Phase

Temp., °C 750 750 750 Pressure, atm 3.3 3.3 3.3 Gas Comp., %N2 100 100 100 Flow Rate, seem 700 700 700

Carbornation Phase

Temp., °C 550 550 550 Pressure, atm 15 15 15 Feed Gas Comp., %C0 5.3 5.4 5.5 %C02 0 0 0 %h 2 0 0 0 %h 2o 20.1 20.0 20.0 %n 2 74.6 74.6 74.4

Flow Rate, seem 250 250 250

Space Velocity (STP) , hr’1 1425 1425 1425 (CO+CO2 ) Breakthrough, COX* 1.1 1.0 0.9 0.7 0.8 0.6 0.5 0.4 0.3 0.2 0.1 0.0 . 02 . 06 . 1.0 0.8 0.6 0.4 0.2 0.0 Figure 5-31.Figure Comparison ofDimensionless Breakthrough CurvesasaFunction ofSorbent Particle Size. Test 25 Test Test 26 Test Dimensionless Time, t* Time, Dimensionless m b b m 1.2 c-c -c c 1.4 1.1 1.0 0.9 0.7 0.8 0.6 0.5 0.4 0.3 0.2 0.0

Fractional Sorbent Conversion, X CO (Tv H 169

5.4.10 The Nature o£ the Shift Reaction

Although the experimental results proved that the shift

reaction occurred over the range of reaction conditions

studied, the driving force for the shift reaction is not

clear. The commercial shift processes (discussed in Chapter

2) are catalytic, but the reaction temperature is 150°C to

200°C lower than the temperatures used in this study. The

literature contains little information on the shift reaction

at 500°C or higher, because adverse equilibrium conversion makes the process impractical when only the shift reaction occurs. Kondrat'ev and Ziskin (1943) reported that the shift

reaction does occur at 700 and 800°C in the absence of a

catalyst. A number of metals and metal oxides, including iron and magnesia, are known shift catalysts (Rofer-DePooter,

1984) . However, no references to either CaO or A1203 catalyzing the reaction are known.

Several possible driving forces for the shift reaction at the conditions of the current study may be listed:

1. The shift reaction is strictly a homogeneous gas phase reaction;

2. The shift reaction is catalyzed by CaO, CaC03, and/or

MgO, which are the primary species present;

3. The shift reaction is catalyzed by the stainless steel reactor walls and/or the porous stainless steel disks used to separate the sorbent sections; 170

4. The shift reaction is catalyzed by the impurities

present in the calcined sorbents.

In order to learn more about the nature of the shift

reaction at high temperature, a series of tests using an empty

reactor as well as a variety of packing materials was carried

out. Reaction conditions were as shown in Figure 5-32, where

fractional conversion of CO via the shift reaction is

compared. The equilibrium fractional conversion is 0.91 under

these conditions. Results from the three tests in which the packing was capable of reacting with C02 (dolomite, limestone,

and marble chips) are based upon the postbreakthrough period

after carbonation was effectively complete.

From Figure 5-32, we see that the lowest CO conversion

occurred in the empty reactor test, which suggests that the

degree of the homogeneous shift reaction and the catalytic

effect of the reactor walls is small. A catalyst is required

if the shift reaction is to approach equilibrium under the

reaction conditions. However, other test results do little to clarify which compound provides the main driving force for the shift reaction. The higher CO conversion using dolomite compared to the marble chips suggests that MgO exerts a catalytic effect (marble chips contain little or no MgO). The higher CO conversion using limestone may be attributed to the impurities in the sorbent. If we compare the results of the empty reactor and the marble chips tests, it seems that CaO or

CaC03 also exerts some catalytic effect on the shift reaction. Equilibrium: 0 .91 T = 550 oC P = 15atm Feed Gas Composition - 5.6% CO 0.67 20.0% H2O -74.4% N2

Test 50 Test 51-1 Test 49 Test 58 Test 57 Test 61 Empty Dolomite AI2O3 Shift Limestone Marble Reactor Catalyst Chips

Figure 5-32. Fractional Extent of the Shift Reaction Using Various Reactor Packing. 172

It is surprising that the highest CO conversion occurred when the reactor contained A1203 because no previous work was found which claimed A1203 as a shift catalyst. The commercial iron*- cobalt shift catalyst, manufactured by United Catalyst, Inc., produced slightly lower shift conversion than did A1203.

However, the reaction temperature was significantly higher than the design temperature for the catalyst. Indeed, experimental results suggested that H2 formed by the shift reaction reduced the iron oxide in the catalyst to a lower oxidation state.

In order to clarify the shift mechanism, more tests have to be made under stringent conditions so that every possible driving force is isolated for the examination. The above

"compound oriented" analysis is obviously too rough. However, it does appear that heterogeneous catalysis is responsible, at least in part, for the shift reaction, but it is not possible to attribute the catalytic effect to a single component.

5.4.11 Sorbent Structural Change Along the Reactor Axis

The sorbent structure change along the reactor axis was studied by measuring pore volume and pore size distribution.

Two VA cycle tests (33 and 51) were specifically conducted for this purposes. Standard conditions were used for both tests except that calcination in test 33 was conducted in a mixture of H20 and N2, and for test 51, the calcination atmosphere was pure N2. The shift/carbonation phase of cycle 2 was 173

intentionally terminated at t* = X* ~ 0.4 to provide an axial

distribution in the extent of carbonation. Roughly speaking,

carbonation in the first (top) section should approach

completeness with a decreasing extent of carbonation in

sections 2 and 3 and essentially no carbonation in section 4

(bottom).

Pore size distribution curves as well as the pore volume

for each section of tests 33 and 51 are shown in Figure 5-33

(a through d) . All samples show a clear difference in the pore size distributions, with the H20/N2 calcination atmosphere producing pores which are approximately twice as large as the pores produced in N2. Total pore volumes were reasonably

similar in each atmosphere, and the pore volume increased from

top to bottom in accordance with the expected variation in extent of carbonation. The range of pore diameters in each of the four sections from test 51 was from approximately 0.05 to

0.3 /xm with the peak in the distribution curve dropping from about 0.18 /xm at minimum carbonation (bottom section) to 0.1

/xm at maximum carbonation (top section) . For test 33, the same trend was also observed and the range of pore diameters was from approximately 0.1 to 0.5 /xm with the peak dropping from 0.32 /xm in the bottom section to 0.22 /xm in the top section. 2.0 1.8 1.8 Toil 33 (Soolton ll. PV - 0.07cm’/g 7«it 33 (Sxcllon 2). PV - O.O7cm»/0 Toil SI (Section I). PV - O.OScm’/g *08 1.8 Tell 51 ISacllon 2). PV - O.OgcmVg - Xo 1.4 •4 1.2 J 3> i JO I « * O.B >7J 0.8 > £ 0.6 s2 0.8 a jj1 0.4 03 0.0 OJO I « * I i l «10» I 1 « I I • 101 X I I I < 10-1 1 > < • 10* Cfamoftr (mferon) Dtemoler (micron)

2 .0 2.0 Toil 33 (Section 3). PV - 0.19cm>/g Test 33 (Section 4). PV - 0.28cmVg Toil 51 (Socllon 3). PV - 0.1 ScmVg 1.8 Test 61 (Section 4), PV - 0.32cm*/o i 1.4 3 1.4 8* $ 1 1 ,0 71 0.8 5 0 .6 «2 —• 0.8 5 0.8 5 ? 0.4 3 M 02 0.0 0.0 x * •• • > > i ■ i t x * •• I I 10 » x I .11 10-1 x I < I to-r Olimotor (micron) . Otomotor (micron)

Figure 5-33. Pore Volume and Pore Size Distribution Curves From Axial Sections for Sorbent Calcined in Na (test 51) and H20/Na (test 33). 175

5.5 Conclusions

The single cycle tests confirm the results established by the preliminary studies discussed in the previous chapter, and provide experimental confirmation of the simultaneous occurrence of the shift and carbonation reactions.

Calcination can be conducted over a temperature range of 750°C to 900°C under either N2 or a mixture of N2 and steam. The calcination reaction becomes very slow at 650°C. A higher calcination temperature tends to increase C02 product concentration while reducing the required amount of calcination sweep gas. No significant adverse effect of high calcination temperature was observed in the first carbonation

shift cycle. Equilibrium for the combined shift carbonation reactions can be closely approached at 15 atm, in the temperature range of 500°C to 600°C and at space velocities as high as 3400 hr'1 (STP) . Total prebreakthrough concentrations of carbon oxides of approximately 300 ppm (dry basis) in the reactor product were routinely achieved while concentrations of less than 50 ppm (dry basis) were achieved at the most favorable conditions. These concentrations correspond to factional removal of total carbon from about

0.995 to greater than 0.999.

Carrying out the shift and carbonation reactions in a single processing vessel without the addition of a catalyst provides the basis for a possible alternate and simpler process for the production of hydrogen from synthesis gas. Although low cost dolomite is used as the C02 sorbent precursor, it will be necessary, if the process is to be economical, for the sorbent to retain structural integrity and activity through numerous calcination-carbonation cycles.

Multicycle tests to evaluate the sorbent durability will be discussed in the next chapter. CHAPTER 6

EXPERIMENTAL RESULTS AND DISCUSSION: MULTICYCLE TESTS

This chapter presents the results of multicycle tests to evaluate sorbent durability. The reaction conditions were chosen based on the single cycle studies described in the previous chapter. Results of five cycle tests using dolomite and marble chips are first discussed and multiple methods to compare sorbent durability are described. Effects of reaction parameters on sorbent durability over multicycle tests are covered in sections 6.2 to 6.5. Finally, two tests with favorable conditions were extended to eleven and ten test cycles.

6.1 Comparison of Sorbent Performance on Five-Cycle Tests

Table 6-1 summarizes the reaction conditions for all the multicycle tests. As discussed in the previous chapters, three different sorbent precursors -- dolomite, limestone and marble chips -- were selected for this research. Single cycle tests indicated that the reactivity of limestone was very poor compared with dolomite and marble chips; the presence of impurities in the limestone made it impossible to calculate dimensionless breakthrough curves for sorbent performance comparison. Hence the limestone was not chosen for multicycle tests. Although no detailed composition analysis was

177 Table 6-1

Summary of Reaction Conditions for Multicycle Tests

65 66 67 68 69 72 Test 27 56 60 62 63 64 D D D D D D Sorbent* D D D D MC D 11 5 5 5 5 10 Cycles 5 5 5 5 5 5 13.2 13.2 13.2 13.2 13.2 13.2 Initial Mass, g 13.2 13.2 13.2 13.2 13.2 13.2 149- 149- 149- 149- 149- Particle Size 149- 149- 149- 149- 149- 149- 149- 210 210 210 210 210 210 Range, pm 210 210 210 210 210 210

Calcination Cycle 800 850 850 900 750 750 750 750 750 750 750 750 Temp., ”C 3.3 3.3 3.3 I 3.3 3.3 3.3 3.3 3.3 3.3 3.3 3.3 Pressure, atm. 100 100 50 100 100 100 100 100 100 100 100 100 Gas Comp., % N2 0 0 50 0 % HjO 0 0 0 0 0 0 0 0 700 700 700 700 250 700 Flow Rate, SCCM 700 700 700 700 700 700

Carbonation Cycle 550 550 550 550 550 550 Temp., °C 550 650 600 550 550 550 15 15 15 15 1 Pressure, atm. 15 15 15 15 15 15 15 Feed Gas Comp., 6.6 6.6 13.7 5.6 5.6 5.6 5.6 5.6 5.6 6.6 6.6 %CO 5.6 3.1 3.1 5.1 0 0 0 0 0 0 0 3.1 3.1 % co2 4.6 4.6 4.6 8.4 0 0 0 0 0 0 0 4.6 % h 2 20.0 26.4 26.4 26.4 26.4 51.8 % h 2o 20.0 20.0 20.0 20.0 20.0 14.0 74.4 59.3 59.3 59.3 59.3 21.0 % n 2 74.4 74.4 74.4 74.4 74.4 80.4 250 250 250 250 250 96.6 Flow Rate, SCCM 250 250 250 525 250 250 1425 1425 1425 1425 1425 551 Space Velocity 1425 1425 1425 3000 1425 1425 (stp), h r' 178 'Sorbent D: dolomite MC: marble chips 179 available, it was assumed that marble chips contain mainly

calcium carbonate.

Two five cycle tests (27 and 63) using dolomite and marble chips were conducted under standard conditions. While there are a number of ways to evaluate durability, all show a gradual decrease with cycle number.

Results of Base Multicycle Run -- Test 27

Figure 6-1 compares the CO and C02 breakthrough curves for the first and fifth cycles of test 27. Data from the thermal conductivity and flame ionization detectors are combined on this semi-log plot to show the complete breakthrough curves ,with concentrations ranging from approximately 10 to 100,000 ppm (0.001% to 10.0%). Flame ionization detector results were used for all concentrations below 5000 ppm (0.5%) with thermal conductivity results used at the higher concentrations.

In the first cycle, CO concentration was at or below the

10 ppm level for the first 72 minutes. Thereafter the traditional S-shaped breakthrough curve was observed and the

CO concentration increased to 26,000 ppm (2.6%) at the end of the test. Initial CO concentrations in cycle 5 were also near the 10 ppm level, but initial breakthrough occurred after 30 minutes. The slope of the breakthrough curve decreased indicating a somewhat slower global reaction rate and the final CO concentration was 33,000 ppm (3.3%), considerably Mol Fraction (dry basis), ppm 105 1 0 1 X X © — © — © Figure 6-1. Comparison of CO and C02and CO Breakthrough of Comparison 6-1. Figure 0 0 0 0 100 80 60 40 20 X - First Cycle - First * Fifth Cycle * Fifth CO 2 0 C 1 1 1 1 of Multiple-Cycle Test. Multiple-Cycle of Cycles Fifth and First the in Curves t Time, minute Time,

[~£J 0

0 ^ I Test 27 Test 0 140 1 20 1

* 180 181

higher than in the first cycle. At the end of the run only

the shift reaction was occurring, and the result suggests that

sorbent aging has a negative effect on the kinetics of the

shift reaction at conditions were carbonation does not occur.

Prebreakthrough C02 concentrations were similar in both

the first and fifth cycles, and were in the general range of

90 to 220 ppm (0.009 to 0.022%) . However, C02 breakthrough in

the fifth cycle clearly occurred at an earlier time. Final

C02 concentration in the fifth cycle was lower than in the

first cycle, in agreement with the previously described reduction in the global kinetics of the shift reaction.

An alternate way of comparing sorbent durability is shown

in Figure 6-2 where CO and C02 breakthrough time is plotted versus cycle number. The breakthrough times in this figure are arbitrarily defined as the times required for the CO product gas composition to reach 100 ppm and for the C02 product gas composition to reach 500 ppm. Over the five cycles the C02 breakthrough time decreased from 94 to 75 minutes while the CO breakthrough time decreased from 85 to 55 minutes. These numbers represent an average decrease in breakthrough time of from 4% to 7% per cycle.

Still another means of evaluating multicycle durability is provided in Figure 6-3 where fractional calcium conversion,

X*, is plotted versus cycle number at two dimensionless times, t* = 0.5 and t* = 1.0. The fact that X’ = 0.5 at t* = 0.5 indicates that effectively complete carbon oxide removal was Breakthrough Time, minute 100 80 60 20 0 Test 27 Test Figure 6-2. CO and C02 Breakthrough Times as a as Times C02and CO 6-2.Breakthrough Figure 1 CO at 100 ppm in ppm 100 at CO product (dry basis) (dry product Function of Cycle Number. Cycle of Function 2 Cycle Number Cycle a- (dry basis) (dry a- CO

2 at 500 ppm inppmproduct 500 at 4 3 5

6 182 Fractional Calcium Conversion, X 0.6 0.7 0.8 0.9 0.5 0.4 0.3 0.2 0.1 0.0 .0 Figure 6-3. Fractional Calcium Conversion at Selected at Conversion Calcium Fractional 6-3. Figure Test 27 Test Dimensionless Times as a Function of Function a as Times Dimensionless Cycle Number. Cycle Cycle Number Cycle t* = 1.0 = t* t* = 0.5 = t* 5

6 183 184

achieved during the early reaction stages of all cycles. At t* = 1.0, the value of X* increased from 0.83 in cycle l to a maximum of 0.84 in cycle 3 and then decreased to 0.79 in cycle

5. Thus the average sorbent capacity decrease was only about

1% per cycle. For practical purposes, the carbonation reaction was complete at t* = 1.0 since final values of X* ranged from 0.84 in cycle 1 and 3 to 0.80 in cycle 5.

Results of Five Cvcle Test Using Marble Chios

Because of the high C02 capacity and relatively favorable first-cycle reactivity, as described in the previous chapter, a five-cycle test (test 63) was conducted at standard reaction conditions using marble chips as the sorbent precursor.

Figure 6-4 compares the CO and C02 breakthrough curves from cycles l and 5 (using both the FID and TCD analyzers) . A prebreakthrough steady-state was reached in cycle 1 with CO concentration in the range of 30 to 40 ppm and C02 concentration of about 200 ppm. These concentrations, which correspond to 99.7% carbon oxide removal, compare favorably with the prebreakthrough concentrations obtained using dolomite. By the fifth cycle, however, no steady-state postbreakthrough period was observed. Instead, both the CO and C02 concentrations increased continuously after the reaction was initiated. I |

Mol Fraction (dry basis), ppm 105 104 103 2 0 1 101 101 0 0 1

. 03 . 07 . 11 . 1.5 1.3 1.1 0.9 0.7 0.5 0.3 0.1

Figure 6-4. Comparison of CO and C0 and CO of Comparison 6-4. Figure 0 - 0 Curves in the First and Fifth Cycles Fifth and theFirst in Curves of Test 63. Test of DimensionlessTime, t* O O O O a O O Q O O 2 Breakthrough - First Cycle First - - Fifth Cycle Fifth - CO CO 2 Test 63 Test

185 186

Poor multicycle durability of CaC03 sorbent precursors was observed in the previous studies using a TGA reactor

(Silaban, 1993), and the results shown in Figure 6-4 confirm those observations. Further evidence of the more rapid deterioration of the high CaC03 sorbent compared to dolomite is shown in Figure 6-5 where the dimensionless times, t*, at which the CO and C02 concentrations in the reactor product reached 100 and 500 ppm, respectively, are plotted versus cycle number. Data from test 27 are included for comparison.

The negative slopes of each of the four lines indicate poorer performance of both sorbents in each succeeding cycle.

Although first-cycle times are comparable, the performance of marble chips decreased more rapidly and the superior durability of the dolomite is clearly evident.

6.2 Effect of Carbonation/Shift Temperature

Tests 56 and 60, coupled with earlier test 27, provide a series of five-cycle tests in which the shift/carbonation temperature was varied while standard conditions were used for all other reaction parameters. A prebreakthrough steady-state was achieved in all cases and prebreakthrough CO and C02 concentrations as a function of cycle number are shown in

Figure 6-6 and 6-7, respectively. At both 550°C and 600°C, the prebreakthrough concentrations were reasonably constant as the number of cycles increased. The variation in the concentrations of both components at 650°C is attributed, at Dimensionless Time, t 1.0 0.9 0.7 0.8 0.4 0.5 0.6 0.3 0.2 0.0 0.1 1 0 Figure 6-5. Comparison of the Multicycle Deterioration of Deterioration the Multicycle of Comparison 6-5. Figure □ CO □ basis) .(dry ppm 100 at CO o ' 2 at 500 ppm (dry basis) (dry ppm 500 at ooie Ts 7 adMrl hp (Test63). Chips (TestMarble and 27) Dolomite 2 CycleNumber Test 27 3 Test 63 Test 4 5

6 187 ! I

Prebreakthrough CO Concentration, ppm 1600 1400 0 0 0 1 0 0 2 1 800 400 600 0 0 2 0

1 3 5 6 5 4 3 2 1 0 Figure 6-6. Prebreakthrough CO Concentration as a as Concentration CO Prebreakthrough 6-6. Figure jb. Function of Temperature and Cycle Number. Cycle and Temperature of Function -m- CycleNumber Test 60 (600°C) 60 Test -B- f ■ f ■ (ft ■ (ft ■ (ft Test 27 (550°C) 27 Test B ---

-a 188 Prebreakthrough CO 2 Concentration, ppm 6000. 4000 5000 1000 2000 3000 Figure 6-7.Figure Prebreakthrough C02 Concentration as a Function ofTemperature and Cycle Number.

Cycle Number Cycle 189 190 least in part, to analytical uncertainty. These concentrations, on the order of 1000 to 6000 ppm (0 .1% to

0.6%) are difficult to analyze with precision. The lower concentration limit for the TCD is about 5000 ppm while the upper concentration limit for methanation and FID analysis is about 1000 ppm. Although the quantitative values at 650°C are somewhat uncertain, the results suggest improved performance as the number of cycles increased. The increase in prebreakthrough concentrations with increasing temperature reflects the effect of temperature on the equilibrium of the combined reactions.

Figure 6-8 shows the fractional removal of total carbon oxides, FCOX, during the prebreakthrough period as a function of temperature and cycle number. At 550°C, , FCOX was equal to

0.996 in each of the five cycles. At 600°C, FCOX varied between 0.979 and 0.986, while at 650°C, FCOX increased from

0.904 in cycle 1 to 0.941 in cycle 5. The apparent increase at 650°C is due to the decreasing C02 concentrations shown in

Figure 6-7.

While the relatively constant prebreakthrough concentration results are quite positive, some deterioration in sorbent performance with cycle number was observed. In general, the duration of the prebreakthrough period decreased with increasing cycle number. This is shown in Figures 6-9 and 6-10. In Figure 6-9, fractional carbon oxide removal

(based upon TCD analysis) is plotted versus dimensionless time Fractional Carbon Oxide Removal, FCOX 1.00 0.98 0.94 0.96 0.92 0.90 0.88 Figure 6-8.Figure FractionalRemoval ofCarbonTotal Oxides as 10 aFunction ofTemperature and Cycle Number. 2 CycleNumber 4 3 5 6

161 Fractional COx Removal 0.8 0.9 0.7 0.6 0.4 0.5 0.3 0.2 0.0 0.1 0.0 Figure 6-9. Fractional Removal of Total Carbon Oxides Carbon Total of Removal Fractional 6-9. Figure 0.2 During the First and Fifth Cycles of Test 56. Test of Cycles Fifth and First the During 0.4 Dimensionless Time, t* Time, Dimensionless 0.6 5th cycle 5th 0.8 . 1.2 1.0 1cycle st

1.4 192 Concentration, ppm 0.0 et6 (600C) 60 Test CO • CO ■ Figure 6-10.Figure CO and 0.2

2 Fifth Cycle Fifth First Cycle First 0.4 Firstand Fifth Cycles60.ofTest Dimensionless t* Time,

C03 0.6 Breakthrough fromCurvesthe 0.8 1.2

1.4 193 194

for cycles l and 5 of test 56 (650°C) . In cycle 1, active

breakthrough occurred over the range of t* values between 0 .9

and 1.2, while in cycle 5 breakthrough began at t* ~ 0.7.

Another measure of sorbent deterioration is that the final

conversion of CaO to CaC03 decreased from 0.89 in cycle l to

0.78 in cycle 5.

In Figure 6-10, CO and C02 breakthrough curves, using

combined FID and TCD results, from the first and fifth cycles

of test 60 (600°C) are plotted versus dimensionless time.

Prebreakthrough CO concentrations were approximately constant

at 400 ppm in both cycles, but the CO concentration reached

1000 ppm at t* = 0.96 in cycle 1 and at t* = 0.76 in cycle 5.

The prebreakthrough C02 concentration was actually lower in

cycle 5, 570 ppm versus 890 ppm, but the duration of the C02

prebreakthrough period was shorter in the fifth cycle.

The deterioration in performance with increasing cycle

number was most severe at 650°C. This is illustrated in

Figure 6-11 where the dimensionless time at which the

fractional carbon oxide removal decreased to 70% is plotted versus cycle number. At 650°C, this occurred at t* = 0.99 in

the first cycle, but at t* = 0.78 in the fifth cycle. At

600°C, the decrease was from 0.99 in cycle 1 to 0.87 in cycle

5. All values of t* at 550°C were somewhat lower but deterioration was moderate, from t* = 0.83 in cycle 1 to t' =

0.75 in cycle 5. These lower values of t* at 550°C are the results of the achievable fractional conversion of CaO being t* corresponding to FCOX = 0.7 1.0 0.9 0.8 0.7 Figure 6-11.Figure Sorbent inDeterioration Multicycle 2 4 5 4 3 2 1 TestsasaFunction of Temperature. Cycle Number Cycle 196 lower at the low temperature. For example, the final fractional conversions of CaO for the first cycles of the three tests were 0.84, 0.99, and 0.89 at 550, 600, and 650°C, respectively. By the fifth cycle, the final fractional conversions were 0.80, 0.89, and 0.78.

Structural Property Measurements

At the conclusion of test 56, sorbent was removed from each of the four axial sections of the reactor and pore volume and pore size distribution were measured using the mercury porosimeter. The results are presented in Figure 6-12 (a through d ) . In all four sections, the pore volume was in the range of 0.02 to 0.03 cm3/g, compared to 0.36 cm3/g after initial calcination. No distribution curve was visible in the top two sections, and only small peaks in the 0.2 to 0.3 fim diameter range were present in the bottom sections. The lower pore volumes for test 56 are consistent with the fact that the shift/carbonation reaction was carried out until carbonation was essentially complete. The overall fractipnal calcium conversion at the end of the fifth cycle was 0.77 and the degree of carbonation was relatively independent of axial position.

6.3 Effect of Carbonation/Shift Space Velocity

In test 62, dolomite was tested over five cycles at the larger space velocity of 3000 hr*x (STP), more than double the i ! i

Log Dill. Vol (dWd.ogO) Onl/g) Log OiM. Vol (dV/dLogV) (ml/g) . • 0.0 0.1 0.2 0.5 0.6 0.7 0.8 0.9 1.0 0.0 0.2 0.9 0.9 t.o 0.6 0.8 et5 Scin3.P .3 c’/g cm’ .031 - (Section 3). PV 56 Test Test 56 (Section t J. PV - .022 cm*/g t .022 J. (Section 56 - Test PV 10-1 Figure 6-12. Pore Volume and Pore Size Distribution Size Pore and Volume Pore 6-12. Figure Diameter, um 10 to-* Cycle of Test 56. Test of Cycle of Axial Sorbent Samples Following Fifth Following Samples Sorbent ofAxial * X3 3 4SS 3 4S • ( Q O TJ > o Ol o» 0.3 a • - > TJ 0-5> a 0< * => o o> 8 E o> * 0.1 0.0 0.2 0.5 0.7 0.9 1.0 0.2 0.7 0.9 1.0 S • S • Test 56 (Seelton 4). PV - .030 cmJ/g .030 - 4). (Seelton PV 56 Test etS Scln2.P .2 m3/o cm .027 - S62). (Socllon PV Test to** 10 * 1 1 9 < S •3 <3S • 4I I Diameter, um Diameter, um 10 10

* -' s i s i s 3

4S 197. 198

standard value of 1425 hr'1 (STP). Standard values were used

for the other reaction parameters. Previous single-cycle tests at the larger space velocity showed only a- small increase in prebreakthrough CO and C02 concentrations. The objective of the multicycle test, therefore, was to determine the effect of the larger space velocity • on sorbent durability.

CO breakthrough curves based upon the combined FID and

TCD analyzers for the five cycles are shown in Figure 6-13.

Because of the large space velocity there was little evidence of a prebreakthrough steady state; instead, CO concentrations increased continually with time. The important result from

Figure 6-13 is that breakthrough performance in cycles 2 through 5 was quite similar; little or no deterioration in sorbent performance occurred after the first cycle. CO concentrations during the early stages of the reaction were in the range of 10 to 40 ppm, and were only marginally larger than the prebreakthrough CO concentrations found in each cycle of the comparable test at 1425 hr'1 (STP) space velocity.

The fact that little or no deterioration in sorbent performance occurred after the first cycle is emphasized in

Figure 6-14 where the dimensionless times required for the CO and C02 concentrations in the product gas to reach 100 ppm and

500 ppm, respectively, are plotted versus cycle number.

Comparable results from test 27 at 1425 hr'1 (STP) space velocity are included for comparison purposes. Except for the CO Concentration (dray basis), ppm 105 103 104 2 0 1 1 0 1 10°

. 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 Figure 6-13.Figure CO Breakthrough Curvesforthe Five Cycles62.ofTest

Dimensionless Time, t‘ Time, Dimensionless 199 i !

Dimensionless Time, t 0 . 1 0.9 0.8 0.7 0.6 0.3 0.4 0.5 0.2 0.1 0.0 0 ; o CO at 100 ppm (dry basis) (dry ppm 100 at CO o ; o CO o . Figure 6-14. Comparison of the Multicycle Deterioration the Multicycle of Comparison 6-14. Figure 2 1 at 500 ppm (dry basis) (dry ppm 500 at of Dolomite as a Function of Space Velocity. Space of Function a as ofDolomite 2 Cycle Number Cycle ©- 4 3 — B Test 27 Test V=1425hr1 1 = SV Test 63 Test SV = 3000hri = SV 5

6 200 201

fact that the decrease between cycles 1 and 2 was more

significant at 3000 hr*1 (STP) ; the overall results for the two

tests are quite similar. Small improvements in performance

actually occurred in test 62 between cycles 3 and 4, and between cycles 4 and 5. Thus, we conclude that increased

space velocity does not' have a negative effect on sorbent durability.

6.4 Effect of Carbonation/Shift Gas Composition

As in the singe-cycle studies, the five-cycle tests using various feed gas compositions focused on the effect of the H20 to CO ratio using the simple gas composition and a comparison of the simple gas with a simulated coal gas.

H,0 to CO Ratio

Economic considerations will require that the shift- carbonation step operate with the lowest H20 to CO ratio consistent with effective shift reaction and the production of

H2 product of the required purity. In most previous multicycle tests, such as test 27, the "standard" feed gas composition having a H20 to CO ratio of 3.57 was used. In test 64, the H20 content was reduced to provide a H20 to CO ratio of 2.5, approximately 30% lower than the standard value.

Prebreakthrough CO and C02 concentrations along with the fractional removal of total carbon oxides in the prebreakthrough period of test 64 are shown as a function of 202

cycle number in Figure 6-15. CO concentrations (dry basis)

ranged from 29 ppm in cycle 1 to 62 ppm in cycle 5. The

maximum C02 concentration was 274 ppm in cycle 2 while the

minimum of 214 ppm occurred in cycle 4. The combined COx

concentrations corresponded to 0.995 to 0.996 fractional

removal of total carbon oxides. These concentrations are only

marginally higher than those achieved in earlier tests using

a HjO to CO ratio of 3.57 and otherwise equivalent conditions.

The fractional carbon oxide removal in those tests ranged from

0.996 to 0.998.

Only moderate sorbent deterioration occurred during the

five cycles as indicated in Figure 6-16, where fractional

carbon oxide removal (as determined by TCD analysis) and

fractional conversion of CaO are plotted as a function of

dimensionless time for the first and fifth cycles of test 64.

(C0+C02) concentrations were below the TCD detection limit until t* ~ 0.87 in cycle 1 and until t* ~ 0.73 in cycle 5. The

slope of the breakthrough curve was also somewhat smaller in

cycle 5. When FCOX - 1.0, X* ~ t* and the fractional sorbent

conversion curves for the two cycles overlapped; the curves began to deviate once COx breakthrough began. At t* = 1.1, X*

in cycle 1 was 0.93 and the comparable value in cycle 5 was

0.89. Prebreakthrough Concentration, ppm (dry basis) 0 0 1 300 200 Figure 6-15.Figure Prebreakthrough CO and COa Concentrations Test 64 Test 2 1 and FractionalRemoval of TotalCarbon Oxidesasa Function ofCycle Number.64. Test Cycle Number Cycle CO 3 CO 2 4 5 1.00 0.99 0.98 0.97 0.96

Fractional Removal of COx OJ o to Fractional COx Removal . 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0 Figure 6-16. Fractional COx Removal and Fractional Sorbent Fractional COxand Removal Fractional 6-16. Figure Dimensionless Time in First and Fifth Cycles Fifth and First in Time Dimensionless Calcium Conversion as a Function of Function a as Conversion Calcium fTs 4 NJ 64 Test of DimensionlessTime, t*

1.0 1.1 6 . 0 8 . 0 0.1 0.4 0.7 0.9 2 . 0 0.3 0.5 0 . 0

Fractional Sorbent Conversion, X O 205

Simulated Coal Gas Feed

In test 6 6 , the reactor feed contained all four major

components of coal gas -- CO, C02, H2, and H20. Reaction

conditions for this five cycle test were, with the exception of the shift-carbonation feed gas composition, identical to

those used in the previous test 27. The primary purpose was

to compare five-cycle sorbent durability for the simple and more complex feed gases.

The total carbon oxide content of the feed gas in test 6 6

(6 .6 % CO and 3.1% C02) was 73% greater than used in test 27.

As a consequence, the prebreakthrough CO and C02 concentrations were somewhat larger, and the duration of the prebreakthrough period (in dimensional time) was decreased.

Sorbent durability, however, did not suffer.

Prebreakthrough CO and C02 concentrations as well as fractional removal of total carbon oxides for test 6 6 are plotted as a function of cycle number in Figure 6-17. CO concentrations ranged from 46 to 135 ppm while C02 concentrations ranged from 225 to 262 ppm; the average prebreakthrough concentrations were 101 and 240 ppm for CO and

C02, respectively. Although the prebreakthrough concentrations, particularly for CO, were considerably higher than those for test 27, the fractional removal of total carbon oxides remained extremely high, varying only between 0.997 and

0.998 for the five cycles. Prebreakthrough Concentration (dry basis), ppm 300 Figure 6-17. Prebreakthrough CO and C0 and CO Prebreakthrough 6-17. Figure and Fractional Removal of Total Carbon Oxides Carbon Total of Removal Fractional and as a Function of Cycle Number. Test Test Number. Cycle of Function a as Cycle NumberCycle 2 Concentrations Test 66 Test 6 6 . 1.00

Fractional Removal of Carbon Oxides, FCOx C\ o to 207

Sorbent durability for test 6 6 is illustrated in Figure

6-18 where fractional carbon oxide removal and fractional

sorbent conversion for the first and fifth cycles are plotted versus dimensionless time. Product gas analysis for this

figure is based upon the TCD. C0+C02 concentrations were below the TCD detection limit for t* s 0.75 in both cycles.

The global rates of the combined reactions were approximately equal in both cycles as indicated by the similar slopes of the breakthrough curves. The obvious changes over the five cycles were the slightly earlier breakthrough in the fifth cycle and the reduction in final fractional CaO conversion, from 0.91 in

cycle 1 to 0.85 in cycle 5.

6.5 Effect of Calcination Conditions

Test 67, which also lasted five cycles, duplicated the reaction conditions of test 6 6 except that the calcination temperature was 800°C instead of the standard 750°C. Although the previous TGA studies (Silaban, 1993) suggested that higher calcination temperature was detrimental to sorbent durability, a high calcination temperature is needed in order to reduce the required flow rate of calcination sweep gas. Figure 6-19 compares the CO and C02 breakthrough curves from the first and fifth cycles of test 67. In this figure, concentrations are based upon the combined FID and TCD analyzers. The general shapes of the curves are similar for each cycle, and gradual Fractional COx Removal 0 . 1 0.7 8 . 0 0.9 0.5 6 . 0 0.3 0.4 2 . 0 0.0 0.1 .1 Figure 6-18. Fractional CO, Removal and Fractional Sorbent Fractional CO,and Removal Fractional 6-18. Figure 0.0 0.2 Time in First and.Fifth Cycles of Test Test of Cycles and.Fifth First in Time Dimensionless of Function a as Conversion Calcium 0.4 DimensionlessTime, t* 0.6 0.8 1.0 1.2 -Cycle 1 -Cycle •Cycle 5 •Cycle et6 : 66 Test 6 6 . 1.4

0.9 0.7 0.8 0.5 0.6 0.0 0.2 0.3 0.4 0.1

Fractional Sorbent Conversion, X CO o to Mol Fraction (dry basis), ppm 1 0 1 0.0 Figure 6-19. CO and C0 and CO 6-19. Figure 0.2 Fifth Cycle Fifth First Cycle First CO CO 2 0.4 First and Fifth Cycles of Test 67. Test of Cycles Fifth and First Dimensionless Time, t Time, Dimensionless

. 08 .( 1 0.8 0.6 2 Breakthrough Curves from the from Curves Breakthrough t i Test 67 Test

209 210

sorbent deterioration is again shown by the shorter duration

of the prebreakthrough period during the fifth cycle.

Single cycle tests suggested that adding steam to the

calcination sweep gas would change the structure of the

calcined sorbent, but would not have a negative effect on the

subsequent shift/carbonation kinetics. Test 6 8 and 69 were

carried out to determine the effect of a further increase in

calcination temperature to 850°C using both dry N2 and a H20

/N2 mixture.

Increasing the calcination temperature and adding H20 to

the calcination atmosphere had, at most, a small effect on prebreakthrough CO and C02 concentrations and fractional

removal of carbon oxides as shown in Table 6-2. CO

concentrations in the 2 0 cycles of the four tests ranged from

46 to 170 ppm while C02 concentrations ranged from 174 to 340 ppm. The average fractional removal of total carbon oxides was

0.998 in tests 6 6 and 67 (750°C and 800°C calcination in N2)

and 0.997 in tests 6 8 and 69 (850°C calcination in N2 and

50%H2O/N2) .

The pattern of gradual sorbent deterioration in these

tests is illustrated in Figures 6-20 and 6-21. The parameter,

t i W , which is the ratio of dimensionless time required to reach a specified concentration in cycle i to the dimensionless time required to reach the same concentration in

cycle 1, is plotted versus cycle number. Results from tests

6 6 through 69 are included. With this representation, the Table 6-2 • | Prebreakthrough Concentrations of CO and C02 in Tests 6 6 Through 69 , f

Prebreakthrough CO Concentrations (ppm)

Test 6 6 Test 67 Test 6 8 ! Test 69 Cycle 750°C, N2 800°C, N2 850°C, N2 j 850°C, H20/N2

1 1 1 1 140 1 2 2 | 147 2 81 138 152 . 131 3 46 1 0 0 159 . 1 1 1 4 134 1 1 1 157 ! 118 5 135 117 170 | 126

Avearge Cone. 1 0 1 1 2 1 152 | 126

5rebreakthrough C02 Concentration (ppm)

Test 6 6 Test 67 Test 6 8 . Test 69 Cycle 750°C/ N2 800°C, N2 850°C, N2 j 850°C, H20/N2

1 225 2 1 2 288 i 340 2 262 240 258 | 295 3 245 174 242 j 286 4 234 185 246 242 5 234 143 278 ; 259

Avearge Cone. 240 191 262 ! 284

Avg. (co+co2) .998 .998 .997 j .997 Removal ti*/t1 * at CO = 200 ppm 1.0 1.1 0.9 0.7 0.8 0.6 0.5 0.3 0.4

Figure 6-20. Comparison of the Decrease in Sorbent in the Decrease of Comparison 6-20. Figure 0 et6 5 50%H20-50%N2 850 100%N2 69 Test 100%N2 850 68 Test ■0— 100%N2 800 67 Test ■A—750 66 Test — e — e 2 4 5 4 3 2 1 I ____ | _____ Number for Tests Tests for Number Reactivity for CO as a Function of Cycle of Function a as CO for Reactivity ep,C Comp. Temp.,°C I ___ Calcination: Cycle Number Cycle | _____ | ___ 6 6 | ______through 69. through | ___ | _____ I ___ L

212 ti*/t1 * at C O 2 = 500 ppm 1.0 1.1 0.8 0.9 0.7 0.6 0.5 0.4 0.3

Figure 6-21. Comparison of the Decrease in Sorbent in theDecrease of Comparison 6-21. Figure A et6 80 100%N2 100%N2 850 800 68 Test 67 ■A— Test — e 2 4 5 4 3 2 1 < . I . I < I et6 70 100%N2 750 66 Test et6 80 50%H2O-50%N2 850 69 Test Number for Tests Tests for Number Reactivity for C0 for Reactivity ep,C Comp. Temp.,°C Calcination: Cycle Number Cycle ____ 2 66 as a Function of Cycle of Function a as through 69. through I , ___ 1 L

213 214

dimensionless time ratio for the first cycle is always 1.0 and

the slope of the line provides an immediate visual image of

the rate of deterioration. In previous comparisons of this

type, the CO concentration reference was chosen to be 100 ppm.

However, in these tests which involved significantly higher

COx feed concentrations, the prebreakthrough CO concentration was often greater than 100 ppm (see Table 6-2). Hence, 200 ppm was chosen as the CO reference concentration; the C02 reference concentration was 500 ppm as usual.

Although the test 66 results for both CO (Figure 6-20) and C02 (Figure 6-21) are somewhat scattered, a consistent order showing more rapid reactivity loss following higher temperature calcination in dry N2 is established. However,

the addition of H20 in test 69 seemed to retard sorbent deterioration somewhat. The results from test 69 are clearly superior to the test 68 results, and are comparable to results using lower calcination temperature in dry N2. The maximum five-cycle deterioration for CO occurred in test 68 using

850°C calcination in dry N2 where (t-jVt/) = 0.68, and the maximum overall C02 deterioration also occurred in test 68 where (tsVtj*) = 0.75. These results are consistent with an average deterioration rate of 5 to 6% per cycle at 850°C and about 2% per cycle at 750°C. The overall average rate of deterioration at 850°C in the H20 - N2 calcination gas was about 3% per cycle. 215

From a process viewpoint, the higher sorbent deterioration rate at high calcination temperature would be balanced by a decrease in the required flow rate of calcination sweep gas and an increase in the C02 concentration in the product gas. At 850°C, the equilibrium partial pressure of C02 over CaO is approximately 0.5 atm, compared to about 0.08 atm at 750°C. Consequently, in steady-state calcination operating near equilibrium conditions, the required sweep gas flow rate would be 6 to 7 times less at

850°C than at 750°C, and the C02 concentration in the product gas would be 6 to 7 times higher.

6.6 Eleven-Cycle Test Results

Prebreakthrough CO and C02 concentrations and fractional removal of carbon oxides for 11-cycle test 65 are shown in

Figure 6-22. Also included are the comparable values from previous five-cycle test 27 at the same reaction conditions.

C02 prebreakthrough concentrations for the 16 cycles of the two tests ranged between about 170 and 280 ppm, while prebreakthrough CO concentrations were in the 10 to 30 ppm range. Fractional carbon oxide removal was either 0.996 or

0.997 in each of the 16 cycles. The fact that prebreakthrough concentrations did not increase through 11 cycles is particularly encouraging.

In spite of the relatively constant prebreakthrough CO and C02 concentrations, sorbent performance did deteriorate in Prebreakthrough Concentration, ppm (dry basis) 200 300 100 Figure 6-22.Figure Prebreakthrough CO and C02 Concentrationsand FractionalRemoval ofTotal Carbon Oxidesas a Function ofCycle65. Number. Test et 65 Test Test 27 Test Cycle Number 02 C CO Feed Gas Composition Gas Feed V- 45r (STP) 1425hr - SV T-550C P - 44 N 74.4% 00 HO H 20.0% FCOx .% CO 5.6% 15atm

1.00 0.99 0.98 0.97 0.96

Fractional Removal of COx

test 65. The most obvious sign of deterioration was a

decreased duration of the prebreakthrough period. This effect

is illustrated in Figure 6-23 where the dimensionless times

required for the CO concentration to reach 100 ppm and for the

C02 concentration to increase to 500 ppm are plotted versus

cycle number. For both components the overall average

decrease in dimensionless time was about 0.05 per cycle.

6.7 Ten Cycle Results Under Xsobaric Conditions

In most of the previous runs, calcination was carried out

at 3.3 atm and the carbonation/shift reaction was at 15 atm in

order to produce high purity H2. However, for possible

commercial processes, a dual recirculating fluidized-bed

reactor system will be most efficient. Such a process will be much easier to operate if both the calcination and

carbonation/shift reactors are at the same pressure.

Ten-cycle test 72 was carried out to evaluate system performance when both calcination and shift/carbonation were

carried out at 1 atm. 900°C calcination temperature was chosen because the C02 equilibrium pressure at this temperature is about 1 atm; hence, pure C02 can theoretically be produced during calcination. The feed gas contained all major components of coal gas with a steam to CO ratio of 3.78.

Feed concentrations of CO and C02 -- 13.7% and 5.1% -- were higher than the standard total carbon oxide feed concentration

-- 5.6% CO -- to promote rapid shift and carbonation at low Dimensionless Time, t 1.0 0.8 0.9 0.7 0.6 0.4 0.5 0.3 0.2 0.1 0.0

Figure 6-23. Dimensionless Time Required for the CO the for Required Time Dimensionless 6-23. Figure 1 I I I 1 1 L 1 I 1 I I I I I I I 23456789 Test 65 Test Concentration to Reach 100 ppm and for the for and 100ppm toReach Concentration Function of Cycle Number. Test 65. Test Number. Cycle of Function a as 500ppm toreach CO,Concentration CycleNumber O 100ppm - CO CO 0 11 10 2 500ppm -

218 219 pressure. The lower space velocity was chosen to increase the contact time between the gases and sorbent.

Component breakthrough curves for the first cycle of test

72 are shown in Figure 6-24. Prebreakthrough CO and C02 concentrations were sufficiently high to permit TCD analysis throughout the test. The concentrations are on a logarithmic scale to emphasize that there was no true prebreakthrough steady state established. The CO concentration, in particular, showed a continuous increase with time. This increase becomes effectively invisible when plotted on a linear scale. During the early portion of the reaction (for approximately the first 40 minutes) the fractional removal of total carbon oxides was about 0.97.

Figure 6-25 compares the fractional carbon oxide removal,

FCOX, and fractional calcium conversion, X’, as a function of dimensionless time, t\ for the first, sixth, and the tenth cycles of test 72. In cycle 1, FCOX began at about 0.97 and remained above 0.9 for t* s 0.7. The reaction became quite slow at t* > 1 .1 , and almost 10% carbon oxide removal was still occurring when the test was terminated at t* = 1.5. The final fractional sorbent conversion was X* = 0.87, which is comparable to the values obtained from tests using 15 atm shift-carbonation pressure and lower calcination temperature.

In the sixth cycle, FCOX exceeded 0.95 during the initial stages but dropped below 0.9 for t* > 0.5. At the end of the cycle, when t* = 1.45, FCOX was about 0.07, and X* was only Mole Fraction (dry basis), ppm iue -4 C, C02, CO, H 6-24. and Figure 0 40 4 40 8 50 2 50 560 540 520 500 480 460 440 420 400 CO First, Sixth, and Tenth Cycles of Test 72. Test of Cycles Tenth and Sixth, First, Time, minute Time, CO 2 Breakthrough Curves for the for Curves Breakthrough 2 Test 721 Test

220 Fractional COx Removal 1.1 1.0 0.8 0.9 0.7 0.6 0.5 0.4 0.2 0.3 0.0 0.1

. 02 . 06 . 10 . 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0

Figure 6-25. Fractional COx Removal and Fractional COxand Removal Fractional 6-25. Figure Sorbent Calcium Conversion for the First, the for Conversion Calcium Sorbent Sixth, and Tenth Cycles of Test 72. Test of Cycles Tenth and Sixth, DimensionlessTime, t*

Fractional Sorbent Conversion, X H to to 222

0.78. Further deterioration in performance is obvious at the

end of the tenth cycle. FCOX exceeded 0.9 only for t’ s 0.3 and the final fractional sorbent conversion was only 0.54.

A measure of sorbent performance on a cycle-by-cycle basis is shown in Figure 6-26 where the dimensionless time, t*, at which the fractional carbon oxide removal decreased to

0.9 (FCOX = 0.9) and the fractional sorbent conversion, X*, at t* = 1.0 are plotted versus cycle number. While sorbent deterioration does occur, it is no more severe at the reaction conditions of test 72 than at the earlier reaction conditions where higher shift/carbonation pressure and lower calcination temperature were used. This is illustrated in Figure 6-27 where the ratio, tiVtx*, of the dimensionless time required for FCOX to decrease to 0.9 in cycle i to the same dimensionless time in cycle 1 is plotted versus cycle number.

Results from test 72 are compared to results from test 65 which utilized 15 atm shift-carbonation pressure and 750°C calcination temperature. The values of tx* for test 65 were larger because of the higher reaction pressure, but there is little, if any, difference in the rate of sorbent deterioration. In both tests, the average rate of sorbent deterioration over 10 cycles was just over 5% per cycle.

6.8 Conclusions

Multicycle tests evaluated sorbent durability at reaction conditions suggested by the single cycle tests. Dolomite was i

t* or X 0.8 0.9 0.7 0.6 0.5 0.4 0.3 0.2 0.0 Figure 6-26. Sorbent Performance Deterioration Performance Sorbent 6-26. Figure t* for FCOX = 0.9 = FCOX t*for Through the Ten Cycles of Test 72. Test of Cycles theTen Through Cycle Number Cycle X* at t* = 1.0 = t* at X* Test 72 Test

223 ti-/t1 * at FCOX = 0.9 0.9 0.7 0.8 0.6 0.5 0.4 0.3 0.2 0.0 Figure 6-27.Figure ComparisonSorbent of Deterioration Test 72 Test 65Test Ratesof 65 Tests and 72. CycleNumber

224 the primary sorbent while marble chips were also tested for

comparison. The effect of calcination temperature and gas

composition, carbonation/shift temperature, space velocity,

and gas composition were studied. It was concluded that dolomite has better durability than marble chips. A deterioration rate of about 5% per cycle was determined for dolomite under the standard conditions, while for marble

chips, the average deterioration per cycle was about 14%.

Calcination at high temperature and in dry N2 resulted in an increased multicycle reactivity loss, but the addition of H20 to the calcination sweep gas retarded sorbent deterioration.

For the shift/carbonation reaction conditions, increasing temperature or decreasing H20 to CO ratio reduced the sorbent reactivity. However, no negative effect on sorbent durability was observed by increasing space velocity. Although the fractional removal of C0X was reduced by lowering the shift/carbonation pressure, sorbent durability was comparable with the results achieved using standard conditions. CHAPTER 7

SUMMARY, CONCLUSIONS AND RECOMMENDATIONS

Different technologies for producing hydrogen have been developed in the past fifty years; among these, steam reforming of methane or light hydrocarbons is the primary commercial process. Because coal represents a major energy resource in the world, hydrogen from coal may become dominant in the future. A technology that allows for utilization of coal and which is environmentally acceptable is coal gasification. Coal gas has numerous potential applications such as electrical power generation through fuel cell and production of hydrogen, methanol and other chemicals.

Synthesis gas from a hydrocarbon reformer or a coal gasifier may be either completely converted to hydrogen for ammonia production or adjusted to a certain H2/CO ratio for methanol synthesis through the water-gas shift reaction. The conventional shift process consists of multiple catalytic reaction steps with C02 removal using low temperature scrubbing, and associated gas-to-gas heat exchangers between the shift and C02 removal steps. Methanation or pressure swing adsorption (PSA) is usually employed if high purity H2 is needed.

A novel concept of combining the shift reaction with high temperature C02 removal for direct hydrogen production has been investigated. Potential advantages of the new process

226 227

include reduced capital costs by combining the reaction steps

and C02 separation into a single processing vessel. In

addition, the number of heat exchangers and excess steam

requirements would be reduced, providing improved thermal

efficiency. Finally, the new process would eliminate the need

for the expensive and sulfur-sensitive shift catalysts.

Calcium oxide based sorbents were used to absorb C02

through the gas-solid reaction:

CaO(s) + C02{g) * CaC03(s) (7-1)

By reversing the reaction (calcination), CaO could be

regenerated for multicycle use. By combining the carbonation

reaction with the shift reaction

CO(g) + H20(g) « H2(g) + C02(g) (7-2)

carbon monoxide conversion was increased and high purity of H2 was produced.

A laboratory-scale fixed-bed reactor system was used to

study the feasibility of the simultaneous shift reaction and high temperature C02 removal process for hydrogen production.

Experiments were conducted to determine the effects of temperature, pressure, space velocity and inlet gas composition on system behavior during both calcination and shift/carbonation stages. Dolomite was the primary sorbent precursor and alternate sorbents were tested for comparison.

Sorbent structural properties were measured using mercury porosimeter for selected tests. 228

There were two phases in this experimental research —

single cycle tests to examine the effect of reaction parameters and multicycle tests to evaluate sorbent durability. Each complete single cycle test consisted of a

calcination cycle and a shift/carbonation cycle, and each multicycle test consisted of five or more complete single

cycles.

The following conclusions were reached as a result of the single cycle tests:

1. Calcination can be conducted over a temperature range of 750°C to 900°C under either N2 or a mixture of N2 and steam.

The calcination reaction became very slow at 650°C.

2. C02 concentration in the calcination product gas can be increased by increasing temperature, reducing pressure, reducing sweep gas flow rate, and adding steam to the calcination gas.

3. No significant adverse effect of high calcination temperature was observed in the subsequent first shift/carbonation cycle. Adding steam to the calcination gas enhanced the calcination rate and created larger pores in the calcined solid than calcination in a pure N2 atmosphere, but had little effect on kinetics in the first carbonation cycle.

4. Shift-carbonation results using feed gas containing only CO, H20, and N2 provided definite proof that the water-gas shift and carbonation reactions occurred simultaneously over a wide range of temperature, pressure, and space velocity. 229

5. Equilibrium for the combined shift-carbonation

reactions was closely approached at 15 atm, in the temperature

range of 500°C to 650°C and at space velocities as high as 3400

hr'1 (STP) . Total concentrations of carbon oxides in the

reactor product during the prebreakthrough period of

approximately 300 ppm (dry basis) were routinely achieved while concentrations of less than 50 ppm (dry basis) were

achieved at the most favorable conditions.

6 . Prebreakthrough carbon oxide removal decreased as the

shift/carbonation temperature increased. Fractional C0X

removal reached a maximum of 0.999 at 500°C, exceeded 0.99

from 400°C to 550°C, and was about 0.93 at 650°C.

7. Lowering the shift/carbonation pressure decreased the

reaction rate and increase prebreakthrough CO and C02

concentration.

8 . As space velocity increased, prebreakthrough CO and

C02 concentrations increased slightly. However, combined reaction equilibrium was closely approached at space velocities as high as 3400 hr'1.

9. Reducing the H20 to CO ratio in the simulated coal gas

feed did not seriously reduce the prebreakthrough CO* removal.

The fractional carbon oxide removal was roughly the same when

the ratio was reduced from 4:1 to 2:1.

10. Commercial dolomite (from National Lime CO. Findley,

OH) was more reactive than limestone and marble chips, both of which had larger calcium content. The excess pore volume 230 associated with calcination of MgCOa made the calcium in the dolomite more accessible to the C02.

11. Heterogeneous catalysis is believed to be

responsible, at least in part, for the shift reaction at the experimental conditions studied. However, the specific

component responsible for the catalysis could not be identified.

Based on the single cycle results, multicycle tests were conducted and the follows conclusions were reached:

1. All multicycle tests showed a gradual decrease in sorbent reactivity with increasing cycle number.

2. At 550°C and 15 atm, prebreakthrough concentrations of

CO and C02 didn't change significantly over 5 cycles.

However, the duration of the prebreakthrough period and the global reaction rate both decreased with cycle number.

3. Dolomite had better durability than marble chips.

Although the prebreakthrough CO and C02 concentrations were comparable using marble chips and dolomite, the performance of the marble chips deteriorated more rapidly. At standard shift/carbonation conditions, the deterioration rate for dolomite was about 5% per cycle for dolomite compared to 14% per cycle for marble chips.

4. Increasing calcination temperature using dry N2 sweep gas increased the rate of reactivity loss, but the addition of

H20 to the calcination gas retarded sorbent deterioration. 231

5. No negative effect on sorbent durability was observed

by increasing space velocity.

Two tests were extended to eleven and ten cycles. The

eleven cycle test confirmed the conclusions reached in an

earlier five cycle test at the same reaction conditions. The average rate of sorbent deterioration was about 5% per cycle over the eleven cycles. For the ten cycle test, both the

calcination and shift/carbonation were conducted at 1 atm and the calcination temperature was increased to 900°C. The

fractional removal of C0X was reduced in the early stages of the shift-carbonation cycle. However, the sorbent durability was comparable with the results achieved at higher pressure and lower calcination temperature.

Favorable results from this research show that hydrogen production via simultaneous shift reaction and high temperature C02 removal is feasible. The following additional work is recommended:

1. An economic analysis should be conducted to evaluate whether the proposed idea is commercially attractive compared to the conventional shift processes for hydrogen production.

2. A small fluidized-bed process with a shift/carbonation reactor and a CaO regenerator is recommanded to further study sorbent performance and process operability. 232

3. Dolomite (from Findley, Ohio) proved to be superior to

limestone as a sorbent precursor in the fixed-bed study.

However, dolomites having different properties are available

throughout the country. A screening study should be conducted

to determine if the favorable results are also true for other

dolomites.

4. Synthesis gas has been considered as the starting point for hydrogen production in this research. It would be

interesting to determine if the simultaneous shift and

carbonation reactions could be further combined with steam

reforming of methane:

CHt(g) + 2 HzO(g) * 4 H2(g) + COz(g) (7-3)

CHt(g) + HzO(g) * 3Hz(g) + CO(g) (7-4)

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The author was born in Suqian, the People's Republic of

China on March 8, 1965, a day all women took as a holiday, but his mom worked real hard to have this boy - her seventh and last child.

After finishing his high school in Suqian, he was gracefully admitted to the University of Science and

Technology of China in Hefei, China, and finished his BS in engineering in July 1988. Then he got his industrial experience from Anqing Petrochemical Company. In August 1990, he came to the United States and pursued graduate degrees at

Louisiana State University, Baton Rouge, Louisiana. He obtained a MS in chemical engineering in 1992 and is presently a candidate for the Doctor of Philosophy.

He was happily married to Liping Yu and they have a beautiful son, Kevin. They are very confident and firmly believe that hard work will bring about a happy life.

Now, having spent thirty years in education, he wishes to feed back what he has to the society and get the best out of his life.

237 DOCTORAL EXAMINATION AND DISSERTATION REPORT

Candidate: Chun Han

Major Field: Chemical Engineering

Title of Dissertation: Simultaneous Water-Gas Shift Reaction and Carbon Dioxide Separation for Direct Hydrogen Production from Synthesis Gas

Approved: & ° l L Major Professor and chairman

ean of 'the Graduate School

EXAMINING COMMITTEE:

•cr

Date of Bx»nH nation:

December 7, 1994